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Renewable hydrocarbons for jet fuels from biomass and plastics via microwave-induced pyrolysis and hydrogenation processes

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RENEWABLE HYDROCARBONS FOR JET FUELS FROM
BIOMASS AND PLASTICS VIA MICROWAVE-INDUCED
PYROLYSIS AND HYDROGENATION PROCESSES
By
XUESONG ZHANG
A dissertation submitted in partial fulfilment of
the requirements for the degree of
DOCTOR OF PHILOSOPHY
WASHINGTON STATE UNIVERSITY
Department of Biological Systems Engineering
MAY 2016
ProQuest Number: 10139694
All rights reserved
INFORMATION TO ALL USERS
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ProQuest 10139694
Published by ProQuest LLC (2016). Copyright of the Dissertation is held by the Author.
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To the Faculty of Washington State University:
The members of the Committee appointed to examine the dissertation of XUESONG
ZHANG find it satisfactory and recommend that it be accepted.
___________________________________
Hanwu Lei, Ph.D., Chair
_________________________________
Shulin Chen, Ph.D.
___________________________________
Joan Q. Wu, Ph.D.
ii
ACKNOWLEDGEMENTS
Though only my name appears on the cover of this dissertation, many great people have
contributed to its production. I owe my sincere gratitude to all those people who have made
this dissertation possible and because of whom my graduate experience has been one that I
will cherish forever.
On the occasion of the completion of the thesis, I would like to convey my appreciation
firstly to my supervisor for Prof. Hanwu Lei. Without his patient supervision, every detail of
this thesis such as experiments and writing cannot be accomplished successfully. I also want
to thank my supervisor Prof. Hanwu Lei for his care and help not only in my study but also
on my work and life in the period of pursuing my Ph.D degree. I cannot overcome the
difficulty encountered without his encouragements and supports. And I have also learnt a lot
from him, such as active minds, serious work attitude and generous heart.
I would like to thank Professor Shulin Chen and Professor Joan Q. Wu for serving as my
committee members, and for their valuable comments and suggestions. Their work has
demonstrated that concern for development of modern technology in biofuel production,
should always transcend academia and provide a quest for our time.
I would also like to thank my colleagues in Prof. Lei’s bioenergy lab, Dr. Lu Wang, Dr. Yi
Wei, Lei Zhu, Moriko Qian, and Xiaolu Zhu for their friendships and collaborations. There
are also many other scientists helping me greatly with my research e.g. Dr. Aftab Ahamed and
Dr. Xin Li in the Bioproducts, Sciences, and Engineering Laboratory (BSEL), Jonathan
Lomber at the Department of Biological Systems Engineering in WSU Pullman, Dr. Valerie
Lynch-Holm in Franceschi Microscopy & Imaging Center (FMIC). I am thankful for all their
advice, suggestions, discussions and assistances during my doctoral research. The staffs at
WSU are always friendly and extremely helpful. I would like to thank Joanna Dreger for her
administrative support.
iii
I would like express thanks to Chinese Scholarship Council for their financial support granted
through predoctoral fellowship.
Most importantly, none of this would have been possible without the love and patience of my
family. My immediate family, to whom this dissertation is dedicated, has been a constant
source of love, concern, support and strength all these years. I would like to express my
heart-felt gratitude to my parents Xianquan Zhang and Jiamei Li, My brother Xuedong Zhang;
my parents-in-law Guangzhen Gu and Tongyao Li, and especially my beloved wife Yexuan
Gu. My family has aided and encouraged me throughout this endeavor. Moreover, I have to
give a special mention for the support given by all my friends, both in China and in the USA;
I warmly appreciate the generosity and understanding of them.
iv
RENEWABLE HYDROCARBONS FOR JET FUELS FROM
BIOMASS AND PLASTICS VIA MICROWAVE-INDUCED
PYROLYSIS AND HYDROGENATION PROCESSES
Abstract
by Xuesong Zhang, Ph.D.
Washington State University
May 2016
Chair: Hanwu Lei
This dissertation aims to enhance the production of aromatic hydrocarbons in the catalytic
microwave-induced pyrolysis, and maximize the production of renewable cycloalkanes for jet
fuels in the hydrogenation process. In the process, ZSM-5 catalyst as the highly efficient
catalyst was employed for catalyzing the pyrolytic volatiles from thermal decomposition of
cellulose (a model compound of lignocellulosic biomass). A central composite experiment
design (CCD) was used to optimize the product yields as a function of independent factors
(e.g. catalytic temperature and catalyst to feed mass ratio). The low-density polyethylene (a
mode compound of waste plastics) was then carried out in the catalytic microwave-induced
pyrolysis in the presence of ZSM-5 catalyst. Thereafter, the catalytic microwave-induced
co-pyrolysis of cellulose with low-density polyethylene (LDPE) was conducted over ZSM-5
catalyst. The results showed that the production of aromatic hydrocarbons was significantly
enhanced and the coke formation was also considerably reduced comparing with the catalytic
microwave pyrolysis of cellulose or LDPE alone. Moreover, practical lignocellulosic biomass
(Douglas fir sawdust pellets) was converted into aromatics-enriched bio-oil by catalytic
microwave pyrolysis. The bio-oil was subsequently hydrogenated by using the Raney Ni
catalyst. A liquid-liquid extraction step was implemented to recover the liquid organics and
remove the water content. Over 20% carbon yield of liquid product regarding lignocellulosic
v
biomass was obtained. Up to 90% selectivity in the liquid product belongs to jet fuel range
cycloalkanes. As the integrated processes was developed, catalytic microwave pyrolysis of
cellulose with LDPE was conducted to improve aromatic production. After the liquid-liquid
extraction by the optimal solvent (n-heptane), over 40% carbon yield of hydrogenated
organics based on cellulose and LDPE were achieved in the hydrogenation process. As such,
real lignocellulosic biomass with LDPE were transformed into aromatics via co-feed catalytic
microwave pyrolysis. It was also found that close to 40% carbon yield of hydrogenated
organics were garnered. Based on these outcomes, the reaction kinetics regarding
non-catalytic co-pyrolysis and catalytic co-pyrolysis of biomass with plastics were also
presented. In addition, the techno-economic analysis of the catalytically integrated processes
from lignocellulosic biomass to renewable cycloalkanes for jet fuels was evaluated in the
dissertation as well.
vi
TABLE OF CONTENTS
Page
ACKNOWLEDGEMENTS ................................................................................................. iii
ABSTRACT ......................................................................................................................... v
LIST OF TABLES ............................................................................................................ xvii
LIST OF FIGURES............................................................................................................. xx
CHAPTER ONE ................................................................................................................... 1
INTRODUCTION ................................................................................................................ 1
1.1 MOTIVATION ................................................................................................................. 1
1.2 COMPONENTS OF LIGNOCELLULOSIC BIOMASS................................................................ 3
1.3 AVAILABILITY OF WASTE POLYMERS ............................................................................... 5
1.3.1 Availability of waste plastics .................................................................................. 5
1.3.2 Availability of waste tires ....................................................................................... 5
1.4 BIOMASS CONVERSION ROUTES ..................................................................................... 6
1.5 BIO-OIL UPGRADING ...................................................................................................... 8
1.6 CATALYTIC CO-PYROLYSIS OF BIOMASS WITH POLYMERS ............................................... 10
1.6.1 The importance of catalytic co-pyrolysis .............................................................. 10
1.6.2. Chemistry of catalytic co-pyrolysis ..................................................................... 12
1.6.3 Mechanism of catalytic co-pyrolysis .................................................................... 13
1.7 HYDROGENATION PROCESS OF AROMATICS ................................................................... 16
1.8 OBJECTIVES ................................................................................................................ 17
1.9 OUTLINES ................................................................................................................... 18
vii
1.10 REFERENCES ............................................................................................................. 27
CHAPTER TWO ................................................................................................................ 34
RENEWABLE GASOLINE-RANGE AROMATICS AND HYDROGEN-ENRICHED FUEL
GAS FROM CELLULOSE VIA CATALYTIC MICROWAVE-INDUCED PYROLYSIS ... 34
2.1 ABSTRACT .................................................................................................................. 34
2.2 INTRODUCTION ........................................................................................................... 35
2.3 EXPERIMENTAL ........................................................................................................... 38
2.3.1 Material ............................................................................................................... 38
2.3.2 Catalyst preparation ............................................................................................. 38
2.3.3 Catalytic microwave pyrolysis and analysis ......................................................... 39
2.3.4 Experimental methods and data processing .......................................................... 41
2.4 RESULTS AND DISCUSSION ........................................................................................... 43
2.4.1 Optimization of catalyst ....................................................................................... 43
2.4.2 Product yields ...................................................................................................... 44
2.4.3 Analysis of the bio-oils ........................................................................................ 46
2.4.4 Analysis of non-condensable gases ...................................................................... 52
2.5 CONCLUSIONS ............................................................................................................. 56
2.6 REFERENCES ............................................................................................................... 58
PRODUCTION OF GASOLINE-RANGE HYDROCARBONS FROM
MICROWAVE-INDUCED PYROLYSIS OF LOW-DENSITY POLYETHYLENE OVER
ZSM-5 ................................................................................................................................ 60
3.2 INTRODUCTION ........................................................................................................... 61
3.2 MATERIALS AND METHODS ......................................................................................... 66
3.2.1 Materials.............................................................................................................. 66
3.2.2 Catalytic microwave degradation of low density polyethylene (LDPE) ................ 67
viii
3.2.3 Experimental design and optimization ................................................................. 68
3.2.4 Analytic methods ................................................................................................. 69
3.4 RESULTS AND DISCUSSION ........................................................................................... 72
3.4.1 Response surface analysis .................................................................................... 72
3.4.2 The analysis of the pyrolysis-oils by GC/MS ....................................................... 76
3.4.3 The analysis of gaseous fraction by Micro-GC ..................................................... 84
3.4.4 The analysis of coke deposition on the catalyst .................................................... 87
3.4.5 Mechanism analysis for aromatic hydrocarbons formation ................................... 89
3.5 CONCLUSIONS ............................................................................................................. 91
3.6 REFERENCES ............................................................................................................... 93
CHAPTER FOUR ............................................................................................................... 96
PRODUCTION OF RENEWABLE JET FUEL RANGE ALKANES AND AROMATICS
VIA INTEGRATED CATALYTIC PROCESSES OF INTACT BIOMASS.......................... 96
4.1 ABSTRACT .................................................................................................................. 96
4.2 INTRODUCTION ........................................................................................................... 97
4.3 EXPERIMENTAL SECTION ........................................................................................... 101
4.3.1 Materials............................................................................................................ 101
4.3.2 Catalyst preparation ........................................................................................... 102
4.3.3 Catalytic microwave pyrolysis of ligonocellulosic biomass ................................ 102
4.3.4 Hydrotreatment of bio-oils from catalytic microwave pyrolysis ......................... 103
4.3.5 Analytical techniques ......................................................................................... 103
4.3.6 Experimental methods and data processing ........................................................ 104
4.4 RESULTS AND DISCUSSION ......................................................................................... 106
ix
4.4.1 Catalytic transformation of lignocellulosic biomass into aromatics .................... 106
4.4.2 Hydrotreatment of bio-oils derived from catalytic microwave pyrolysis ............. 108
4.4.3 Analysis of gaseous fraction............................................................................... 122
4.5 CONCLUSIONS ........................................................................................................... 125
4.6 REFERENCES ............................................................................................................. 126
CHAPTER FIVE .............................................................................................................. 129
FROM LIGNOCELLULOSIC BIOMASS TO RENEWABLE CYCLOALKANES FOR JET
FUELS .............................................................................................................................. 129
5.1 ABSTRACT ................................................................................................................ 129
5.2 INTRODUCTION ......................................................................................................... 130
5.3 EXPERIMENTAL ......................................................................................................... 135
5.3.1 Materials............................................................................................................ 135
5.3.2 Catalyst preparation ........................................................................................... 136
5.3.3 Catalytic microwave-induced pyrolysis of lignocellulosic biomass .................... 137
5.3.4 Hydrogenation of model compounds in diverse solvents .................................... 138
5.3.5 Hydrotreatment of bio-oil derived from catalytic microwave pyrolysis .............. 139
5.3.6 Analytical techniques ......................................................................................... 140
5.3.7 Data evaluation .................................................................................................. 141
5.4 RESULTS AND DISCUSSION ......................................................................................... 143
5.4.1 Catalytic transformation of lignocellulosic biomass into C8- C16 aromatics ........ 143
5.4.2 Solvents effect on the hydrogenation of naphthalene .......................................... 145
5.4.3 In suit hydrogenation of extracted bio-oils ......................................................... 147
5.4.4 Catalyst characterization .................................................................................... 159
x
5.4.5 Reaction pathway for the conversion of lignocellulosic biomass into jet fuel range
cycloalkanes ............................................................................................................... 162
5.6 REFERENCES ............................................................................................................. 166
CHAPTER SIX ................................................................................................................. 170
DEVELOPMENT OF A CATALYTICALLY GREEN ROUTE FROM DIVERSE
LIGNOCELLULOSIC BIOMASS TO HIGH-DENSITY CYCLOALKANES FOR JET
FUELS .............................................................................................................................. 170
6.1 ABSTRACT ................................................................................................................ 170
6.2 INTRODUCTION ......................................................................................................... 171
6.3 EXPERIMENTAL ......................................................................................................... 174
6.3.1 Materials............................................................................................................ 174
6.3.2 Catalyst preparation ........................................................................................... 174
6.3.3 Catalytic microwave-induced pyrolysis of diverse ligonocellulosic biomasses ... 175
6.3.4 Hydrotreatment of liquid organics derived from catalytic microwave pyrolysis .. 176
6.3.5 Analytical techniques ......................................................................................... 177
6.3.6 Experimental methods and data evaluation ........................................................ 178
6.4 RESULTS AND DISCUSSION ......................................................................................... 180
6.4.1 Catalyst characterization of as-prepared Raney nickel ........................................ 180
6.4.2 Optimization of liquid organics from catalytic microwave pyrolysis .................. 182
6.4.3 Hydrotreatment of extracted organics................................................................. 190
6.5 CONCLUSIONS ........................................................................................................... 201
6.6 REFERENCES ............................................................................................................. 203
CHAPTER SEVEN........................................................................................................... 206
SYNTHESIS OF HIGH-DENSITY JET FUEL FROM PLASTICS VIA CATALYTICALLY
INTEGRAL PROCESSES ................................................................................................ 206
7.1 ABSTRACT ................................................................................................................ 206
xi
7.2 INTRODUCTION ......................................................................................................... 207
7.3 EXPERIMENTAL ......................................................................................................... 211
7.3.1 Materials............................................................................................................ 211
7.3.2 Catalyst preparation ........................................................................................... 212
7.3.3 Catalytic microwave-assisted degradation of low-density polyethylene (LDPE) 214
7.3.4 Hydrotreatment of liquid organics derived from catalytic microwave degradation
................................................................................................................................... 214
7.3.5 Analytical techniques ......................................................................................... 215
7.3.6 Experimental methods and data evaluation ........................................................ 217
7.4 RESULTS AND DISCUSSION ......................................................................................... 218
7.4.1 Catalytic transformation of LDPE into liquid organics ....................................... 218
7.4.2 Hydrotreatment of liquid organics derived from catalytic microwave degradation
................................................................................................................................... 221
7.4.3 Reaction pathway for the conversion of plastics (LDPE) into JP-5 navy fuel...... 233
7.5 CONCLUSIONS ........................................................................................................ 235
7.6 REFERENCES ............................................................................................................. 237
CHAPTER EIGHT ........................................................................................................... 240
FROM PLASTICS TO DIVERSE GRADES OF JET FUELS VIA COMBINED
CATALYTIC CONVERSIONS ......................................................................................... 240
8.1 ABSTRACT ................................................................................................................ 240
8.2 INTRODUCTION ......................................................................................................... 241
8.3 MATERIALS AND METHODS ....................................................................................... 244
8.3.1 Materials............................................................................................................ 244
8.3.2 Catalyst preparation ........................................................................................... 245
8.3.3 Catalytic microwave-assisted degradation of low-density polyethylene (LDPE) 246
xii
8.3.4 Hydrogenation of raw organics derived from catalytic microwave degradation .. 247
8.3.5 Analytical techniques ......................................................................................... 247
8.3.6 Experimental methods and data evaluation ........................................................ 250
8.4 RESULTS AND DISCUSSION ......................................................................................... 251
8.4.1 Catalyst characterization .................................................................................... 251
8.4.2 Catalytic transformation of LDPE into liquid organics ....................................... 255
8.4.3 Hydrogenation of raw organics derived from catalytic microwave degradation .. 257
8.5 CONCLUSIONS ........................................................................................................... 276
8.6 REFERENCES ............................................................................................................. 279
CHAPTER NINE .............................................................................................................. 282
OPTIMIZING CARBON EFFICIENCY OF JET FUEL RANGE ALKANES FROM
CELLULOSE CO-FED WITH POLYETHYLENE VIA CATALYTICALLY TANDEM
PROCESSES .................................................................................................................... 282
9.1 ABSTRACT ................................................................................................................ 282
9.2 INTRODUCTION ......................................................................................................... 283
9.3 EXPERIMENTAL SECTIONS .......................................................................................... 287
9.3.1 Materials............................................................................................................ 287
9.3.2 Catalyst preparation ........................................................................................... 287
9.3.3 Co-feed catalytic microwave-induced pyrolysis of cellulose and LDPE ............. 288
9.3.4 Hydrogenation of liquid organics derived from co-feed catalytic microwave
pyrolysis ..................................................................................................................... 289
9.3.5 Analytical techniques......................................................................................... 289
9.3.6 Experimental methods and data evaluation ........................................................ 292
9.4 RESULTS AND DISCUSSION ......................................................................................... 294
9.4.1 Catalyst characterization .................................................................................... 294
xiii
9.4.2 Product yield distributions from co-feed catalytic microwave pyrolysis ............. 297
9.4.3 The effect of catalytic temperature on the co-feed catalytic microwave pyrolysis 299
9.4.4 The effect of LDPE to cellulose ratio on the co-feed catalytic microwave pyrolysis
................................................................................................................................... 305
9.4.5 Hydrogenation of liquid organics derived from co-feed catalytic microwave
pyrolysis for jet fuels .................................................................................................. 309
9.4.6 Reaction pathway for the conversion regrading co-feeding of cellulose with LDPE
into jet fuels................................................................................................................ 314
9.5 CONCLUSIONS ........................................................................................................... 317
9.6 REFERENCES ............................................................................................................. 318
CHAPTER TEN................................................................................................................ 322
ENHANCEMENT OF JET FUEL RANGE ALKANES FROM CO-FEEDING OF
LIGNOCELLULOSIC BIOMASS WITH PLASTICS VIA TANDEM CATALYTIC
CONVERSIONS............................................................................................................... 322
10.1 ABSTRACT .............................................................................................................. 322
10.2 INTRODUCTION ....................................................................................................... 323
10.3 EXPERIMENTAL SECTIONS ........................................................................................ 328
10.3.1 Materials.......................................................................................................... 328
10.3.2 Catalyst preparation ......................................................................................... 328
10.3.3 Co-feed catalytic microwave-induced pyrolysis of lignocellulosic biomass and
plastics ....................................................................................................................... 329
10.3.4 Hydrogenation of liquid organics derived from co-feed catalytic microwave
pyrolysis..................................................................................................................... 329
10.3.5 Analytical techniques ....................................................................................... 330
10.3.6 Experimental methods and data evaluation....................................................... 332
10.4 RESULTS AND DISCUSSION ....................................................................................... 334
10.4.1 Catalyst characterization .................................................................................. 334
10.4.2 Product yield distributions from co-feed catalytic microwave pyrolysis ........... 337
xiv
10.4.3 The effect of catalytic temperature on the co-feed catalytic microwave pyrolysis
................................................................................................................................... 341
10.4.4 The effect of plastics to biomass ratio on the co-feed catalytic microwave
pyrolysis ..................................................................................................................... 346
10.4.5 Process robustness: the recyclability of the well-promoted ZSM-5 catalyst ...... 350
10.4.6 Hydrogenation of liquid organics derived from co-feed catalytic microwave
pyrolysis for jet fuels .................................................................................................. 354
10.4.7 Reaction pathway for the conversion regrading co-feeding of biomass with
plastics into jet fuels ................................................................................................... 358
10.5 CONCLUSIONS ......................................................................................................... 361
10.6 REFERENCES ........................................................................................................... 363
CHAPTER ELEVEN ........................................................................................................ 367
TECHNO-ECONOMIC EVALUATION OF RENEWABLE CYCLOALKANES FOR JET
FUELS FROM LIGNOCELLULOSIC BIOMASS ........................................................... 367
11.1 ABSTRACT .............................................................................................................. 367
11.2 INTRODUCTION ....................................................................................................... 368
11.3 PROCESS MODEL DESCRIPTION ................................................................................. 373
11.4 RESULTS AND DISCUSSION ....................................................................................... 383
11.4.1 Overall mass balances ...................................................................................... 383
11.4.2 Capital cost ...................................................................................................... 384
11.4.3 Operating cost .................................................................................................. 387
11.4.4 Economic assessment of the biorefinery process .............................................. 388
11.4.5 Sensitivity Analysis.......................................................................................... 389
11.4.6 Limitations of the Analysis............................................................................... 390
11.5 CONCLUSIONS ......................................................................................................... 391
11. 6 REFERENCES .......................................................................................................... 393
xv
CHAPTER TWELVE........................................................................................................ 395
THERMAL BEHAVIOR AND KINETIC STUDY FOR CATALYTIC CO-PYROLYSIS OF
BIOMASS WITH PLASTICS ........................................................................................... 395
12.1 ABSTRACT .............................................................................................................. 395
12.2 INTRODUCTION ....................................................................................................... 396
12.3 EXPERIMENTAL ....................................................................................................... 399
12.3.1 Materials.......................................................................................................... 399
12.3.2 Catalyst preparation ......................................................................................... 400
12.3.3 Thermogravimetric analysis ............................................................................. 400
12.3.4 Kinetic study.................................................................................................... 401
12.4 RESULTS AND DISCUSSION ....................................................................................... 402
12.4.1 Pyrolysis of materials and their blends ............................................................. 402
12.4.2 Kinetic parameters ........................................................................................... 409
12. 5 CONCLUSIONS ........................................................................................................ 415
12. 6 REFERENCES .......................................................................................................... 417
CHAPTER THIRTEEN .................................................................................................... 419
GENERAL CONCLUSIONS AND FUTURE WORK SUGGESTIONS ........................... 419
13.1 GENERAL CONCLUSIONS .......................................................................................... 419
13.2 FUTURE RESEARCH SUGGESTIONS ............................................................................ 425
LIST OF PAPERS IN THE PERIOD OF APPLICATION OF PH.D DEGREE ......... ERROR!
BOOKMARK NOT DEFINED.
xvi
LIST OF TABLES
Table 1.1 Basic elementary composition of fast pyrolysis oil and crude oil ------------------------ 8
Table 2.1 Experimental design and product yield distribution ---------------------------------------- 42
Table 2.2 Textural and acid properties of the ZSM-5 zeolites ---------------------------------------- 43
Table 3.1 Experimental design and product yield distribution ---------------------------------------- 71
Table 3.2 Identified hydrocarbon products for gasoline versus LDPE pyrolysis-oils (%) -------- 78
Table 3.3 Products selectivity for aromatic hydrocarbons species on the base of catalytic
temperature by GC/MS area at the same reactant to catalyst ratio of 3 ----------------------------- 79
Table 3.4 Products selectivity for aromatic hydrocarbons species on the base of reactant to
catalyt ratio by GC/MS area at the same catalytic temperature (375°C) ---------------------------- 83
Table 4.1 Experimental design and product yield distribution ---------------------------------------- 105
Table 4.2 Major paraffins and olefins in hydrotreated oils as a function of reaction
temperature --------------------------------------------------------------------------------------------------- 111
Table 4.3 Major chemical compounds in hydrotreated oils as a function of reaction
temperature --------------------------------------------------------------------------------------------------- 113
Table 4.4 Major chemical compounds in hydrotreated oils as a function of reaction
temperature --------------------------------------------------------------------------------------------------- 117
Table 5.1 Effect of various solvents on the hydrogenation reaction of naphthalene --------------- 144
Table 5.2 Products distribution and partial cycloalkanes carbon selectivity as a function of
reaction temperature ---------------------------------------------------------------------------------------- 151
Table 5.3 Products distribution and partial cycloalkanes carbon selectivity on the basis of
reaction time ------------------------------------------------------------------------------------------------- 155
Table 5.4 Products distribution and partial cycloalkanes carbon selectivity in the presence of
various catalysts --------------------------------------------------------------------------------------------- 157
Table 5.5 Textural properties of Ni-Al alloy, NP Ni, Raney-Ni 4200 and home-made Raney
Ni catalysts --------------------------------------------------------------------------------------------------- 159
Table 6.1 Textural properties of Ni-Al alloy powder and home-made Raney Ni catalyst -------- 180
xvii
Table 6.2 Products distribution and main aromatics carbon selectivity from diverse
lignocellulosic biomasses at 500 °C ---------------------------------------------------------------------- 188
Table 6.3 Products distribution and partial cycloalkanes carbon selectivity from Douglas fir
as a function of reaction temperature --------------------------------------------------------------------- 194
Table 6.4 Products distribution and partial cycloalkanes carbon selectivity as a function of
diverse biomass sources ------------------------------------------------------------------------------------ 198
Table 7.1 Textural properties of as-received Raney-Ni 4200 catalysts ------------------------------ 213
Table 7.2 Experimental design and product yield distribution of hydrotreated organics --------- 217
Table 7.3 Products distribution and partial alkanes’ carbon selectivity as a function of reaction
temperature --------------------------------------------------------------------------------------------------- 227
Table 7.4 Products distribution and partial alkanes’ carbon selectivity as a function of reaction
temperature --------------------------------------------------------------------------------------------------- 231
Table 8.1 Textural properties of well-promoted ZSM-5 and Raney Ni catalysts ------------------ 250
Table 8.2 Products distribution and partial alkanes’ carbon selectivity as a function of reaction
temperature --------------------------------------------------------------------------------------------------- 262
Table 8.3 Products distribution and partial cycloalkanes carbon selectivity on the basis of
reaction time ------------------------------------------------------------------------------------------------- 268
Table 8.4 Products distribution and partial cycloalkanes carbon selectivity with respect to
catalyst to reactant ratio ------------------------------------------------------------------------------------ 270
Table 8.5 Products distribution and partial alkanes’ carbon selectivity in the presence of
various catalysts --------------------------------------------------------------------------------------------- 273
Table 9.1 Experimental design and product yield distribution ---------------------------------------- 293
Table 9.2 Textural properties of well-promoted ZSM-5 and home-made Raney Ni catalysts --- 294
Table 9.3 Detailed carbon yield distribution and product carbon selectivity as a function of
catalytic temperature ---------------------------------------------------------------------------------------- 303
Table 9.4 Detailed carbon yield distribution and product carbon selectivity on the basis of
LDPE to cellulose ratio ------------------------------------------------------------------------------------- 308
Table 9.5 Products distribution and partial alkanes’ carbon selectivity for hydrogenation of
diverse liquid organics -------------------------------------------------------------------------------------- 312
Table 10.1 Experimental design and product yield distribution -------------------------------------- 333
xviii
Table 10.2 Textural properties of well-promoted ZSM-5 and home-made Raney Ni catalysts -- 334
Table 10.3 Detailed carbon yield distribution and product carbon selectivity as a function of
catalytic temperature ---------------------------------------------------------------------------------------- 344
Table 10.4 Detailed carbon yield distribution and product carbon selectivity on the basis of
plastics to biomass ratio ------------------------------------------------------------------------------------ 349
Table 10.5 Detailed carbon yield distribution and product carbon selectivity with respect to
catalyst reused times ---------------------------------------------------------------------------------------- 353
Table 10.6 Products distribution and partial alkanes’ carbon selectivity for hydrogenation of
diverse liquid organics -------------------------------------------------------------------------------------- 356
Table 11.1 Proximate and elemental analyses of diverse lignocellulosic biomasses -------------- 372
Table 11.2 Experimental design and product yield distribution of hybrid poplar ------------------ 374
Table 11.3 Experimental design and product yield distribution of loblolly pine ------------------- 375
Table 11.4 Experimental design and product yield distribution of Douglas fir -------------------- 376
Table 11.5 Products distribution and main aromatics carbon selectivity from diverse
lignocellulosic biomasses at 500 ºC ---------------------------------------------------------------------- 379
Table 11.6 Products distribution and partial cycloalkanes carbon selectivity as a function of
diverse biomass sources ------------------------------------------------------------------------------------ 381
Table 11.7 Summary of estimated capital costs for required equipment in the integrated
processes ------------------------------------------------------------------------------------------------------ 385
Table 11.8 Summary of operating costs involved in the processes from hybrid poplar to
cycloalkanes -------------------------------------------------------------------------------------------------- 386
Table 11.9 Annual income from the sales of final products ------------------------------------------- 388
Table 12.1 Representative results regarding kinetic parameters for the co-pyrolysis of
cellulose with LDPE ---------------------------------------------------------------------------------------- 409
Table 12.2 Representative results regarding kinetic parameters for the co-pyrolysis of Doulas
fir sawdust with LDPE ------------------------------------------------------------------------------------- 413
xix
LIST OF FIGURES
Fig. 1.1 Schematic representation of lignocellulosic biomass with the structure of cellulose,
hemicellulose and lignin------------------------------------------------------------------------------------ 4
Fig. 1.2 Schematic presentation for the productions of aromatics from catalytic co-pyrolysis
of cellulose with LDPE in the presence of ZSM-5 catalyst ------------------------------------------- 14
Fig. 1.3 Possible reaction scheme regarding the formations of naphthalene and indene from
catalytic co-pyrolysis of cellulose with PS by using HZSM-5 as the catalyst ---------------------- 15
Fig. 2.1 Chemical composition of bio-oils on the base of catalytic temperature at the same
WHSV-1 of 0.067 h------------------------------------------------------------------------------------------ 46
Fig. 2.2 Chemical composition of bio-oils as a function of catalytic temperature at the same
temperature of 375 °C -------------------------------------------------------------------------------------- 50
Fig. 2.3 The composition of gaseous fraction with respect of catalytic temperature at the same
WHSV-1 (0.067 h) ------------------------------------------------------------------------------------------- 54
Fig. 2.4 The proposed reaction pathway for furfural conversion to xylenes and toluene over
modified ZSM-5 --------------------------------------------------------------------------------------------- 55
Fig. 2.5 The composition of gaseous fraction as a function of WHSV-1 at the same catalytic
temperature (375 °C)---------------------------------------------------------------------------------------- 56
Fig. 3.1 The schematic diagram of the microwave-assisted pyrolysis system integrated with
zeolite catalysis process ------------------------------------------------------------------------------------ 67
Fig. 3.2 Effect of the interaction of the independent variables on pyrolysis-oil (A) and gas (B) 75
Fig. 3.3 The chemical composition of the pyro-oils from GC/MS analysis with ZSM ----------- 75
Fig. 3.4 The composition of gaseous fraction with respect of catalytic temperature at the same
reactant to catalyst ratio of 3 ------------------------------------------------------------------------------- 84
Fig. 3.5 The development of the gaseous composition as a function of reactant to catalyst
ratio at the same catalytic temperature (375°C)--------------------------------------------------------- 86
Fig. 3.6 Effect of the interaction of the independent variables on coke yield ---------------------- 88
Fig. 3.7 The overall reaction pathway for converting LDPE into aromatic hydrocarbons ------- 90
xx
Fig. 4.1 Product distribution of organics based on carbon number and various functional
groups --------------------------------------------------------------------------------------------------------- 106
Fig. 4.2 Effect of the interaction of the independent variables on selectivity of paraffins -------- 109
Fig. 4.3 Proposed reaction pathway for converting phenolic monomers (phenol, guaiacol,
catechol) into cyclohexanol and 1, 2-cyclohexanediol------------------------------------------------- 115
Fig. 4.4 Chemical composition of hydrotreated oils on the basis of initial pressure -------------- 118
Fig. 4.5 Chemical composition of hydrotreated oils as a function of reaction time --------------- 119
Fig. 4.6 Chemical composition of hydrotreated oils on the basis of catalyst loading ------------- 121
Fig. 4.7 Effects of Reaction temperature (A), Reaction time (B), catalyst loading (C) on the
composition of gaseous fraction--------------------------------------------------------------------------- 124
Fig. 5.1 Overall carbon yield distribution (A) and main aromatics carbon selectivity (B) for
extracted bio-oils from catalytic microwave pyrolysis of Douglas fir sawdust pellets ----------- 142
Fig. 5.2 Overall products distribution (A) and partial aromatics carbon selectivity (B) with
resptect with catalyst loading ------------------------------------------------------------------------------ 150
Fig. 5.3 The XRD patterns of the Ni-Al alloy powder, NP-Ni, Raney-Ni 4200, and
home-made Raney Ni --------------------------------------------------------------------------------------- 160
Fig. 5.4 SEM images of the Ni-Al alloy powder (A), NP-Ni (B), Raney-Ni 4200 (C), and
home-made Raney Ni (D) ---------------------------------------------------------------------------------- 161
Fig. 5.5 Proposed reaction pathways for the conversion of lignocellulosic biomass into jet
fuel range cycloalkanes ------------------------------------------------------------------------------------- 164
Fig. 6.1 The XRD patterns of the Ni-Al alloy powder and home-made Raney Ni catalyst ------ 181
Fig. 6.2 SEM images of the Ni-Al alloy powder (A) and Raney Ni catalyst (B) ------------------ 182
Fig. 6.3 Effect of the interaction of the independent variables on organic yield from hybrid
poplar (A), loblolly pine (B) and Douglas fir (C) yields ---------------------------------------------- 185
Fig. 6.4 Overall products distribution (A) and partial cycloalkanes carbon selectivity (B) with
respect with catalyst loading ------------------------------------------------------------------------------- 195
Fig. 7.1 SEM image of the as-received Raney Ni catalyst -------------------------------------------- 213
Fig. 7.2 Overall carbon yield distribution (C mol%), (A) and carbon selectivity of main
chemical compounds (C mol%), (B) in the liquid organics from catalytic microwave
degradation of LDPE --------------------------------------------------------------------------------------- 220
xxi
Fig. 7.3 Chemical composition of hydrotreated organics on the basis of initial pressure--------- 223
Fig. 7.4 Chemical composition of hydrotreated organics with respect to catalyst to reactant
ratio ------------------------------------------------------------------------------------------------------------ 225
Fig. 7.5 Reaction pathways for the conversion of plastics (LDPE) into jet fuels ------------------ 234
Fig. 8.1 Reaction pathways for the conversion of plastics (LDPE) into jet fuels ------------------ 244
Fig. 8.2 NH3-TPD profiles of parent ZSM-5 and well-promoted ZSM-5 catalyst ----------------- 251
Fig. 8.3 The XRD patterns of the Ni-Al alloy powder, NP-Ni, Raney-Ni 4200, and
home-made Raney Ni catalyst ----------------------------------------------------------------------------- 253
Fig. 8.4 SEM images of the Ni-Al alloy powder (A), NP-Ni (B), Raney-Ni 4200 (C), and
home-made Raney Ni (D) ---------------------------------------------------------------------------------- 254
Fig. 8.5 Overall carbon yield distribution (C mol%) and carbon selectivity of main chemical
compounds (C mol%)--------------------------------------------------------------------------------------- 255
Fig. 8.6 Chemical composition of hydrotreated organics on the basis of reaction temperature - 259
Fig. 8.7 Chemical composition of hydrotreated organics with respect to initial pressure -------- 264
Fig. 8.8 Carbon selectivity of main alkanes with respect to initial pressure ------------------------ 266
Fig. 8.9 Chemical composition of hydrotreated organics with respect to recycle times ---------- 276
Fig. 9.1 NH3-TPD profiles of parent ZSM-5 and well-promoted ZSM-5 catalyst ----------------- 295
Fig. 9.2 The XRD patterns of the Ni-Al alloy powder and home-made Raney Ni catalyst ------ 296
Fig. 9.3 SEM images of Ni-Al alloy powder (A) and home-made Raney Ni catalyst (B)-------- 297
Fig. 9.4 The overall carbon yields regarding product distribution from co-feed catalytic
microwave pyrolysis in light of catalytic temperature ------------------------------------------------- 301
Fig. 9.5 The overall carbon yields regarding product distribution from co-feed catalytic
microwave pyrolysis as a function of LDPE to cellulose ratio --------------------------------------- 306
Fig. 9.6 Proposed reaction pathways for the conversion of cellulose and LDPE into jet fuel
range alkanes ------------------------------------------------------------------------------------------------- 314
Fig. 10.1 NH3-TPD profiles of parent ZSM-5 and well-promoted ZSM-5 catalyst --------------- 335
Fig. 10.2 The XRD patterns of the Ni-Al alloy powder and home-made Raney Ni catalyst ----- 336
Fig. 10.3 SEM images of Ni-Al alloy powder (A) and home-made Raney Ni catalyst (B) ------ 337
Fig. 10.4 The overall carbon yields regarding product distribution from co-feed catalytic
microwave pyrolysis in light of catalytic temperature ------------------------------------------------- 341
xxii
Fig. 10.5 The overall carbon yields regarding product distribution from co-feed catalytic
microwave pyrolysis as a function of plastics to biomass ratio--------------------------------------- 347
Fig. 10.6 The overall carbon yields regarding product distribution from co-feed catalytic
microwave pyrolysis on the basis of catalyst recycle times ------------------------------------------- 352
Fig. 10. 7 Proposed reaction pathways for the conversion of lignocellulosic biomass and
plastics into jet fuel range alkanes ------------------------------------------------------------------------ 360
Fig. 11.1 Simplified flow diagram of a biorefinery process for the production of cycloalkanes
for jet fuels --------------------------------------------------------------------------------------------------- 372
Fig. 11.2 The organic composition from hybrid poplar as a function of catalytic temperature -- 377
Fig. 11.3 The organic composition from loblolly pine as a function of catalytic temperature --- 378
Fig. 11.4 The organic composition from Douglas fir as a function of catalytic temperature ----- 379
Fig. 11.5 Overall input-output analysis for MAECP and HG systems in this biorefinery -------- 384
Fig. 11.6 Sensitivity analysis on return of investment (ROI) ----------------------------------------- 389
Fig. 12.1 TG curves of cellulose with LDPE (A); and Douglas fir sawdust (DF) with LDPE
(B) ------------------------------------------------------------------------------------------------------------- 404
Fig. 12.2 DTG curves of cellulose with LDPE (A); and Douglas fir sawdust (DF) with LDPE
(B) ------------------------------------------------------------------------------------------------------------- 407
Fig. 12.3 Variation of ∆W for biomass (cellulose and DF) and LDPE blends---------------------- 407
Fig. 12.4 Plots of ln(-ln(1-x)/T2) vs 1/T of LDPE (A), Cellulose (B), Cellulose-LDPE (C),
Cellulose-LDPE-Catalyst calculated by using TGA --------------------------------------------------- 412
Fig. 12.5 Plots of ln(-ln(1-x)/T2) vs 1/T of LDPE (A), DF (B), DF-LDPE (C),
DF-LDPE-Catalyst calculated by using TGA ----------------------------------------------------------- 415
xxiii
Dedication
To my beloved parents and wife who provided encouragement
and support during the study and before
xxiv
CHAPTER ONE
INTRODUCTION
1.1 Motivation
Extensive consumption of natural resources (such as coal, petroleum, and natural gas) is
predicted to grow continuously at the annual rate of 1.6% in the next two decades.1 The
petroleum fuel resources are recognized as valuable supply of finite natural energy, because
their current depleted rate is orders of magnitude higher than their corresponding regeneration
cycle. The massive utilization of petroleum fuel resources results in the augmented emissions
of detrimental pollutants (e.g., sulfur dioxide, oxides of nitrogen, and carbon dioxide),
triggering a sequence of environmental problems.2 Carbon dioxide, as a strong greenhouse
gas released into the atmosphere by means of petroleum resources combustion, contributes to
global warming.3
Despite a variety of environmental issues caused by the use of petroleum fuel resources,
petroleum-derived liquid hydrocarbons still served as the most attractive and feasible
transportation fuels, including gasoline and jet fuels.4 For instance, 7 billion barrels of
petroleum fuels were consumed for the United States in 2010; with 71% of petroleum fuels
(5.2 billion barrels) went toward satisfying the combined demand of gasoline, jet, and diesel
fuels.5 Within this context, the petroleum resources will be exhausted worldwide after 2042
without considering projected growth of petroleum fuel consumption.6
1
These issues are inherently attributed to the limitations of petroleum resources. To control the
pollutant emissions and mitigate the energy crisis, several efforts are currently underway to
seek new eco-friendly sources of energy for future generations.3 The renewable energy
sources can not only dispel the negative environmental impacts, but also eliminate the
dependence on the utilization of petroleum fuels.7 For the sake of
substituting renewable
biofuels for petroleum-derived fuels, the U.S. renewable fuels standard (RFS2) claimed that
the domestic supply of alternative fuels would increase to 36 billion gallons by 2022; besides
the U.S. Energy Information Administration stated that irrespective of varied prices of the oil
prices the production of renewable biofuels will soar in the next three decades.8
In fact, current demand for renewable biofuels originates from starch-based feedstock,
competing with edible resources.9 Ideally, the second generation biofuels produced from
lignocellulosic biomass (e.g., woods, grass, energy crops, and agricultural residues) are
preferred, which can be compatible with the existing energy infrastructure.8 Lignocellulosic
biomass has been identified as the most promising feedstock for the production of renewable
biofuels.10-15 That is because lignocellulosic biomass is the most abundant and renewable
source of non-edible biomass. A total of 220 billion dry tons of annual lignocellulosic
biomass could be sustainably produced worldwide.6 For the United States alone, the annual
production (1.3 billion dry tons) of lignocellulosic biomass could offset 50% carbon
consumption for gasoline and diesel.16,
17
Unlike the petroleum sources, lignocellulosic
biomass has been considered as carbon-neutral sources, conducing to mitigating global
2
warming effects.4 From the economic point of view, lignocellulosic biomass as the feedstock
to manufacture biofuels is much more inexpensive than edible biomass (such as corn starch)
and crude oil.18 Accordingly, the effective implementation of lignocellulosic biomass
necessitates appropriate technologies to produce large amounts of biofuels.3
1.2 Components of lignocellulosic biomass
A fundamental understanding regarding the characteristics of lignocellulosic biomass is
essential to rationally illustrate the mechanisms of catalytic co-pyrolysis. In this respect, we
will summarize the properties of lignocellulosic biomass in this section. Lignocellulosic
biomass is a complicated biopolymer, primarily consisting of cellulose, hemicellulose, and
lignin (Fig. 1).19, 20 Typically, lignocellulosic biomass contains 40 - 50 wt% of cellulose, 15 –
30 wt% of hemicellulose, and 15 – 30 wt% of lignin.9 Cellulose is a linear polysaccharide
polymer of glucose strongly linked by β-1, 4-glycoside linkage, acting as the framework of
lignocellulose cell walls (Fig. 1.1).21
A variety of hydroxyl groups are also presented in
cellulose chains, giving rise to the formation of hydrogen bonds.22 Cellulose is made up of
regions: crystalline (high crystallinity) and amorphous (low crystallinity) regions.23 Moreover,
cellulose microfibrils are linked to each other by hemicellulose and/or pectin and covered by
lignin.22
3
Fig. 1.1 Schematic representation of lignocellulosic biomass with the structure of cellulose,
hemicellulose and lignin.
In contrast to cellulose, hemicellulose is more structurally amorphous, random, and has a
heterogeneous composition. The structure of hemicellulose contains various pentoses (xylose
and arabinose), hexoses (glucose, galactose, mannose, rhamnose, and fucose), and uronic
acids (glucuronic acid, methyl glucuronic acid, and galacturonic acid).9 Short and branched
chains of hemicellulose is conducive to building a network with cellulose microfibrils and
lignin, rendering the lignocellulose matrix extremely rigid.22 Lignin is the second abundant
organic composition in nature, which is water insoluble and optically inert. It is a complex
aromatic and hydrophobic amorphous biopolymer of propyl-phenol groups.24 It is constructed
of three basic phenol-containing components: p-coumaryl, coniferyl, and sinapyl alcohols
(Fig. 1).25 These units are bound together by C–O (β-O-4, α-O-4, and 4-O-5 linkage) and
C–C (β-5, 5-5, β-1, β–β linkage).26 Lignin also plays a crucial role in the cross-linking
between cellulose and hemicellulose to garner a rigid three-dimensional structure of the cell
walls. In addition, lignin stores approximately 40% of energy of lignocellulosic biomass
4
because of its high carbon content.25 In general, softwood contains more lignin fraction than
hardwood and most agricultural residues.22
1.3 Availability of waste polymers
1.3.1 Availability of waste plastics
The availability of plastics as a co-reactant during co-pyrolysis is sufficient for the future
sustainability of fuels production.6 Plastics are abundantly used in normal life worldwide. For
instance, the total consumption of plastics in Western Europe was estimated at 33.4 million
tons in 1996; while it was augmented to 48.3 million tons in 2007 by an average increasing
rate of 4% per year.27 After its initial use, less than 10% of plastics can be recycled,28 and the
high estimation of over 60% of solid waste plastics is discarded in open space or landfills
worldwide.27 As reported by World Bank,29 waste plastics occupy 8 – 12% of municipal solid
waste (MSW) produced in different counties. It was also evaluated that the proportion of
global waste plastics in MSW will increase to 9 – 13% in 2025. Nevertheless, the disposal of
waste plastics become a major concern in light of damages towards the environment.
1.3.2 Availability of waste tires
It was found that around 2.8 billion tires were manufactured in 2012, which correspond to 17
million tons.30 There tires will eventually categorized or considered as waste tires.31 In
particular, the considerable tires consumption in Asia and Oceania will make these regions
become the dominant waste tires sources in next two decades. Taking China for example, the
waste tires in 2010 was 5.2 million tons, accounting for 60 wt% of the global waste tires.32
5
Waste tires have a vital influence in the increase of urban waste stream and threaten the
environment. In general, approximately 64% of waste tires are disposed in landfills or
illegally dumped or stockpiled; no more than 13% of waste tires can be recycled.33 Since
waste tires are not readily degraded in landfills, yet easily float to the top over time owing to
trapped gases. As a result, the waste tires break landfill covers.6 Incineration of waste tires
liberate the toxic gas, containing carcinogenic and mutagenic chemicals. Within the waste
refinery concept, the valorization of waste tires is a preliminary driving force for research and
innovation.
1.4 Biomass conversion routes
To date, the numerous routes for transforming biomass into biofuels are gradually divided
into two broad categories: biochemical and thermochemical conversion routes.34 Standard
biochemical conversion routes typically contain a mild thermal chemical pretreatment step in
order to enhance the plant cell more amenable to hydrolysis of the most recalcitrant biomass
fractions to soluble sugars in the presence of enzymes.35 The obtained sugars could be
catalytically36, 37 or biologically38 upgraded. Conversely, thermochemical conversion routes
include direct combustion, gasification, liquefaction, and pyrolysis, using significantly
harsher to depolymerize biomass.
Pyrolysis is a thermochemical conversion process which runs at 350−650 °C in the absence
of oxygen to produce solid, liquid, and syngas.39 The technology can be grouped into two
6
categories depending on different heating rate and temperature. Slow pyrolysis is heating the
biomass to 350 to 600 °C on heating rates of 0.1 to 1 °C/s, with a reaction time lasting
hours.40 Research has found that the product yields approximately equal the quantities of
solid, liquid, and syngas.39, 41 With longer reaction time and lower heating rate, more char is
formed during the pyrolysis.42 Fast pyrolysis was first developed on fluidized bed reactor in
early 1990s.43 Compared with slow pyrolysis, fast pyrolysis is one effectual pathway for the
conversion of biomass to generate liquid products under the condition of higher reaction
temperature (400−650 °C), higher heating rates (10−1000 °C/s), and short residence time
(seconds to minutes).44, 45 Elevated-temperature thermal decomposition of biomass in an inert
gas environment followed by subsequent condensation of vapors results in receiving a large
fraction of biomass-derived pyrolysis liquids.46
Microwave-assisted pyrolysis technology is one of the most ideal methods of enhancing and
accelerating chemical reactions due to effective heat transfer profiles through microwave
irradiation.47 In comparison with conventional thermal heating with certain drawbacks
involving heat transfer resistance, heat loss to surroundings and damage to reactor walls
mostly as a result of continuous electrical heating;48 the microwave pyrolysis which has the
potential for fast and selective heating, energy and cost reduction could overcome these
limitations.49-51 The microwave-induced pyrolysis method has been successfully investigated
from plant residues,11, 52-54 wood,55 sewage sludge 56, 57 and microalgae.58, 59
7
Table 1.1 Basic elementary composition of fast pyrolysis oil and crude oil.1
Composition
Pyrolysis oil
Crude oil
Water (wt%)
15 – 30
0.1
pH
2.8 – 3.8
–
Density (kg/L)
1.05 – 1.25
0.86 – 0.94
Viscosity 50 °C (cP)
40 – 100
180
HHV (MJ/kg)
16 – 19
44
C (wt%)
55 – 65
83.86
H (wt%)
5–7
11 – 14
O (wt%)
28 – 40
<1
N (wt%)
< 0.4
<1
S (wt%)
< 0.05
<4
Ash (wt%)
< 0.2
0.1
H/C
0.9 – 1.5
1.5 – 2.0
O/C
0.3 – 0.5
~0
1.5 Bio-oil upgrading
Pyrolysis liquid oil owns several potential applications on the base of its remarkable merits
such as limited content of sulfur and nitrogen, fuel for static application and enriched
commodity chemicals.60 Nevertheless, there are some detrimental properties of bio-oil such
as poor thermal and chemical stabilities and high viscosity as a function of high oxygen
content; the energy density of fast pyrolysis oil is much lower than that of petroleum fuels
and pyrolysis oil is reactively immiscible with hydrocarbon fuels.44 Table 1.1 gives a detailed
indicators between fast pyrolysis oil and crude oil. It is therefore unable for further use in
combustion engines. In particular, most organic chemicals in pyrolysis oil are in low content,
leading to their recovery not only technically difficult but also economically unattractive.46
8
Accordingly, pyrolysis oil has to undergo downstream catalytic upgrading to reject oxygen
functional groups prior to be used as a conventional transportation fuel.
Bio-oil upgrading is an arduous task because of the complex compounds in bio-oils. The
phases of bio-oil upgrading can be categorized into the conceptual elements of feed
purification, chemical modification, heteroatom removal, cracking, and separations.61
Improving the quality of pyrolysis oil toward properties identical to those of hydrocarbon
fuels is necessary.62 Efforts to expel the high oxygen content of bio-oil by certain upgrading
technologies is a priority.41,
63
A variety of studies have been undertaken to garner this
objective through upgrading techniques. Among them, catalytic deoxygenation has been
widely explored for upgrading bio-oil and dominantly includes two techniques: catalytic
cracking and hydrodeoxygenation (HDO).14 Catalytic cracking is a technique that involves
the introduction of solid acid catalysts to the pyrolysis process and it is conducted at the
atmospheric pressure without the use of hydrogen.64 However, catalytic cracking of pyrolysis
oil seem to be unpromising due to low grade products obtained with a low carbon yield and
high coke formation (8 – 25% wt% regarding the feed), resulting in a short catalyst lifetime.6
HDO of bio-oil adopts petroleum oil hydrotreating technique and gain desired hydrocarbons
at a pressurized hydrogen atmosphere by using supported metal catalysts.65, 66 This process
has been well attracted and developed owing to the substantial increase in high grade
products.67 Nevertheless, HDO process requires high-severity reaction conditions (such as
high reaction temperature and hydrogen pressure), improving the complexity of the reactor.
High operating cost related to noble catalysts used, significant catalyst deactivation, and
9
considerable hydrogen consumption are also endured in HDO process.68
An alternative way is to directly upgrade pyrolytic volatiles assisted by a catalyst prior to
quenching volatiles to attain bio-oil, which is called catalytic fast pyrolysis (CFP).
The
in-situ catalytic fast pyrolysis (in-situ CFP) is conducted by combing the catalyst with
biomass; when the catalyst located in a downstream reactor solely catalyzes the pyrolytic
vapors, the process is referred as ex-situ catalytic fast pyrolysis (ex-situ CFP).69 CFP is more
amenable to directly convert biomass into high quality bio-oils with enhanced stability, since
it can avoid polymerization and re-evaporation of pyrolysis oil.70-77 Zeolite-based catalysts
(e.g., ZSM-5) have been recognized as the highest-efficiency catalysts in CFP to produce
considerable petrochemicals (aromatics and olefins).78-82
1.6 Catalytic co-pyrolysis of biomass with polymers
1.6.1 The importance of catalytic co-pyrolysis
However, even in the presence of highly efficient catalyst, catalytic pyrolysis of
lignocellulose can solely manufacture low carbon yield (10 – 30%) of aromatic hydrocarbons;
large amounts of solid residues, including both biochar and coke (carbon yields usually above
30%), are achieved in the process.1,
12, 81, 83, 84
As aromatics are considered as the high
energy-density hydrocarbons production,85 such low carbon yield of aromatic hydrocarbons
from catalytic pyrolysis is not cost-effective to scale up the process in a biorefinery. Further,
the process is commonly plagued by the high production of coke deposited on the catalyst
because the coke could rapidly deactivate the catalyst and reduce its lifetime, resulting in the
10
catalytic process to be impractical.86 Therefore, these are the huge challenges in the face of
commercializing the catalytic pyrolysis for the production of renewable petrochemicals.
It is discerned that the petrochemicals with the low carbon yield of aromatics and high coke
formation are mainly associated with the oxygen-enriched intrinsic nature and hydrogen
deficiency of lignocellulosic biomass.1, 83, 87 In addition, the hydrogen to carbon effective
(H/Ceff) ratio plays a significant role in coke formation and converting efficiency of biomass
into advanced biofuels.88-90 Thus the hydrogen-deficient (H/Ceff usually less than 0.3)
biomass produces low carbon yields of petrochemicals and large formation of coke when the
lignocellulose were transformed over zeolite catalysts.78, 80 In order to improve the carbon
efficiency of aromatics and minimize the coke formation, it is reasonable that the
incorporation of high H/Ceff ratio co-reactants with lignocellulosic biomass in the catalytic
pyrolysis could help mitigate these issues. It is observed that co-feeding of lignocellulosic
biomass with hydrogen-rich feedstock in the catalytic pyrolysis can modify the reaction
mechanism of oxygen removal by substituting decarbonylation and decarboxylation with
dehydration.84, 89, 91, 92
Synthetic polymers (e.g., waste plastics) represent a cheaper and abundant hydrogen sources,
which can be used to improve carbon efficiency of aromatics and lower the coke formation in
the catalytic co-pyrolysis.78, 80, 83, 87, 93 Tremendous quantities of synthetic polymers waste are
generated each year worldwide. For instance, the global generation of waste electrical and
electronic equipment (WEEE) is close to 40 million tons per year.94 These polymers waste
11
cause serious damage toward the environment and have detrimental impact on human
health.80 Yet the current fate to manage polymers waste is generally limited by means of
landfills and incineration.95, 96 The state of the art on tertiary recycling technologies like CFP
offers a promising alternative to valorize polymers waste.97, 98 In light of these premises, the
co-feeding of lignocellulosic biomass with polymers waste in catalytic pyrolysis is
remarkably beneficial for the environment and energy recapture.
1.6.2. Chemistry of catalytic co-pyrolysis
In terms of chemistry and steps for catalytic co-pyrolysis, they are comparable with those of
catalytic fast pyrolysis (CFP). Basically, catalytic co-pyrolysis is also necessitated to be
carried out in the presence of a catalyst. The presence of a catalyst could enhance the targeted
product yield and selectivity in CFP.99 Among the widely used catalysts, zeolite-based
catalysts are particularly attractive due to their performances in the increase of aromatic
yield.71 Likewise, zeolite-based catalysts (e.g., HZSM-5) have been mainly applied as the
catalysts in the catalytic co-pyrolysis of lignocellulosic biomass with plastics.78, 79, 89, 93, 100
However, most previous studies regarding catalytic co-pyrolysis of biomass with plastics
have been usually performed by using a high catalyst to feed ratios (e.g. 10:1) to maximize
petrochemical yields. The co-feeding of biomass and plastics are commonly subject to
catalytic pyrolysis process at the catalytic temperature of 400 – 700 °C.79, 87, 89
It is noticed that current catalytic co-feed catalytic pyrolysis of biomass and plastics are
frequently conducted in fixed-bed reactors.6 By contrast, the continuous feeding fluidized-bed
12
reactor80, 90 and the tandem micro-pyrolyzer reactor with downstream packed-bed reactor87
have been applied in the catalytic co-pyrolysis of biomass with plastics as well. In addition, a
couple of groups implemented the co-feeding of biomass with plastics in CFP by using
micro-pyrolysis reactors.78, 79, 89, 93, 100 There were a positive synergy between biomass and
plastics in the augment of aromatic yield and solid residues reduction. For the slow catalytic
co-pyrolysis conducted in abovementioned studies, it is detrimental to biomass conversions;
nonetheless, the extended reaction time not only enhances the degradation of polymers to
small molecular, but also promotes the catalytic reforming of biomass-derived oxygenates
and polymers-derived olefins to yield aromatic production.
1.6.3 Mechanism of catalytic co-pyrolysis
The reaction mechanism of catalytic co-pyrolysis can be relatively complex due to various
types of materials introduced in the process.89 The catalytic co-pyrolysis mechanism can be
classified into two categories: the mechanism among biomass and polymers during thermal
degradation, and the mechanism among the pyrolytic volatiles at the catalytic sites.87 For
catalytic co-pyrolysis of biomass and plastics model compounds, it was observed that there
was a positive synergy for aromatic production in the co-feeding of cellulose and LDPE in
CFP.93 The phenomenon was mainly attributed to the interactions between cellulose-derived
oxygenates and LDPE-derived olefins.78 It was discerned that cellulose-derived oxygenates
could react with LDPE-derived light olefins in the presence of ZSM-5 catalyst.78 Fig. 1.2
depicts the reaction pathway for the conversion of cellulose and LDPE into aromatic
production in the catalytic co-pyrolysis over ZSM-5 catalyst. It can be seen that furan
13
compound (e.g., furan and furfural) derived from cellulose could react with light olefins
(ethylene and propylene) evolved from LDPE to form aromatics (e.g, BTX) through
Diels-Alder reactions followed by dehydration reaction. In the catalytic co-pyrolysis, furan
compounds and light olefins serves as the diene and dienophile compounds, respectively.88, 91,
101
The Diels-Alder reactions followed by dehydration reaction can enhance the aromatics
yield and reduce the coke formation via polymerization of furans.78, 102 It was also evidenced
that LDPE-derived hydrocarbons (such as olefins and alkanes) can also act as the hydrogen
donor for cellulose-derived oxygenates, resulting in the decrease of coke formation in the
zeolite-catalyzed conversions.103, 104
Fig. 1.2 Schematic presentation for the productions of aromatics from catalytic co-pyrolysis
of cellulose with LDPE in the presence of ZSM-5 catalyst.83, 93
14
Fig. 1.3 Possible reaction scheme regarding the formations of naphthalene and indene from
catalytic co-pyrolysis of cellulose with PS by using HZSM-5 as the catalyst.88, 89
Comparing to PE and PP, PS cannot produce enough small olefins for biomass-derived
oxygenates to produce aromatics in the catalytic co-pyrolysis. The interactions among
biomass and plastics proceeds in a different manner in the co-feed catalytic pyrolysis.
Generally, naphthalenes can be produced through the Diels-Alder condensation, dehydration
of benzene and furans, and a cascade of alkylation among toluene and intermediate allene.88
Based on the mechanism of naphthalenes formation in CFP, possible reaction scheme for the
formations of naphthalene and indene from catalytic co-pyrolysis of cellulose with PS by
using HZSM-5 as the catalyst is sketched in Fig. 1.3. Styrene as a major product from the
thermal degradation of PS could undergo the successive alkylation with allene derived from
furans to form indene in the presence of the catalyst.88 As such, the indene produced could
15
further react with allene, giving rise to the generation of naphthalene.88
1.7 Hydrogenation process of aromatics
It is widely known that the oxygen content in the fast pyrolysis oils was removed
through HDO process in the form of H2O. In this regard, large amount of costly
hydrogen was consumed by the dehydration reaction, which significantly reduces the
hydrogenating efficiency. Since advanced jet fuels from HDO process are expected to
be enlarged in a biorefinery scale, high capital costs caused by tolerance of severe
condition
and
low
selectivity
of
liquid
products
render
these
processes
uneconomical.105, 106 Accordingly, the production of jet fuels from renewable biomass
resources calls for ideal technologies with efficient solid-phase catalysts to make the
processes economically feasible under a mild, environmentally friendly reaction
condition.
In contrast, oxygen content of bio-oil can be partially or even completely eliminated
by zeolite cracking during fast pyrolysis of biomass.107 Carlson et al. have reported
that biomass-derived feedstock can be directly converted into aromatics with ZSM-5
as catalyst in a single catalytic pyrolysis step.71, 108 Lei and his colleagues have focused
on the production of aromatics through catalytic microwave-induced pyrolysis of
biomass-derived feedstock.15,
82
Up to 92.60% selectivity towards aromatic
hydrocarbons was obtained, belonging to jet fuel range (C8 – C16) aromatics. In this
sense, the certain range aromatics with low oxygen content are prone to be
16
hydrogenated into cycloalkanes together with minor aromatics under low-severity
conditions. The hydrogenation process is regarded as a hydrogen saving process, in
which hydrogen is converted in the form of cycloalkanes rather than water.
1.8 Objectives
The goal of this proposal is to improve the carbon yield of aromatics and reduce coke
formation in the catalytic microwave pyrolysis and obtain renewable cycloalkanes and
aromatic hydrocarbons for jet fuels.
The specific objectives and plans include:
1) To determine the production of catalytic microwave pyrolysis of cellulose and reaction
pathway;
2) To explore the production of catalytic microwave pyrolysis of low-density polyethylene
and reaction mechanism;
3) To investigate the interaction between cellulose and low-density polyethylene during
catalytic microwave pyrolysis;
4) To compare diverse solvents effect on the hydrogen saving process and produce jet fuel
range cycloalkanes from C8 – C16 aromatic hydrocarbons;
5) To maximize the carbon yield of cycloalkanes for jet fuels from catalytic microwave
pyrolysis of distinct feedstocks and carry out the techno-economic analysis of all feedstocks;
6) To study thermal behaviors on catalytic co-pyrolysis of biomass with plastics and
understand the reaction kinetics for the process.
17
1.9 Outlines
This dissertation includes thirteen chapters. The first chapter is a general introduction
providing the background of catalytic pyrolysis, hydrogenation process, and objective of this
study. The second to twelve chapters are four main parts of the research studies. The chapter
thirteen is the summary of the studies and future research directions. The summaries of
eleven main chapters are introduced below.
Chapter two: A novel pathway was investigated to produce gasoline-range aromatics and
hydrogen-enriched fuel gas by microwave-induced pyrolysis of cellulose integrated with
packed-bed catalysis in the presence of solid phase catalyst. The employed catalyst was
well-promoted ZSM-5 after the couplings of hydrothermal and calcined treatments,
completely converting volatile vapors derived from microwave pyrolysis into aromatics and
non-condensable gases. A central composite experimental design (CCD) was employed to
investigate the effects of catalytic temperature and inverse weight hourly space velocity
(WHSV)-1 on the pyrolysis-oils composition. It was observed that the chemical compounds of
the upgraded bio-oils from catalytic microwave pyrolysis of cellulose were aromatic
hydrocarbons, phenols, and aromatic oxygenates. Aromatic hydrocarbons accounted for the
largest selectivity of these compounds were in the range from 82.93 to 96.60% in bio-oils
depending on alterations of catalytic conditions. Up to 48.56% selectivity towards aromatics
in the upgraded bio-oil belongs to gasoline-range aromatics at the mild condition. The
18
maximum selectivity of aromatic hydrocarbons (96.60%) was gained at packed-bed
temperature of 500 °C and WHSV-1 of 0.067 h.
Chapter three: The microwave-induced pyrolysis of low-density polyethylene (a model of
waste plastics) towards its conversion into biofuels was investigated using ZSM-5 as a
catalyst, generating significant amounts of gasoline-range hydrocarbons. A central composite
experimental design (CCD) was done to investigate the effects of catalytic temperature and
reactant to catalyst ratio on the pyrolysis-oils composition and to achieve the maximum liquid
yield. The optimized condition for maximizing the yield of upgraded oil (32.58 wt. %) was at
450°C and reactant to catalyst ratio of 2. GC-MS analysis showed that mono-ring aromatic
hydrocarbons were enriched and became the most abundant compounds which varied from
74.73% to 88.49% in upgraded pyrolysis-oils, depending on the catalytic pyrolysis
conditions.
Chapter four: A novel pathway was investigated to produce jet fuel range paraffins (alkanes)
and aromatics by catalytic microwave-induced pyrolysis of intact biomass (Douglas fir)
integrated with the hydrotreating upgrading process. The proof-of-principle for the
consecutive two-step process for converting lignocellulosic biomass into jet fuel range
paraffins and aromatics involves the use of the well-promoted ZSM-5 in the process of
catalytic microwave pyrolysis and Raney nickel in the hydrotreating process.
The
production of desired C8 – C15 aromatics was achieved from catalytic depolymerization of
intact biomass at 375 °C. Up to 12.63% selectivity of C8 – C15 paraffins and 19.48%
19
selectivity of hydro-aromatic hydrocarbons were obtained from the hydrotreating of parent
oil under a low-severity condition (for 2 h). A central composite experimental design (CCD)
was employed to investigate the effects of reaction temperatures and initial pressures on the
composition of hydrotreatd oils. We observed that increasing reaction temperature and initial
pressure with prolonged time could enhance the hydrogenation and hydrogenolysis reactions
to form jet fuel range paraffins and aromatics.
Chapter five: The consecutive processes for converting lignocellulosic biomass into jet fuel
range cycloalkanes principally involved the use of the well-promoted ZSM-5 in the process
of catalytic microwave-induced pyrolysis and Raney nickel catalysts in the hydrogen saving
process. Up to 24.68% carbon yield of desired C8 – C16 aromatics was achieved from
catalytic microwave pyrolysis at 500 °C. We observed that solvents could assist in the
hydrogenation reaction of naphthalene; and the optimum result for maximizing the carbon
selectivity (99.9%) of decalin was from the reaction conducted in the n-heptane medium. The
recovery of organics could reach ~94 wt. % after the extracting process. These aromatics in
the n-heptane medium were eventually hydrogenated into jet fuel range cycloalkanes. Various
factors were employed to determine the optimal result under mild conditions. Increasing
catalyst loading, reaction temperature, and prolonged time could enhance the hydrogenation
reactions to improve the selectivity of jet fuel range cycloalkanes. Three types of
hydrogenation catalysts (NP Ni, Raney-Ni 4200, home-made Raney Ni) were chosen to
evaluate the catalytic performance.
20
Chapter six: The consecutive processes for manufacturing high-density cycloalkanes
primarily included the catalytic microwave-induced pyrolysis of diverse lignocellulosic
biomasses (hybrid poplar, loblolly pine, and Douglas fir) over a well-promoted ZSM-5 and
hydrogenation process in the present of the Raney nickel catalyst. Two variables (catalytic
temperature and catalyst to biomass ratio) were employed to determine the optimal condition
for the production of C8 – C16 aromatics in the catalytic microwave pyrolysis. The maximum
carbon yield of desired aromatics was 24.76%, which was achieved from catalytic microwave
pyrolysis of hybrid poplar at 500 °C with the catalyst to biomass ratio of 0.25. We observed
the aromatics derived from catalytic microwave pyrolysis in the n-heptane medium were
completely hydrogenated into renewable high-density cycloalkanes for jet fuels. In the
hydrogenation process, increasing catalyst loading and reaction temperature could promote
the selectivity of high-density cycloalkanes. Results indicated that the hybrid poplar was the
optimal feedstock to obtain the highest selectivity (95.20%) towards high-density
cycloalkanes. The maximum carbon yield of cycloalkanes-enriched hyrogenated organics
based on hybrid poplar was 22.11%.
Chapter seven: The consecutive processes for manufacturing JP-5 navy fuel principally
included the catalytic microwave-induced degradation of low-density polyethylene (a model
compound of waste plastics) and the hydrotreatment of obtained liquid organics. The catalytic
microwave degradtation was conducted at the catalytic temperature of 375 °C and catalyst to
feed ratio of 0.1. The carbon yield of the liquid organics from the catalytic microwave
degradation was 66.18%, mainly consisting of a mixture of aromatic hydrocarbons and
21
aliphatic olefins. Several variables, such as initial pressure and catalyst to reactant ratio, were
employed to determine the optimal condition for the production of alternative jet fuels in the
hydrotreating process. We observed that the aromatic hydrocarbons and aliphatic olefins as
the precursors of jet fuels could be converted into jet fuel range aliphatic alkanes and
cycloalkanes. The hydrotreated organics from the experiment conducted at the reaction
temperature of 250 °C for 2 h, including 31.23% selectivity towards aliphatic alkanes, 53.06%
selectivity towards cycloalkanes, and 15% selectivity towards remaining aromatic
hydrocarbons, which were consistent with the specifications of JP-5 navy fuel.
Chapter eight: The consecutive processes principally proceeded via the catalytic
microwave-induced degradation of low-density polyethylene (a model compound of waste
plastics) followed by hydrogenation of raw organics. The catalytic microwave degradation
was conducted at the catalytic temperature of 375 °C and catalyst to feed ratio of 0.1 and 0.2
to manufacture distinct contents of aliphatic and cyclic hydrocarbons. The carbon yields of
the raw organics from the catalytic microwave degradation was 66.18 and 56.32%. Several
variables were employed to determine the suitable conditions for the production of alternative
jet fuels in the hydrogenation process. We observed that the unsaturated hydrocarbons as the
precursors of jet fuels could be converted into jet fuel range aliphatic and cyclic alkanes.
Since n-heptane medium could assist in the hydrogenation of unsaturated hydrocarbons, all
hydrogenation experiments were conducted in such medium. With the presence of 10 and 20
wt% well-promoted ZSM-5 catalyst, the overall carbon yields of hydrogenated organics with
respect to raw plastics were approximately 54 and 63%, respectively. The alterations of
reaction conditions did not have an effect on the overall carbon yield of hydrogenated
organics. The raw organics from catalytic microwave degradation by using 10 wt%
22
well-promoted ZSM-5 catalyst could be converted into alternatives of JP-5 navy fuel or
additives of Jet A and JP-8 in the presence of 10 wt% Raney Ni catalyst at 200 °C for 2 h.
The raw organics from catalytic microwave degradation by using 20 wt% well-promoted
ZSM-5 catalyst could transformed into high energy-density jet fuels (e.g. RJ-5 and JP-10)
with regard to various variables. Home-made Raney Ni catalyst showed much higher
catalytic performance than that of as-purchased Raney Ni 4200 catalyst.
Chapter nine: Enhanced carbon yields of renewable alkanes for jet fuels were obtained
through a novel pathway from co-feeding of cellulose and low-density polyethylene (LDPE).
The consecutive processes proceeded via the catalytic microwave-induced pyrolysis over
well-promoted ZSM-5 catalyst and hydrogenation process by using home-made Raney Ni
catalyst. It was found that parent ZSM-5 modified by hydrothermal and calcined treatments
resulted in the increase of surface area as well as the formation of more mesopores.
Interestingly, the well-promoted ZSM-5 catalyst had high selectivity toward C8 – C16
aromatic hydrocarbons in the co-feed catalytic microwave pyrolysis. The raw organics with
improved carbon yield (~ 44%) were more principally lumped in the jet fuel range at the
catalytic temperature of 375 °C with the LDPE to cellulose ratio of 0.75. The home-made
Raney Ni catalyst was assayed for hydrogenation of diverse organics species derived from
co-feed catalytic microwave pyrolysis under a low-severity condition. It was observed that
the four species of raw organics in the n-heptane medium were almost completely converted
into saturated hydrocarbons. The overall carbon yield (with regards to co-reactants of
cellulose and LDPE) of hydrogenated organics that mostly match jet fuels was sustainably
enhanced, reaching above 39%. Meanwhile, ~ 90% selectivity toward renewable alkanes for
23
jet fuels was attained.
Chapter
ten:
The
consecutive
processes
proceeded
via
the
co-feed
catalytic
microwave-induced pyrolysis and hydrogenation process. In the co-feed catalytic microwave
pyrolysis by using ZSM-5 as the catalyst, parent ZSM-5 fabricated by hydrothermal and
calcined treatments contributed to the increase of surface area as well as the formation of
more mesopores. Liquid organics with enhanced carbon yield (40.54%) were more
principally lumped in the jet fuel range from the co-feed catalytic microwave pyrolysis
performed at the catalytic temperature of 375 °C with the plastics to biomass ratio of 0.75. To
manufacture home-made Raney Ni catalyst, the BET surface area, pore surface area, and pore
volume of the home-made Raney Ni catalyst were considerably improved when the Ni-Al
alloy was dissolved by the NaOH solution. In the hydrogenation process, we observed the
three species of raw organic derived from the co-feed catalytic microwave pyrolysis were
almost completely converted into saturated hydrocarbons under a low-severity condition. The
improved carbon yield (38.51%) of hydrogenated organics regarding co-reactants of biomass
and plastics predominantly match jet fuels. In the hydrogenated organics, over 90%
selectivity toward jet fuel range alkanes was attained.
Chapter eleven: This study reported a novel pathway to obtain renewable cycloalkanes for
jet fuels from diverse lignocellulosic biomasses and presents an economic analysis of the
combined processes in a small scale biorefinery based on the experimental results. In the
bench-scale experiments, we observed that the maximum carbon yield (24.76%) of desired
24
aromatics in the carbon number range from 8 to 16 was achieved from catalytic microwave
pyrolysis of hybrid poplar at 500 °C, comparing with loblolly pine and Douglas fir used as
feedstock. After the hydrogenation process, the highest selectivity (95.20%) towards jet fuel
range cycloalkanes was obtained. As expected, a scenario using 15 dry tonne woody biomass
per day is developed in a small scale biorefinery; and the hybrid poplar is used as feedstock
to manufacture 0.25 million gallons renewable cycloalkanes for one year. The assessments
showed that the plant including microwave assisted ex-situ catalytic pyrolysis (MAECP) and
hydrogenation (HG) systems is profitable. The equipment costs have the largest contribution
to the total capital investment, whereas the feedstock and chemicals costs have the largest
contribution to the total annual production cost. It is estimated that the Return of Investment
(ROI) is 30.33% for each year. Sensitivity analysis indicates that liquid organics yield and
cycloalkane selling price significantly impact the ROI.
Chapter twelve: The present study aims to investigate the thermal decomposition behaviors
and kinetics of biomass (cellulose/ Douglas fir sawdust) and plastics (LDPE) in a
non-catalytic and catalytic co-pyrolysis over ZSM-5 catalyst by using a thermogravimetric
analyzer (TGA). It was found that there was a positive synergistic interaction between
biomass and plastics according to the difference of weight loss (∆ ), which could decrease
the formation of solid residue at the end of the experiment. The DTG curves demonstrated
that the addition of catalyst tended to reduce the degradation temperature. The first order
reaction model well fitted for both non-catalytic and catalytic co-pyrolysis of biomass with
plastics.
As for the activation energy (E) and pre-exponential factor (A), the addition of
25
plastics in the non-catalytic pyrolysis of biomass could appreciably decrease the activation
energy (E). Moreover, the presence of ZSM-5 catalyst could further decrease the activation
energy (E) based on the result of the blends without catalyst. The activation energy (E) of
Cellulose-LDPE-Catalyst and DF-LDPE-Catalyst are only 89.51 and 54.51 KJ/mol,
respectively.
The kinetics analysis showed that adding catalyst doesn’t change the
decomposition mechanism.
26
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99.
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1387-1390.
100.
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101.
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Lobo, W. Fan and P. J. Dauenhauer, ACS Catalysis, 2012, 2, 935-939.
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102.
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2011, DOI: 10.1039/c0ee00341g, 145-161.
33
CHAPTER TWO
RENEWABLE GASOLINE-RANGE AROMATICS AND HYDROGEN ENRICHED FUEL GAS FROM CELLULOSE VIA CATALYTIC
MICROWAVE-INDUCED PYROLYSIS
2.1 Abstract
A
novel
pathway
was
investigated
to
produce
gasoline-range
aromatics
and
hydrogen-enriched fuel gas by microwave-induced pyrolysis of cellulose integrated with
packed-bed catalysis in the presence of solid phase catalyst. The employed catalyst was
well-promoted ZSM-5 after the couplings of hydrothermal and calcined treatments,
completely converting volatile vapors derived from microwave pyrolysis into aromatics and
non-condensable gases. A central composite experimental design (CCD) was employed to
investigate the effects of catalytic temperature and inverse weight hourly space velocity
(WHSV)-1 on the pyrolysis-oils composition. It was observed that the chemical compounds of
the upgraded bio-oils from catalytic microwave pyrolysis of cellulose were aromatic
hydrocarbons, phenols, and aromatic oxygenates. Aromatic hydrocarbons accounted for the
largest selectivity of these compounds were in the range from 82.93 to 96.60% in bio-oils
depending on alterations of catalytic conditions. Up to 48.56% selectivity towards aromatics
in the upgraded bio-oil belongs to gasoline-range aromatics at the mild condition. The
maximum selectivity of aromatic hydrocarbons (96.60%) was gained at packed-bed
temperature of 500 °C and WHSV-1 of 0.067 h. Gaseous results show that hydrogen was the
34
dominant composition, occupying approximately 40 vol.%. The high amounts of
gasoline-range aromatics and valuable hydrogen is attributed to the technologies of
microwave-assisted pyrolysis and ex-situ catalysis. These findings from this study pave a
new route for biorefinery industries to produce developed products (aromatics and
hydrogen-rich gas) through microwave-induced technologies.
Keywords: Catalytic microwave-induced pyrolysis; aromatics; hydrogen; cellulose; ZSM-5
catalyst
2.2 Introduction
Growing concerns about global climate change and rapid depleting fossil-based petroleum
sources have spurred interest in the utilization of renewable resources for fuels and high
value-added chemicals.1 Cellulose-based carbohydrates are the ubiquitous compounds, which
have been realized to manufacture biofuels such as bioethanol instead of petroleum fuels
derived from fossil sources. Although bioethanol produced from cellulosic biomass has
already been added in automotive engines for improving the combustion, the low energy
density (2/3 of gasoline) makes it impossible completely replace petroleum-derived gasoline.2
In contrast, renewable gasoline-range aromatics rather than ethanol will be drop-in
replacements for petroleum gasoline if current technological innovations contribute to
competitive production costs.2
35
Aromatics have numerous industrial applications, and can be used as fuels additives for
octane enhancement or chemical precursors for a variety of plastics and solvents.3 The
demand for aromatics exceeds 200 billion lbs/year with a total market of close to $100 billion
in 2010 and compound annual growth rates from 3.1-5.4% (Chemicals Economic Handbook).
Aromatics are crucial components of gasoline and fuel oils, up to 85% aromatics are
dominant in Kuwaiti gasoline.4 Current aromatics are exclusively derived from petroleum
sources. Thus discovering a green route to convert renewable resources into aromatics will
pave a new avenue towards renewable resources utilization.
Numerous attempts have been made to substitute petroleum-based aromatics; catalytic fast
pyrolysis (CFP) of biomass-derived carbohydrates (e.g. cellulose) is a potential process used
for converting biomass into aromatics.5-10 However, these conventional techniques were
usually conducted with the harsh condition such as high consumption of energy, elevated
reaction temperature, and high loading of catalysts, resulting in infeasible commercialization
in biorefineries comparing with commercial formulations based on petroleum derived
aromatics. Therefore the production of green aromatics from renewable biomass resources
calls for an ideal technology with an efficient solid-phase catalyst to make the process
economically feasible under a mild, environmentally friendly reaction condition. Instead,
microwave-assisted pyrolysis technology has been proven to be one of the most promising
methods for enhancing and accelerating chemical reactions due to effective heat transfer
through
microwave
irradiation.11
In
comparison
36
with
traditional
pyrolysis,
microwave-induced pyrolysis encloses the potentials of fast and selective heating, easy
control of reaction conditions, low reaction temperatures and energy requirements.12, 13
Aforementioned studies were mainly engaged in the topics of production of green aromatics
in the presence of ZSM-5 zeolite without any treatments. In fact, moderate treatments can be
empolyed to tailor zeolite properties for modifications of porosity and acidity, which is
essential for improving product selectivity and the quality of the bio-oil.14 Groen et al
claimed that hydrothermal treatment is usually conducted to adjust the acidity of ZSM-5
zeolite and improve its stability.15 The extra-framework aluminum species derived from
hydrothermal treatment also act as Lewis acid centers and improve certain catalytic
activities.16 Furthermore, upon removal of the water inside the zeolite pores by calcination at
elevated temperature, strong electrostatic fields are built up.16 Under the condition of the
electrostatic fields, remaining water molecules dissociate into a proton bounding to Brønsted
acid sites and an OH group. Thereupon mild hydrothermal and calcined conditions can favor
the generation of catalytic sites specific efficiency.
Unfortunately, gas from catalytic fast pyrolysis was usually paid less attention by current
researches.17-20 The major constituents in the gaseous products including hydrogen, carbon
monoxide, methane, and other light hydrocarbons can be viewed as significant reactants in a
majority of refinery processes to produce energy and liquid fuels. Costly hydrogen as one of
vital co-products can also be supplied for hydrotreating of aromatics to produce jet fuels.
Accordingly, hydrogen-rich gas products from catalytic pyrolysis should be considered to
37
reach the goal for price-competitive renewable aromatics. Nonetheless, there has not been a
successfully
attempt
to
produce
high-concentration
aromatics
integrated
with
hydrogen-enriched fuel gas from catalytic microwave assisted pyrolysis of cellulosic biomass.
Herein, this study aims to fill this knowledge gap and investigate the production of green
aromatics and hydrogen-enriched fuel gas by catalytic microwave pyrolysis of cellulosic
biomass under distinct reaction conditions including catalytic temperature and inverse weight
hourly space velocity (WHSV)-1. Well-promoted ZSM-5 via mild hydrothermal and calcined
treatments was first used as a catalyst for aromatics and hydrogen-rich gas production in the
process of microwave-mediated pyrolysis.
2.3 Experimental
2.3.1 Material
The cellulose (CAS number 9004-34-6) is purchased from Sigma-Aldrich Corporation (St.
Louis, MO, USA). Cellulose is in the form of microcrystalline powders and particle sizes of
cellulose are averaged at 50 µm. Parent ZSM-5(SiO2/Al2O3 Mole Ratio: 50) is purchased
from Zeolyst International, USA.
2.3.2 Catalyst preparation
In order to determine the optimum catalyst, various treatments upon ZSM-5 are conducted
and compared to the parent ZSM-5. The activity of parent ZSM-5 is improved by suffering
either hydrothermal or calcined treatment. In this regard, hydrothermal treatment is used to
modify the parent ZSM-5 catalyst. Under the gentle stirring, parent ZSM-5 powder is added
38
into deionized water (mass ratio=1) at 60 °C. After addition, the mixture is kept on stirring
for 2 h under this condition. The slurry is then dried at 105 °C till constant weight. The
sequential process is the catalyst calcination: parent ZSM-5 or ZSM-5 treated is calcined at
550°C for 5 h in a muffle furnace. The catalysts are partially pelletized and sieved to 20 – 40
mesh.
2.3.3 Catalytic microwave pyrolysis and analysis
The feedstock is air dried at 105 °C for 24 h to remove the physically bound moisture,
prior to conducting the experiments. Fixed loading of cellulose (20g) loaded in a 500
mL quartz flask is placed inside the Sineo MAS-II batch microwave oven (Shanghai,
China) by a constant microwave power setting (700 W). 0.05g activated carbon
powder is used as the absorber for the microwave-assisted pyrolysis of cellulose. All
reactions of microwave pyrolysis are conducted at the temperature of 480 °C for 10
min. The pyrolysis volatile vapors from the flask passed through a packed bed
catalysis reactor which is filled with catalyst. The packed-bed reactor customized is
constructed of quartz and externally heated by a heating tape. A thermocouple is
introduced between the reactor and the heating tape to measure catalytic temperature.
The separate heating regimes ensure that distinctly separate temperatures in the
microwave oven and packed-bed reactor are maintained.
39
Elemental analysis (C, H and N) of cellulose, char, and coke deposited on spent
catalysts is conducted using a 2400 Series II CHN/O Elemental Analyzer
(PerkinElmer, USA). Oxygen content is calculated by difference.
The textural properties of the catalyst are determined by means of N2
adsorption–desorption (Micromeritics TriStar II 3020 Automatic Physisorption
Analyzer). Fresh catalysts are degassed in vacuum at 300 °C for 1 h. Spent catalysts
are degassed at 150 °C under vacuum for 2 h. The Brunauer–Emmett–Teller equation
is applied to calculate the specific surface area using adsorption data at p/po=
0.05–0.25. The pore volume is evaluated by using the Barrett–Joyner–Halenda (BJH)
method.
The acidity of the catalyst is measured by temperature-programmed desorption (TPD)
of ammonia with a Micromeritics AutoChem II 2920 Chemisorption Analyzer
equipped with a PFEIFFER mass spectrometer. Fresh catalysts are saturated with NH3
at room temperature in a flow of 10% NH3 in nitrogen. After NH3 saturation the
weakly bound NH3 is desorbed prior to the measurement at 120 °C for 1 h with a He
flow rate of 25 ml/min. The desorption curve is then attained at a heating ramp of 10
°C/min from 120 °C to 550 °C at a He flow rate of 25 ml/min.
The chemical composition of the bio-oils is characterized and qualified by Agilent
7890A GC-MS (GC–MS; GC, Agilent 7890A; MS, Agilent 5975C) with a DB-5
40
capillary column. The GC is first programmed to heat to 45 °C for 3 min followed by
heating to 300 °C at a rate of 10 °C/min. The injection sample size is 1 μL. The flow
rate of the carrier gas (helium) is 0.6 mL/min. The ion source temperature is 230 °C
for the mass selective detector. Compounds are identified by comparing the spectral
data with that in the NIST Mass Spectral library. The area percent of compounds
obtained from GC/MS results is utilized to predict products selectivity.
The moisture content in the bio-oils is determined by a Karl Fischer (KF) compact
titrator (V20 Compact Volumetric KF Titrator, Mettler-Toledo).
The gaseous product is collected in a 1 L Tedlar gas bag and then offline analyzed by
an INFICON 3000 Micro-GC (INFICON Inc., Santa Clara, CA, USA) system with a
thermal conductivity detector (TCD). A standard gas mixture consisting of H2, N2,
CH4, CO, CO2, C2H4, C2H6, and C3H6 is used to calibrate the yield of non-condensable
gas. Alkanes and olefins (>C4) are either not detected or negligible in this research.
2.3.4 Experimental methods and data processing
A central composite experimental design (CCD) is employed to optimize the process
conditions and product yields distribution (Table 2.1). The catalytic temperature (X ,
°C) and inverse weight hourly space velocity, WHSV-1 ( X , h ), are chosen as
independent variables. WHSV-1 is defined as the mass of catalyst used in the
packed-bed reator divided by the mass flow rate of pyrolysis volatile vapors. In this
41
experiment, the mass of catalyst related to reactant to catalyst ratio is varied from 4.54
to 12.58 g, while the mass flow rate is kept constant. The WHSV-1 and packed-bed
temperature are ranged from 0.045 to 0.126 h, 269 to 481 °C, respectively.
Table 2.1 Experimental design and product yield distribution.
Runa
Catalytic
Yield (wt.%)
WHSV-1 (h)
temperature (ºC)
Bio-oil
Gas
Char
Coke
C-1
300
0.100
31.55
49.21
16.12
3.12
C-2
300
0.050
33.25
48.14
15.78
2.83
C-3
450
0.100
31.27
51.28
16.07
1.38
C-4
450
0.050
33.41
49.48
15.98
1.13
C-5
269
0.067
34.77
46.01
15.95
3.27
C-6
375
0.067
33.12
49.14
16.04
1.70
C-7
375
0.067
33.11
48.83
16.32
1.74
C-8
375
0.067
32.93
49.49
15.78
1.80
C-9
375
0.126
30.35
52.35
15.65
2.16
C-10
375
0.067
33.04
49.71
15.60
1.65
C-11
481
0.067
30.65
52.73
15.63
0.99
C-12
375
0.067
33.19
49.03
16.04
1.74
C-13
375
0.045
35.20
47.51
15.97
1.32
C-14
500
0.067
30.42
53.06
15.78
0.74
C-15
375
0.054
34.75
48.25
15.46
1.54
C-16
-
-
36.18
46.89
16.93
-
a
C-1 to C-13 was conducted based on central composite design; C-14 and C-15 were
added as the controls; C-16 is the control in the absence of catalyst.
42
The coke mass is determined by the difference before and after catalytic pyrolysis. The
weight of non-condensable gas is calculated using the following equation:
ℎ
=
−
− ℎ
−
(1)
The liquid, gas, coke, and coke yields are calculated by their corresponding masses divided
by the initial reactant mass.
Table 2.2 Textural and acid properties of the ZSM-5 zeolites.a
Catalyst properties
A
B
C
D
E
F
G
BET surface area (m2/g)
386.9
379.6
382.7
405.6
400.5
396.2
306.3
Pore (17< dp/Å < 3000) volume (cm3/g))
0.078
0.082
0.080
0.091
0.107
0.097
0.063
Pore (17< dp/Å <3000) surface area (m2/g)
55.3
47.9
49.0
77.6
85.4
74.1
52.1
Mesopore size (20< dp/Å <500) distribution (%)
98.47
97.41
97.67
98.44
98.95
98.80
70.23
5.7
6.9
6.5
4.7
5.0
5.2
4.6
Average pore size 22
a
A: parent ZSM-5; B: hydrothermally treated ZSM-5 (Powder); C: hydrothermally treated ZSM-5 (Pellet); D: calcined
ZSM-5 (Powder); E: hydrothermally treated and calcined ZSM-5 (Powder); F: hydrothermally treated and calcined
ZSM-5 (Pellet); G: spent F at 500 ºC with WHSV-1 of 0.067 h.
2.4 Results and discussion
2.4.1 Optimization of catalyst
Various treatments upon ZSM-5, compared to the parent ZSM-5, were shown in Table
2.2. BET surface area, pore volume, and pore surface area were significantly improved
by the combined treatments. Furthermore, ZSM-5 promoted by means of the
combined treatments resulted in the generation of secondary porosity (mesoporosity)
in the ZSM-5 zeolite matrix. The average pore size of the modified ZSM-5 (Entry-F in
43
Table 2) is 5.2 nm, which is very close to naphthalene diameter (5.5 nm), therefore
aromatics are prone to be adsorbed in the pores.21 Because the overall acid amounts
can be inferred from the relative peak areas of the NH3 desorption curves; it can be
found that the acidity are similar. For the catalysts modified by hydrothermal
treatment, the maximum of the peak shifted towards low temperatures, which
accompanied with a decrease in percentage of strong acid sites; while catalysts solely
treated by calcination performed same acidity. Hence, optimal ZSM-5 used in the
catalytic reaction was modified by the combined treatments and then pelletized and
sieved to 20 – 40 mesh.
2.4.2 Product yields
The product yield distributions are summarized as a function of catalytic temperature
and inverse weigh hourly space velocity (WHSV)-1 in Table 2.1. It was found that the
yields of bio-oil and gas were in the range from 30.42 to 36.18 wt.% versus 46.01 to
53.06 wt.% respectively, showing inverse trends owing to catalytic conditions. The
water content of most obtained bio-oils was around 10 wt.%, based on cellulose
loading. Results indicate that catalytic temperature had significant effects on both
yields of bio-oil and gas. The yield of bio-oil declined with the increase of the catalytic
temperature, while the yield of gas increased with the similar increasing tendency of
catalytic temperature. It is worth noting that the product yields were also affected by
WHSV-1. Increasing WHSV-1 contributed to improving the yield of gas, whereas the
yield of bio-oil gradually decreased. The optimal condition for maximum bio-oils
44
yield was presented to be a catalytic temperature of 269 ºC and WHSV-1 of 0.067 h.
By contrast, the maximum gas yield was predicted at 500 °C and WHSV-1 of 0.067 h.
When pyrolysis without catalyst, the bio-oil yield went up to 36.18 wt.% as gas yield
simultaneously decreased to 46.89 wt.%. Similar phenomenon was observed by other
studies as a result of catalytic temperature and WHSV-1 affecting product (bio-oil, gas)
yields.23, 24
Solid carbonaceous residue (char and coke) from degraded and catalytic processes can
be distinguished due to the ex-situ catalysis. The char yield remained virtually constant
at 16 wt. %, suggesting that cellulose loaded in the microwave oven was almost
decomposed to volatile compounds. Given the coke deposited on the catalyst resulting
in decrease of the active sites and micropores blockage,25 the carbonaceous
compounds were a crucial element to be taken into consideration in the study of
formation mechanism when zeolites are employed as catalysts. The coke deposition
yields as a function of catalytic temperature and WHSV-1 is also depicted in Table 1,
varying from 0.74 to 3.27 wt.%. As expected, the minor coke deposition was affected
by short residence times set at 10 mins, which abated the formation of coke precursors.
Formation and deposition of coke was attenuated as the catalytic temperature was
increased, declining to below 1 wt.%. It was noticeable that spent ZSM-5 derived from
catalytic process at 500 ºC still showed high BET surface area, pore volume, and pore
surface area (Table 2.2). Results imply that elevated temperature promoted the
cracking reaction of pyrolytic volatiles on the catalyst towards small molecules.
45
Rather, as more catalysts were loaded, more retention time of catalysis will be
provided for pyrolytic volatiles to form more polycyclic aromatics, easily generating
coke on the catalyst.
Fig. 2.1 Chemical composition of bio-oils on the base of catalytic temperature at the
same WHSV-1 of 0.067 h.
2.4.3 Analysis of the bio-oils
2.4.3.1 Chemical composition of bio-oils
In order to further understand the chemical reaction of catalytic microwave pyrolysis of
cellulose to maximize aromatics content, the chemical compositions of bio-oils were
identified and partially quantified by GC/MS. It was observed that the chemical compounds
of the upgraded bio-oils from catalytic microwave pyrolysis of cellulose were aromatic
hydrocarbons, phenols, and aromatic oxygenates. Aromatic hydrocarbons accounted for the
largest selectivity of these compounds were in the range from 82.93 to 96.60% in bio-oils
46
depending on alterations of catalytic conditions. The maximum selectivity of aromatic
hydrocarbons (96.60%) was received at packed-bed temperature of 500 °C and WHSV-1 of
0.067 h as described in the experiment of Run C-14. Phenols in the upgraded bio-oils were
from 2.51 to 13.41% followed by aromatic oxygenates with an identical tendency from 0.53
to 3.66%. Compared with the optimal result with fresh catalyst at 500 °C and WHSV-1 of
0.067 h, spent catalyst reused under this condition still displayed fairly decent catalytic
performance, achieving 67.43% aromatic hydrocarbons together with 9.40% phenols.
2.4.3.2 The effects of catalytic temperature on the chemical composition of bio-oils
Fig. 2.1 shows representative results of bio-oils regarding the chemical compounds on the
base of catalytic temperature at the same WHSV-1 of 0.067 h. In comparison with the bio-oil
obtained by non-catalytic pyrolysis of cellulose, the amount of aromatic hydrocarbons in
bio-oils derived from catalytic microwave pyrolysis significantly increased. It can be seen
that aromatic hydrocarbons were in the range from 88.07% at 269 °C to 96.60% as catalytic
temperature increased to 500°C, implying that a high catalytic temperature favored high
selectivity of aromatic hydrocarbons. Aromatic hydrocarbons obtained were predominantly
comprised of toluene, xylenes, trimethylbenzenes, indane, indene, naphthalene, and their
derivatives. Up to 48.56% selectivity of gasoline-range aromatics (mono-ring aromatic
hydrocarbons) were dominant in upgraded bio-oil at the mild condition of 269 °C, indicating
a decreased trend as the catalytic temperature increased. Among gasoline-range aromatics,
xylenes are the petrochemical intermediates to synthetize fibers, plasticizers and resins, for
which there is an urgent market demand in petrochemical industry.26 On the contrary, it is
47
noted that a total selectivity (39.52 - 49.83%) of naphthalene and methyl substituted
naphthalene, such as methylnaphthalene and dimethylnaphthalene, were received along with
the increment of catalytic temperature from 269 to 500 °C. As reported previously for the
formation of mesopores in the modified ZSM-5 catalysts by the treatment and the
interconnection of the 10-membered ring pores, the pores can provide enough volume for
aromatization and hydrogen transfer reactions. The active sites located at these areas may
serve as the reaction sites for the generation of the double-ring aromatic hydrocarbons. Such
high selectivity of naphthalene and its derivatives can be hydrotreated to cyclic alkanes with
the addition of hydrogen from the process of catalytic microwave pyrolysis. Therefore they
can be considered as the precursors of jet and rocket fuels in the aviation field.
Over 20 wt.% aromatics were obtained under catalytic conditions, which is similar to other
studies.27 It is well known that oxygenates were converted into aromatic hydrocarbons
through a series of decarbonylation, decarboxylation, dehydration, and oligomerization
reactions, rejecting the oxygen content in the form of CO2, CO, H2O.5, 28 It was found that
decarboxylation and dehydration were the dominant deoxygenated reactions in the catalytic
process.
Phenols are vital synthetic materials in chemical industries for synthesizing PF resin,
medicines and so forth.29 A maximum selectivity of phenols (10.26%) including phenol,
methyl phenol, and ethyl phenols was gained at the catalytic temperature of 269 °C; whereas
aromatic hydrocarbons showed the lowest selectivity at this condition.
48
These results
indicate that high catalytic temperature favored the scission of hydroxyl group from phenyl to
form aromatic hydrocarbons. It is also noteworthy that the rest contents of upgraded bio-oils
were aromatic oxygenates, mainly consisting of benzofuran and its derivatives which can be
converted into synthetic resins used in manufacturing and adhesives for food packaging.
The concentrations of four main compounds (p-xylene, naphthalene, phenol, and furfural)
were also quantified by GC/MS. The concentration of p-xylene increased from 0.010 g/mL
without catalysis to 0.080 g/mL at the catalytic temperature of 269 °C, whilst the highest
concentration of naphthalene (0.075 g/mL) was obtained at the catalytic temperature of 500
°C compared with the lowest concentration of 0.019 g/mL in the absence of catalyst. For the
concentration of phenol, the maximum was 0.073 g/mL at the catalytic temperature of 269
°C, while the concentration of phenol in raw bio-oil was only 0.038 g/mL. In the catalytic
groups, there was no furfural detected in comparison with non-catalytic pyrolysis at the
concentration of 0.086 g/mL. Thereafter the furanic compounds which may act as
intermediates, diffused into the modified catalyst pores and went through a series of
decarbonylation, decarboxylation, dehydration, and oligomerization reactions to form
monocyclic aromatics.5, 28 The cooperation of the entrained aromatic component inside the
inorganic zeolite framework possibly created a rather unique reaction route. The inorganic
zeolite framework acted as the dehydration, oligomerization, aromatization and methylation
center for the generation of aromatic hydrocarbons. polymethylbenzenes mainly worked as
the entrained aromatic compounds because naphthenic species usually have an high energy
barrier.30 The crucial reaction mechanism of hydrocarbon pool may be that furans reacted
49
with entrained hydrocarbon species in the catalyst including a series of complicated steps to
form the aromatic hydrocarbons and olefinic products during the catalytic cycle.
Fig. 2.2 Chemical composition of bio-oils as a function of catalytic temperature at the
same temperature of 375 °C.
2.4.3.3 The effects of WHSV-1 on the chemical composition of bio-oils
The chemical composition of bio-oils from catalytic microwave pyrolysis at 375 °C as
a function of WHSV-1 was explained in Fig. 2.2 Compared with the lowest aromatic
hydrocarbons content (26.64%) evolved from pyrolysis without the addition of
catalyst, the content increased gradually from 82.93 to 94.24 % with the increase of
WHSV-1 from 0.045 to 0.126 h, which is consistent with the similar trend claimed by
other researches.23, 31 The phenols fraction was fluctuant caused by WHSV-1, which
first increased from 8.33% (the control) to 13.41% (WHSV-1 of 0.045 h) and then
50
decreased to 5.33% with the increase of WHSV-1 to 0.126 h. As such, the content of
aromatic oxygenates had the same tendency as phenols from 2.35 to 0.53%. In the
upgraded bio-oils, there were no any non-aromatic oxygenates detected. It was confirmed that
pyrolysis volatiles generated were completely catalyzed online when going through the
catalytic reactor. Thus introduction of the catalysts into pyrolysis could convert oxygenated
compounds into aromatics and phenols effectively, which is identical to other research.32
The WHSV-1 also had a significant effect on the species of aromatics. As for the total amount
of gasoline range aromatics and double-ring aromatic hydrocarbons, increasing the WHSV-1
contributed to dramatically enhancing the former selectivity from 38.96 to 49.54 % in
upgraded bio-oils; while results obtained for the latter content were around 43%. Hence most
volatile oxygenates from microwave pyrolysis were converted into mono-cyclic aromatic
hydrocarbons, the mono-cyclic aromatic hydrocarbons may not readily react with oxygenates
to further form polycyclic aromatic hydrocarbons at this catalytic temperature. It was noted
that the concentrations of p-xylene and naphthalene increased appreciably from 0.010 g/mL
and 0.039 at the control to 0.137 g/mL and 0.069 g/mL at the WHSV-1 of 0.126, respectively;
while the concentration of phenol first went up from 0.039 g/mL at the control to 0.087 g/mL
and subsequently declined to 0.048 g/mL as WHSV-1 increased from 0.045 to 0.126 h. The
variations of the chemical composition in bio-oils affected by the alteration of WHSV-1 were
much alike other researches’ results.5,
33
These outcomes suggest that a higher WHSV-1,
which means more active sites were offered to volatile vapors in unit time, resulted in more
51
severe zeolite cracking, oligomerization, aromatization reactions occurs in the pack-bed
reactor.
2.4.4 Analysis of non-condensable gases
Gas was one of the major co-products of cellulose from packed-bed catalysis integrated with
microwave pyrolysis. The composition of the gases were detected and quantified by
Micro-GC, which would be helpful to investigate the reaction pathways in the packed bed
reactor. The gas yield varied from 46.01 to 53.06 wt.% in comparison with no catalyst loaded
in the packed bed rector observing 46.89 wt.% gas yield as shown in Table 2.1. The large
amount gas should be attributed to the modified catalyst in the process that can facilitate the
cracking reaction of volatiles toward small molecules from decomposition of cellulose. The
non-condensed gas was mainly comprised of hydrogen, carbon dioxide, carbon monoxide,
methane, and light olefins such as ethylene.
The composition of gaseous fraction with respect to catalytic temperature is depicted in Fig.
2.3. Results indicates that hydrogen was the dominant composition from 41.83 to 32.89
vol.%. It is attributed to the technology of microwave-assisted pyrolysis inducing hot spots
where are capable of producing hydrogen-rich fuel gas. Unlike the conventional pattern of
pyrolysis with low heating rate, microwave-mediated pyrolysis with a focused and high
heating rate can readily induce hot spots inside biomass.17 This high amount of hydrogen can
be supplied for the hydrogenation of aromatics or fuel gas. The concentrations of carbon
dioxide, carbon monoxide, and methane gradually went up as the catalytic temperature
52
increased, which verifies that decarbonylation, decarboxylation, and oligomerization
reactions on the catalyst were promoted by increasing catalytic temperature. Owing to the
reducing characteristics of syngas (H2, CO), they could react with oxygen content in
pyrolysis volatiles to generate water and CO2. More aromatic hydrocarbons obtained at high
catalytic temperature as reported previously; hence oxygenated compounds derived from
decomposition of cellulose were reduced by the reducing syngas. As the more reducing
syngas was consumed, consequently the less reducing syngas was gained, which was
consistent with GC results as shown in Fig. 2.3. CO derived from furans decarbonylation
reaction offset the consumption for deoxygenation reaction. Wherefore the content of CO
was kept at the higher level from catalytic pryolysis than parent pyrolysis. High amount of
CO2 not only arose from decarboxylation reaction but also was produced from CO reducing
conversion, which was correspondingly increased as catalytic temperatue was elevated. As
expected, the maximum selectivity of xylenes (12.39%) and toluene (8.77%) in the upgraded
oils were received at catalytic temperature of 500 °C, which suggests that the elevated
temperature favored the formation of xylenes and toluene. High selectivity of p-xylene
(12.39%) at 500 °C arises from the removal of strong acid sites in catalysts after
hydrothermal treatment suppressing disproportionation, transalkylation, which is in
agreement with the former observation.34
53
Fig. 2.3 The composition of gaseous fraction with respect of catalytic temperature at the same
WHSV-1 (0.067 h).
According to the detailed production distributions, The proposed reaction pathway for
furfural conversion to xylens and toluene is described in Fig. 2.4. It was proposed that
furfural initially underwent decarbonylation reaction to form allene (C3H4) and CO.35 Since
there were no any furfural compounds being detected in the catalytic experiments, they could
act as the intemidates to be completetly transformed. The allene could go through either
aromatization to generate benzene or oligomerization and crackings to form propylene,
hydrogen and cyclopentadiene.35 The alkylation of benzene readily occurred under the high
temperature condition due to the active feature of benzene. Thus propylene that served as the
intermediate compound could react with benzene through alkylation to produce xylenes,
toluene and ethylene. Meanwhile, toluene formed could further react with propylene to
54
generate xylenes via alkylation. From the reaction pathway of furfural conversion into
xylenes, proportion of ethylene as by-products could be viewed as an index of aromatic
formation for catalytic pyrolysis of cellulose. Increasing proportion (2.23 – 7.79 vol.%) of
ethylene as a function of catalytic temperature was in accordance with the rising xylenes
yield in Fig. 2.3. Accordingly, the product distribution not only confirms that aromatic
hydrocarbon originated from furans derived from cellulose, but also reaffirms the
convergence of intermediate products to aromatic hydrocarbons.
H
O
O
Al
-O
+
Si
O
O
H
O+ O
Al- Si
O
H
O+ O
Al- Si
O
Decarbonylation
CO
C
+
Aromatization
Alkylation
H H2
+
O
O
-O
Al Si
+
O
Oligomerization
Crackings
H2 +
H
+
O
O
-O
Al Si
H
O+ O
Al- Si
+
Fig. 2.4 The proposed reaction pathway for furfural conversion to xylenes and toluene over modified
ZSM-5.
The development of the gaseous composition as a function of WHSV-1 is illustrated in
Fig. 2.5. The yield of hydrogen was first increased slightly as the catalyst was
employed, albeit in a less pronounced way; it was thereafter decreased due to the
hydrogen-consumed reactions dominating the catalytic process. The yield of carbon
dioxide was found to increase in the most pronounced way from 35.48 vol.% at the
control to 41.15 vol.% at WHSV-1 of 0.126 h, which suggests that more catalyst added
in the process accelerated the decarboxylation reaction in the process. With the
55
increasing addition of the catalyst, the amounts of gases with such as ethylene and
methane generally increased. These gaseous results imply that in addition to aromatics,
catalytic microwave pyrolysis of cellulosic biomass can be tuned to generate
hydrogen-enriched fuel gas, which can be used either directly as energy or fuel
precursors after upgrading.
Fig. 2.5 The composition of gaseous fraction as a function of WHSV-1 at the same catalytic
temperature (375 °C).
2.5 Conclusions
This study investigated that renewable gasoline-range aromatics and hydrogen-riche
fuel gas can be produced from catalytic microwave pyrolysis of cellulose over
modified ZSM-5. GC-MS analysis showed aromatic hydrocarbons were the most
abundant compounds which were approximately 82.93 – 96.60% in upgraded bio-oils,
56
depending on the catalytic pyrolysis conditions. Over 20 wt.% carbon yield of
aromatics were obtained under various catalytic conditions. Decarboxylation and
dehydration were the dominant deoxygenated reactions in the catalytic process. The
non-condensable gases (46.01 - 53.06 wt.% based on cellulose mass) were mainly
composed of hydrogen, carbon dioxide, carbon monoxide, methane. The share of
hydrogen was from 41.83 to 32.89 vol.%, which was the dominant composition in the
gases fraction. The coke deposition yield varied from 0.74 to 3.27 wt.%; spent catalyst
also showed decent catalytic performance due to the quite high BET surface area, pore
volume, and pore surface area. Significant quality of bio-oils and flexibility of
catalytic microwave mild conditions provide competitive merits compared with
conventional methods from petroleum resources. From this perspective it was proved
that cellulosic biomass can be used to generate developed products (aromatics and
hydrogen-rich gas) through catalytic microwave-induced pyrolysis.
57
2.6 References
1.
P. R. Patwardhan, J. A. Satrio, R. C. Brown and B. H. Shanks, Journal of Analytical
and Applied Pyrolysis, 2009, 86, 323-330.
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J. R. Regalbuto, Science, 2009, 325, 822-824.
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T. R. Carlson, J. Jae, Y.-C. Lin, G. A. Tompsett and G. W. Huber, Journal of Catalysis,
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H. Zhang, Y.-T. Cheng, T. P. Vispute, R. Xiao and G. W. Huber, Energy &
Environmental Science, 2011, 4, 2297-2307.
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Y.-T. Cheng and G. W. Huber, Green Chemistry, 2012, 14, 3114-3125.
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V. K. Venkatakrishnan, W. N. Delgass, F. H. Ribeiro and R. Agrawal, Green Chem.,
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A. Zheng, Z. Zhao, S. Chang, Z. Huang, K. Zhao, H. Wu, X. Wang, F. He and H. Li,
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P. Lidstrom, J. Tierney, B. Wathey and J. Westman, Tetrahedron, 2001, 57,
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Q. Bu, H. Lei, L. Wang, Y. Wei, L. Zhu, X. Zhang, Y. Liu, G. Yadavalli and J. Tang,
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E. F. Iliopoulou, E. V. Antonakou, S. A. Karakoulia, I. A. Vasalos, A. A. Lappas and K.
S. Triantafyllidis, Chemical Engineering Journal, 2007, 134, 51-57.
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J. C. Groen, J. A. Moulijn and J. Pérez-Ramírez, Microporous Mesoporous Mater.,
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J. Weitkamp, Solid State Ionics, 2000, 131, 175-188.
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M. Görling, M. Larsson and P. Alvfors, Applied Energy, 2013, 112, 440-447.
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technology, 2013, 135, 659-664.
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T. Namioka, A. Saito, Y. Inoue, Y. Park, T.-j. Min, S.-a. Roh and K. Yoshikawa,
Applied Energy, 2011, 88, 2019-2026.
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M. A. d. Hollander, M. Wissink, M. Makkee and J. A. Moulijn, Applied Catalysis A:
General, 2002, 85-102.
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J. Cai, W. Wu, R. Liu and G. W. Huber, Green Chemistry, 2013, 15, 1331.
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L. Wang, H. Lei, J. Lee, S. Chen, J. Tang and B. Ahring, RSC Adv., 2013, 3, 14609.
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X. Zhang, H. Lei, G. Yadavalli, L. Zhu, Y. Wei and Y. Liu, Fuel, 2015, 144, 33-42.
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A. Marcilla, A. Gómez-Siurana and F. Valdés, Journal of Analytical and Applied
Pyrolysis, 2007, 79, 433-442.
26.
Y. T. Cheng, Z. Wang, C. J. Gilbert, W. Fan and G. W. Huber, Angewandte Chemie,
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K. Wang, K. H. Kim and R. C. Brown, Green Chemistry, 2014, 16, 727.
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A. Corma, G. Huber, L. Sauvanaud and P. Oconnor, Journal of Catalysis, 2007, 247,
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Q. Bu, H. Lei, S. Ren, L. Wang, J. Holladay, Q. Zhang, J. Tang and R. Ruan,
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59
CHAPTER THREE
PRODUCTION OF GASOLINE-RANGE HYDROCARBONS FROM
MICROWAVE-INDUCED PYROLYSIS OF LOW-DENSITY
POLYETHYLENE OVER ZSM-5
3.1 Abstract
The microwave-induced pyrolysis of low-density polyethylene (a model of waste
plastics) towards its conversion into biofuels was investigated using ZSM-5 as a
catalyst, generating significant amounts of gasoline-range hydrocarbons. A central
composite experimental design (CCD) was done to investigate the effects of catalytic
temperature and reactant to catalyst ratio on the pyrolysis-oils composition and to
achieve the maximum liquid yield. The optimized condition for maximizing the yield
of upgraded oil (32.58 wt. %) was at 450°C and reactant to catalyst ratio of 2. GC-MS
analysis showed that mono-ring aromatic hydrocarbons were enriched and became the
most abundant compounds which varied from 74.73% to 88.49% in upgraded
pyrolysis-oils, depending on the catalytic pyrolysis conditions. Both low temperature
and high reactant to catalyst ratio gave rise to the formation of less desirable
polycyclic aromatic hydrocarbons whereas high temperature and high ratio contributed
to mono-ring aromatic hydrocarbons. The primary reaction competing with aromatic
hydrocarbon production was the formation of coke which was negligible even at low
catalytic temperatures. A plausible reaction mechanism was also proposed in order to
60
shed light on the overall catalytic microwave pyrolysis of LDPE for aromatic
hydrocarbons.
Keywords: catalytic microwave pyrolysis; LDPE; gasoline-range hydrocarbons; ZSM-5
catalyst; reaction mechanism.
3.2 Introduction
Plastics are used extensively in daily life as well as in industries acting as an indispensable
ingredient due to the versatility and low cost; thus consumption of virgin plastics has
increased exponentially over the past decades
1, 2.
Current per capita consumption of
exploited polyolefins for any technical and packing usage was quantified at 100 kg in 2010
and predicted to rise to 140 kg in the forthcoming years 3. The primarily used plastics were
polyethylene (PE) and polypropylene with the amount of 13.6 × 109 kg and 8.9 × 109 kg
respectively, accounting for 29% and 19% of total demand of plastics in 20114. Consequently
the plastic wastes (approximately 26 ×109 kg) resulting from this expanding use of plastics
was enormous, yet a large amount of plastic wastes (around 10 ×109 kg) was not recycled 4, 5.
The widespread application of plastics caused huge economic and environmental concerns
about their resourceful disposal, in spite of significant measures that have been taken to
improve the valorization of plastics and recover plastic wastes (only 8% recovery of total
plastics in U.S.) 1, 6.
Landfills and incineration for disposal of waste plastics are commonly utilized among the
61
conventional existing methods
2, 7.
Most waste plastics generated were generally disposed in
landfills, which caused a serious danger towards the environment owing to plastics
degradation and subsequent contaminant generation 8, 9. An even worse perspective for waste
plastics is low density and low biodegradability leading to a tremendous and long-term filling
of landfill sites
4, 10.
Energy recovery by means of energetic valorization (incineration) is
prevalently applied among thermal treatment on the base of the high calorific value of waste
plastics 8. On the other hand, incineration stimulates the release of harmful compounds such
as acid gases, dioxins, and furans into the atmosphere together with heavy metals causing
damage towards the environment and human health
2, 11.
Accordingly there is an urgent
demand for addressing waste plastics in a different way to the conventional treatment systems:
incineration and landfill.
The state of the art on tertiary recycling technologies and implementation is not only more
feasible at different scales but also more economically viable and environmentally friendly
than that of typical technologies 10. Pyrolysis, as an effectual and versatile route for valorizing
polyolefins (2/3 of waste plastics), has been highly developed with respect to environmental
management among the tertiary recycling of waste plastics
10, 12.
Thermal degradation of
waste plastics occurs through a complex free-radical mechanism involving the scission of
long polymeric molecules by exposure to elevated temperatures in an inert gas environment 1,
13.
Broadly speaking, products evolved from thermal decomposition of waste plastics give
rise to a heterogeneous hydrocarbon mixture of paraffins and olefins over a wide range of
molecular weights
9.
Volatile and condensable hydrocarbons with their relevant
62
concentrations varying with the dynamic conditions employed are predominant products
obtained, intensifying the recovery of valuable hydrocarbons and providing the gasoline
precursor for petrochemical industries
12, 14.
In contrast, due to the poor heat conductibility
and thermal decomposition of plastics, the process requires appreciable amounts of energy
and elevated temperature 8; another detriment exhibits that a reactively broad spectrum of
products are generated from the thermal degradation of macromolecules to small molecules
complicating their utilization on an industrial scale at present 8. Therefore, pyrolysis oil has to
undergo downstream catalytic upgrading to overcome the aforementioned problem prior to
being used as a conventional transportation fuel.
Towards this end, catalytic pyrolysis represents an essential step forward in the enhancement
of transforming waste plastics into gasoline-range hydrocarbons, eliminating the necessity for
further processing
15.
The suitable catalysts introduced confers the thermal degradation an
additional value to a gradual extent since an adequate catalyst can narrow the spectrum of
evolved products towards an excellent selectivity contributing to more valuable products
even at low temperatures 7. The catalytic cracking furthermore demands less energy as the
catalysts used either in situ or online reforming of pyrolysis volatiles can reduce the
activation energy, lowering the required temperature and optimizing the quality of the product
stream
16.
Zeolites have been verified to be noticeable in generating evolved hydrocarbons
lumped in gasoline-range derived from the pyrolysis of waste plastics containing
silica-alumina, β-zeolite, Y-zeolite, mordenite, HZSM-5, MCM-41, clinoptilolite
8, 14, 17, 18.
Among zeolite-based catalysts, ZSM performs as a feasible catalyst displaying a sharp
63
selectivity in the formation of branched hydrocarbons fostered by the isomerization and
aromatization reactions as a function of the structure of zeolitic framework, including acid
strength, external surface area and microporous texture
18.
Moreover carbonaceous coke
deposited on ZSM-5 deactivates the catalyst to a lesser extent, which is mainly attributed to
the tridimensional porous structure of ZSM-5 giving way to the circulation of aromatic coke
precursors toward the outside of the zeolite micropores 19.
As a matter of fact, catalytic pyrolysis of polyolefins has been studied by a multitude of
researchers in the presence of HZSM-5 striving to obtain hydrocarbons remarkably composed
of non-condensable olefins along with low concentration of aromatic hydrocarbons in the
liquid fraction
3, 8, 10, 12, 14, 15, 19-21.
Light olefins or aromatic hydrocarbons with low content
derived from catalytic cracking of polyolefins in the presence of HZSM-5 have been
investigated in fluidized bed reactors
8, 15, 20
and conical spouted bed reactors
3, 10, 12, 19,
respectively. Given in situ pyrolysis combining thermal cracking and catalytic cracking in a
single reactor, catalysts placed in the reactors are in intimate contact with waste plastics;
hence char formed as well as the liberated impurities readily deactivate the catalysts in the
integrated reactions 9. In addition, variables associated with catalysts are reactively difficult
to simulate for investigating the influences of catalysts parameters alterations on products
distribution. Nevertheless sequential pyrolysis and catalytic reforming are prone to noticeably
mitigate the problems
3, 9, 22.
A downstream fixed-bed reactor introduced in the study
5
was
conducive to an upward tendency in contents of aromatic hydrocarbons and light olefins
mostly dependent upon the fixed bed conditions.
64
The larger scale implementation of the process was restricted by the sticky nature and low
thermal conductivity of fused polyolefins, thereafter hindering heat and mass transfer
between phases
19.
Hence an adequate reactor
is crucial
to overcome characteristic
limitations of plastics in order to reinforce heat and mass transfer between phases reducing
plastics decomposition to carbonaceous materials 3. Microwave-assisted pyrolysis technology
is one of the most ideal methods of enhancing and accelerating chemical reactions due to
effective heat transfer profiles through microwave irradiation
23.
The Microwave assisted
reactor herein is transparent for effectively decomposing waste plastics in accordance with
satisfactory performances. In comparison with traditional pyrolysis, the microwave-induced
pyrolysis which encloses the potentials of fast and selective heating, easy control of reaction
conditions, low reaction temperatures and energy requirements could somewhat overcome
these limitations
24, 25.
In detail, the microwave-induced pyrolysis approach has been
successfully investigated from plant residues
26, 27;
yet little research is demonstrated for
microwave-assisted pyrolysis of polymeric materials 4, 28-30. Among previous studies, thermal
degradation of polyethylene was investigated28; Other researchers were devoted to improving
the quality of products on the base of an activated carbon bed coupled with microwave
pyrolysis of polyethylene 29. In addition to pure polyolefins, waste polyolefins were disposed
to efficiently convert into valuable chemicals by means of microwave-induced pyrolysis4.
Aforementioned studies were chiefly engaged in the topics of gaseous hydrocarbons evolved
from in situ or ex situ catalytic pyrolysis of polymeric materials; nonetheless there is no
65
literature on characterizing, in depth, the evolution of individual aromatic hydrocarbons in
liquid phase derived from microwave assisted pyrolysis of polyolefins over ZSM-5. As low
density polyethylene (LDPE) with branched carbon chains indicates an intermediate behavior
between high density polyethylene (HDPE) and polypropylene 4; the objective of this study
thereby aims to fill this knowledge gap in catalytic microwave pyrolysis of polyolefins over
ZSM-5 and take advantage of low density polyethylene (LDPE) as a model compound to
predict gasoline-range hydrocarbons. In this study, we will also elucidate the overall reaction
mechanism for the conversion of LDPE to aromatics by microwave-assisted pyrolysis in the
presence of ZSM-5. The effects of catalytic temperature and catalyst to feedstock ratio on
product yields were investigated and determined the thermochemical maximum of
hydrocarbon yields evolved from LDPE catalytic pyrolysis was determined. The
compositions of resulting bio-oil from ZSM-5 assisted cracking in packed-bed catalysis were
characterized by GC-MS; the non-condensable gas contents were analyzed by Micro-GC.
3.2 Materials and Methods
3.2.1 Materials
Low density polyethylene (LDPE) (CAS number 9002-88-4) in the form of pellets is
purchased from Sigma-Aldrich Corporation (St. Louis, MO, USA). The density and melting
point of LDPE are 0.925g/cm3 at ambient temperature and 116 °C, respectively. Prior to the
reactions, ZSM-5 (Zeolyst International, USA; SiO2/Al2O3 Mole Ratio: 50) is oven dried at
105 °C for 12 h and calcined at 550°C for 5 h in a muffle furnace.
66
Fig. 3.1 The schematic diagram of the microwave-assisted pyrolysis system integrated with
zeolite catalysis process.
3.2.2 Catalytic microwave degradation of low density polyethylene (LDPE)
Fig. 3.1 shows the schematic diagram of the microwave-assisted pyrolysis system coupled
with zeolite reforming process. The microwave pyrolysis system is mainly comprised of
several components: a 1000W, 2.45GHz microwave cavity, an infrared temperature sensor for
temperature measurement, a 500mL quartz flask inside the microwave oven in which the
reactant (LDPE) is loaded, and a product cooling and collection system where the
condensable liquid is collected. The temperature of cooling water in the condensers is around
2 °C. LDPE is placed in the 500 mL quartz flask inside the Sineo MAS-II batch microwave
oven (Shanghai, China), a constant microwave power setting (700 W) is carried out. The
system is purged with nitrogen on a flow rate of 1000 mL/min for 15 min prior to pyrolysis
67
reaction to maintain an oxygen-free environment. All reactions of microwave pyrolysis are
conducted at the temperature of 480 °C for 10 min.
The pyrolysis volatile vapors from the flask pass through a packed bed catalysis reactor
which is filled with catalyst. The packed-bed reactor customized is constructed of quartz and
externally heated by a heating tape. A thermocouple is introduced between the reactor and the
heating tape to measure catalytic temperature. The separate heating regimes ensure that
distinctly separate temperatures in the microwave oven and packed-bed reactor are
maintained. Then after the condensation system, the condensable liquid is collected as
pyrolysis-oil. The non-condensable vapors escape as gas at the end of the condensers and are
collected for analysis. The char is left in the quartz flask. The coke is formed on the catalyst,
which is determined by the difference in mass before and after catalytic pyrolysis. The weight
of gas is calculated using the following equation:
ℎ
=
!"#
−
− ℎ
−
(1)
The liquid, gas, coke, and coke yields are calculated by their corresponding masses divided
by the initial reactant mass.
3.2.3 Experimental design and optimization
A central composite experimental design (CCD) is employed to optimize the process
conditions and product yields distribution. The catalytic temperature (X , °C) and reactant to
catalyst ratio (X ), are chosen as independent variables. A central composite experimental
68
design (CCD) is used as given in Table 1. Y% is recorded as the dependent output variables.
The loading of LDPE feed is 20g, and catalyst loading is varied from 4.27 to 15.15g, which
corresponds to reactant to catalyst ratio (1.32 – 4.68). The packed-bed temperature is varied
from 249 to 501°C.
For statistical calculations, the variables X% are coded as x% according to Eq. (2):
x% = (X% − X' /∆X)
(2)
where x% is the coded value of an independent variable, X% is a real value, X' is a raw
value of the variable at the center point, and X ' /∆X is the step change. A total amount of 20
experiments which are performed to optimize the reaction conditions including 7 replications
at the center points (n0 = 7) from a 22-factorial CCD design. The second order polynomials
(Eq. (4)) are calculated with a SAS statistical package (SAS Institute Inc., USA) to evaluate
the response of the dependent variables.
Y% = b' + b X + b X + b X + b X X + b X
(3)
where Yi is predicted response, X1, X2 are independent variables. b0 is the offset term; b1, b2
are linear effects; b11, b22 are squared effects and b21 are interaction terms, which are all
regression coefficients.
3.2.4 Analytic methods
The textural properties are determined by means of N2 adsorption–desorption (Micromeritics
TriStar II 3020 Automatic Physisorption Analyzer). The acidity of the catalyst is measured by
69
temperature-programmed desorption (TPD) of ammonia with a Micromeritics AutoChem II
2920 Chemisorption Analyzer with Vapor Generator.
The chemical composition of the bio-oils is characterized by Agilent 7890A GC-MS
(GC–MS; GC, Agilent 7890A; MS, Agilent 5975C) with a DB-5 capillary column. The GC is
first programmed to heat to 45°C for 3 min followed by heating to 300°C at a rate of
10°C/min. The injection sample size is 1 μL. The flow rate of the carrier gas (helium) is
0.6mL/min. The ion source temperature is 230°C for the mass selective detector. Compounds
are identified by comparing the spectral data with that in the NIST Mass Spectral library. The
area percent of compounds obtained from GC/MS results is utilized to predict products
selectivity.
The gaseous product is collected in a 1L Tedlar gas bag and then offline analyzed by an
INFICON 3000 Micro-GC (INFICON Inc., Santa Clara, CA, USA) system with a thermal
conductivity detector (TCD).
70
Table 3.1 Experimental design and product yield distribution.
Runa
Catalytic
Reactant to
Liquid
Gas
Char
Coke
Temperature (ºC)
Catalyst Ratio
(wt. %)
(wt. %)
(wt. %)
(wt. %)
P-1
300
2
27.44
67.74
1.84
2.98
P-2
300
4
26.41
69.56
1.82
2.21
P-3
375
1.32
26.57
70.57
1.88
0.98
P-4
375
3
24.44
73.2
1.84
0.52
P-5
248.87
3
27.38
67.28
1.85
3.49
P-6
375
3
24.56
72.96
1.93
0.54
P-7
375
3
24.44
73.02
1.89
0.65
P-8
375
3
24.74
72.81
1.89
0.55
P-9
375
3
24.94
72.73
1.78
0.55
P-10
450
2
32.08
65.77
1.81
0.34
P-11
375
3
24.2
73.5
1.79
0.5
P-12
375
4.68
22.16
75.59
1.86
0.4
P-13
501.13
3
29.7
68.22
1.83
0.25
P-14
300
4
26.48
69.74
1.79
1.99
P-15
375
3
24.92
72.54
1.94
0.6
P-16
450
4
27.14
70.84
1.82
0.2
P-17
450
4
27.39
70.63
1.88
0.1
P-18
450
2
32.58
65.13
1.89
0.4
P-19
300
2
27.21
67.74
1.8
3.25
P-20
375
3
25.04
72.53
1.85
0.58
P-21b
-
-
-
35.38
1.84
-
a
P-1 to P-20 was conducted with catalyst addition; P-21 was the control.
b The
rest yield (62.78 wt.%) determined was from waxes of C21+
71
3.4 Results and discussion
3.4.1 Response surface analysis
It is noteworthy that physical factors such as various packed-bed temperature and reactant to
catalyst ratio had a significant influence on products yield distribution during microwave
pyrolysis of Douglas fir pellets using ZSM-5 as a catalyst in aforementioned research.27, 31
Thus, catalytic temperature and reactant to catalyst ratio were herein chosen as independent
variables to investigate the bio-oil distribution from catalytic microwave induced pyrolysis of
LDPE in the present study.
The detailed experimental design and products distribution obtained were shown in Table 3.1
depending on reaction conditions. Linear model equations for pyrolysis-oil (Eq. 4) and gas
(Eq. 5) as a function of catalytic temperature (X1, °C) and reactant to catalyst ratio (X2)
received were presenting the results of the experiments:
Y./012/3%341%2 = 25.00 + 1.14X − 1.41X − 1.05X X + 1.83X + 0.35X
(4)
Y:;3 = 72.52 − 0.061X + 1.67X + 0.84X X − 2.35X − 0.46X
(5)
The model for yield prediction was reduced by using backward statistical analysis, and
parameters were sequentially removed based on the coefficient’s P-value until all remaining
were significant (P-value<0.05). For ANOVA analysis, the P-value of Eq. (4) was 0.0002,
which was less than 0.05, evidencing that Eq. (4) was significant to describe the relationship
between pyrolysis-oil yield and reaction conditions. The coefficient of determination (R2) for
72
Eq. (4) was 0.80, implying that the second order regression model finely represented the
relationships among the independent variables. The P-values for model term X1 , X2, X1X2,
X12 were significant since P-value for these models were 0.0073, 0.0016, 0.0446, 0.0001
respectively, which were less than α=0.05; whilst the P-value related to X22 was 0.3305
suggesting that the model term for X22 was not significant. Accordingly catalytic temperature
and reactant to catalyst ratio had a tremendous influence on pyrolysis-oil yield as P-values of
Eq. (4) was less than α=0.05; meanwhile it is noted that the reduced regression model
precisely illustrated the reaction as a result of the coefficient of determination (0.80).
P-value of Eq. (5) was 0.0002 < α=0.05, it therefore enabled to fit the data for gas yield. The
coefficient of determination (R2) for Eq. (5) was 0.80, indicating that the model can fairly
predict the gas yield among the independent variables. The P-values for model term X2, X12
were 0.0010, <0.0001, respectively; whereas model term X1, X1X2, X22 were larger than
α=0.05. Hence, Eq. (10) was markedly presented the intimate relationship between gas yield
and reaction variables according to the P-value. It is moreover noticeable that the model
could really predict the combination of factors to achieve the best response on the base of the
coefficient of determination. In contrast, char yield always fluctuated at 1.85 wt. % as shown
in Table 3.1, which could not be accurately used to predict neither the maximum yield nor
minimal yield.
The product (pyrolysis-oil and gas) yields were visualized in accordance to the response
surface and contour line in Table 3.1 and Fig. 3.2 (A), (B). Generally speaking, the yields of
73
pyrolysis oil and gas were in the range from 22.16 wt. % to 32.58 wt. % versus 65.13 wt. %
to 75.59 wt. % respectively, showing inverse trends owing to upgrading conditions. The
optimal condition for maximum pyrolysis-oil yield was found to be a catalytic temperature of
450 ºC and reactant to catalyst ratio of 2; the maximum gas yield was predicted at 375 °C and
reactant to catalyst ratio of 4.68. As previously mentioned, the negligible char yield remained
virtually constant at 1.85 wt. %, suggesting that LDPE was almost decomposed to volatile
compounds. The strong acid sites existed in ZSM-5, affording excellent catalyzed
performances, further favoring the formation of small molecule hydrocarbons. Similar
phenomenon was reported elsewhere3, 5 as a function of catalytic temperature and catalyst
loading.
(A)
74
(B)
Fig. 3.2 Effect of the interaction of the independent variables on pyrolysis-oil (A) and gas (B)
(red pots represent actual experimental values).
Fig. 3.3 The chemical composition of the pyro-oils from GC/MS analysis with ZSM (the
serial numbers correspond to those in Table 3.
75
3.4.2 The analysis of the pyrolysis-oils by GC/MS
3.4.2.1 Chemical composition of pyrolysis-oils
In order to further understand the chemical reaction of catalytic microwave pyrolysis of
LDPE and obtain more insight into evolved pyrolysis-oils, the pyrolysis-oils were
characterized and categorized by GC/MS on the basis of the various functional groups; the
chemical compounds of pyrolysis-oils are depicted in Fig. 3.3. Generally the typical
compositions of pyrolysis-oils derived from polyolefins pyrolysis have been categorized into
the lumps of paraffins, olefins, iso-paraffins, aromatics, naphthenes, waxes (C12+), and some
unclassified compositions. However it was observed that aromatic hydrocarbons were found
to be the dominant chemical compounds of the upgraded pyrolysis-oils which arose from
catalytic microwave pyrolysis of LDPE. Owing to the low catalytic temperature, there was a
minor concentration (approximately 6%) of aliphatic hydrocarbons (C11+) in Run P-5; while
neither aliphatic hydrocarbons nor waxes were detected in other compared experiments.
Therefore it is transparent that high temperature favors the cracking and aromatizing
reactions over ZSM-5. Mono-ring aromatic hydrocarbons (C8 - C12) accounted for a large
amount of pyrolysis-oils from 74.73% to 88.49% depending on alterations of reaction
conditions. Previous research3,
5, 14, 15, 19
associated with the yield of mono-ring aromatic
hydrocarbons implied that the highest yield of aromatics in the liquid fraction rarely reached
the ceiling of 30% from catalytic fast pyrolysis of plastics. It is also observed that liquid
fractions also contained other less desired products in aforementioned studies, resulting in
reducing the value of liquid products application. The maximum amount of single-ring
76
aromatic hydrocarbons (88.49%) was attained at packed-bed temperature of 450°C and
reactant to catalyst ratio of 4 as described in the experiment of Run P-17. Mono-ring aromatic
hydrocarbons
obtained
were
predominantly
comprised
of
xylenes,
ethyltoluene,
trimethylbenzene, indane, diethylbenzene, and their derivatives.
In addition, double-ring aromatic hydrocarbon in the evolved pyrolysis-oils varied from 11.51%
to 25.27% indicating an inverse tendency with mono-ring aromatic hydrocarbons.
Double-ring aromatic hydrocarbons as the following content mainly consisted of naphthalene,
methylnaphthalene, dimethylnaphthalene, trimethylnaphthalene, and their derivatives. Even
though naphthalene and its derivatives are usually regarded as undesired products as it is less
value than the monocycle aromatics, nevertheless they can be currently treated as the
precursors of jet or rocket fuels in the aviation field.32 Overall it is affirmed that high yield of
observed aromatic hydrocarbons derived from catalytic pyrolysis of LDPE were attributed to
short chains with a high branched degree for LDPE, which
led to readily cracking
small
molecule olefins; consequently the behavior facilitated a progressive formation of branching
aromatics evolved from small olefins condensation and aromatization.19 It is also observed
that product oils contained a gasoline fraction with the presence of xylenes, ethyltoluene,
trimethylbenzene, indane etc.
Because volatility and octane number of the hydrocarbons are regarded as indicators to
measure gasoline qualities, suited volatility is preferred for smooth operation of petrol
engines and higher octane numbers are required in internal combustion engines. In terms of
77
volatility, neither light hydrocarbons with high volatilities nor heavy hydrocarbons obtaining
low volatilities are selected for engine fuel usage; for octane number, aromatics, isoalkanes
are more desirable than olefins and paraffins. As branching compounds have moderately high
octane numbers and adequate volatilities, branched aromatic hydrocarbon ranged from C8 to
C10 was the maximum amount in gasoline.18 Table 3.2 interprets the identified compounds in
gasoline26 versus selected experiments. The experimental results indicated a higher total
amount of C8 - C9 aromatic hydrocarbons than that of gasoline (56.58%), attaining more
favorable properties for further application.
Table 3.2 Identified hydrocarbon products for gasoline versus LDPE pyrolysis-oils (%).
Name
ID
Gasoline
P-5
P-10
P-12
P-17
Xylenes and derivatives
C8H10
30.48
37.02
49.57
51.29
56.86
Indane
C9H10
0.25
0.88
1.63
1.15
1.03
Trimethylbenzene and derivatives
C9H12
25.85
25.69
15.46
22.99
19.96
3.4.2.2 The effects of catalytic temperature on the chemical composition of pyrolysis-oils
The effect of catalytic temperatures on chemical compositions of upgraded pyrolysis-oils was
elucidated with a consistent reactant to catalyst ratio of 3 as shown in Table 3.3. Nevertheless,
pyrolysis volatiles evaporating from the pyrolysis process without the introduction of catalyst
were principally decomposed to waxes with the carbon number of over 21, which is similar
with other report.3 The aliphatic hydrocarbon (6.53%) were completely converted into
aromatic hydrocarbons, undergoing zeolite cracking and subsequent aromatization when the
temperature was increased from 249°C to 501°C; therefore, an increasing temperature
enhanced selectivity of aromatic hydrocarbons in upgraded pyrolysis-oils.14 It is noted that a
78
significant effect of temperature observed on the mono-ring aromatic hydrocarbons overall
yield, presenting an upward trend from 79.30% to 84.34% as a result of the enhancement of
hydrogen transfer reactions when the temperature was increased from 249°C to 375°C. Thus,
a temperature increase in the range below 375°C gave rise to the generation of mono-ring
aromatic hydrocarbons; while the total amount of the gaseous compounds: ethylene and
ethane derived from the cracking reactions at the catalytic temperature of 501°C were higher
than that at 375°C (Fig. 3.4), implying that a temperature increase in the range above 375 °C
facilitated cracking reactions. Nonetheless, the aromatic components were not easily cracked
under these conditions;19 the yield of mono-ring aromatic components experienced a slight
decrease to 83.34% at 501°C. By contrast, the content of double-ring aromatic hydrocarbons
displayed an increase (14.17% - 16.66%) along with the increased temperature, which was
slightly affected in the temperatures intervals studied.
Table 3.3 Products selectivity for aromatic hydrocarbons species on the base of
catalytic temperature by GC/MS area at the same reactant to catalyst ratio of 3.
Catalytic Temperature (°C)
249
375
501
Xylenes(C8H10)
37.02
52.30
52.75
Ethyltoluene(C9H12)
19.09
12.13
8.83
Trimethylbenzene(C9H12)
6.60
7.99
9.58
Indane(C9H10)
0.88
1.16
1.40
Diethylbenzene(C10H14) and isomers
8.42
5.74
4.33
Methylindan(C10H12) and isomers
2.51
2.37
3.56
Dimethylindan(C11H14) and isomers
0.43
1.91
2.44
Mono-ring aromatic hydrocarbons selectivity
79
Pentamethylbenzene(C11H16)
3.16
0.24
0.11
Trimethylindan(C12H16) and isomers
1.20
0.49
0.52
Mono-ring aromaticsoverall yields (area %)
79.30
84.34
83.34
Naphthalene(C10H8)
3.59
3.70
3.63
Methylnaphthalene(C11H10)
4.92
5.80
6.32
Dimethylnaphthalene(C12H12)
3.42
3.82
4.70
Ethylnaphthalene(C12H12)
-
0.93
0.72
Vinylnaphthalene(C12H10)
0.13
-
-
Trimethylnaphthalene(C13H14) and isomers
2.12
1.42
2.00
Double-ring aromatics overall yields (area %)
14.17
15.66
16.66
Double-ring aromatic hydrocarbons selectivity
Catalytic temperature also had a significant influence on the species of aromatic
hydrocarbons. As mentioned above, the aromatic hydrocarbons were composed of mono-ring
hydrocarbons and double-ring hydrocarbons as depicted in Table 3.3. For mono-ring aromatic
hydrocarbons, xylenes accounted for over 50% at the catalytic temperature of 375°C and
501°C, which may be due to the size selectivity of the zeolite catalyst. Increasing the catalytic
temperature had a considerable effect on xylenes formation (from 37.02% to 52.75%) that
was attributed to the effect of the diffusivity of catalyst reduced at high temperature, thus
xylenes with high diffusivity can readily diffused out from ZSM-5 pores. The total amount of
single ring aromatic hydrocarbons (C9H12) including ethyltolune and trimethylbenzene
revealed an opposite trend, steadily declining from 25.69% to 18.41%. Hence higher
branched single-ring aromatic hydrocarbons were declined in terms of concentration as
catalytic temperature was raised; because alkylation of benzene which is an exothermic
reaction was suppressed at high temperature.33 Likewise, there were distinct downward trend
80
of diethylbenzene(C10H14), pentamethylbenzene(C11H16), trimethyllindan(C12H16) and their
isomers decreasing from 8.42% to 4.33%, 3.16% to 0.11% and 1.20% to 0.52%, respectively.
On the other hand, less branched single-ring aromatic hydrocarbons were slightly increased
involving methylindan (C10H12), dimethylindan (C11H14) and their derivatives as shown in
Table 3.3. As a matter of fact, the concentration of naphthalene was steadily fluctuated
around 3.64%, suffering a slight influence by catalytic temperature variations. There was a
further increase in the concentration of methylnaphthalene and dimethylnaphthalene with
increasing catalytic temperature, 4.92% - 6.32% and 3.42% - 4.70%, respectively. The results
obtained partially coincided with other outcomes.3, 5
3.4.2.3 The effects of reactant to catalyst ratio on the chemical composition of
pyrolysis-oils
The individual composition yield of pyrolysis-oils from catalytic microwave pyrolysis at
375°C as a function of reactant to catalyst ratio is shown in Table 3.4. Previous studies27, 31
indicates that ZSM-5 catalysts introduced expressed an excellent performance, generating a
remarkable amount of hydrocarbons derived from microwave pyrolysis of Douglas fir pellets.
Reactant to catalyst ratio herein was adjusted from 1.32 to 4.68 through changing the catalyst
mass (15.15g - 4.27g) whereas the LDPE mass were held constant in all experiments. The
control experiment without the addition of catalyst, the waxes of C21+ dominated the products.
The mono-ring aromatic hydrocarbons overall yield was increased gradually from 77.08 % to
87.16% with the increasing ratio from 1.32 to 4.68; as more catalyst was added to the packed
bed, more spaced catalytic sites of catalyst were available for oligomerization and
81
aromatization, and consequently mono-ring aromatic hydrocarbons evolved may enter the
spaced zeolite pores to further react with cracked olefins generating double-ring aromatic
hydrocarbons. In contrast, there was a corresponding decline in double-ring aromatic
hydrocarbons overall yields from 22.32% to 12.84% as the ratio was raised. Therefore, a
higher ratio favored the selectivity of mono-ring aromatic hydrocarbons, which is consistent
with the similar trend claimed by this research.15
The reactant to catalyst ratio also had a significant effect on the individual selectivity for
aromatic hydrocarbons. It was noteworthy that xylenes as the largest amount was first
increased from 43.90% to 52.30% (the ratio of 3) and stepwise descended to 51.29% with the
increasing ratio. This was presumably attributed to the high activity of ZSM-5 and intimate
contact with LDPE when more ZSM-5 was filled in the packed bed; xylenes could further
react with intermediate olefins to form corresponding double-ring aromatic hydrocarbons in
the presence of catalyst.34 The second fractions (ethyltoluene and trimethylbenzene) both
showed a slight increase to 22.99% from 19.40%. The content of methylindan was solely
indicated an opposite tendency, decreasing from 3.63% to 2.16%. Nevertheless there were no
dramatic effects on the remaining fractions of mono-ring aromatic hydrocarbons caused by
reactant to catalyst ratio, which were steadily fluctuant at the low amounts. As for
naphthalene and its derivatives, increasing the reactant to catalyst ratio resulted in slight
decline in formations of all double-ring aromatic hydrocarbons including naphthalene (4.25%
- 3.14%), methylnaphthalene (7.87% - 4.47%), dimethylnaphthalene (5.94% - 3.34%),
ethylnaphthalene (1.22% - 0.83%), trimethylnaphthalene and isomers (3.04% - 1.06%). These
82
outcomes suggest that a lower ratio, which means more active sites were offered on the
catalyst surface contributing to more severe zeolite reactions occurring in the pack-bed.
Table 3.4 Products selectivity for aromatic hydrocarbons species on the base of reactant to
catalyt ratio by GC/MS area at the same catalytic temperature (375°C).
Reactant to Catalyst Ratio
1.32
3
4.68
Xylenes(C8H10)
43.90
52.30
51.29
Ethyltoluene(C9H12)
11.95
12.13
14.67
Trimethylbenzene(C9H12)
7.45
7.99
8.32
Indane(C9H10)
1.21
1.16
1.15
Diethylbenzene(C10H14) and isomers
5.91
5.74
6.38
Methylindan(C10H12) and isomers
3.63
2.37
2.16
Dimethylindan(C11H14) and isomers
2.03
1.91
2.22
Pentamethylbenzene(C11H16)
0.08
0.24
0.16
Trimethylindan(C12H16) and isomers
0.92
0.49
0.81
Mono-ring aromatics overall yields (area %)
77.08
84.34
87.16
Naphthalene(C10H8)
4.25
3.70
3.14
Methylnaphthalene(C11H10)
7.87
5.80
4.47
Dimethylnaphthalene(C12H12)
5.94
3.82
3.34
Ethylnaphthalene(C12H12)
1.22
0.93
0.83
Trimethylnaphthalene(C13H14) and isomers
3.04
1.42
1.06
Double-ring aromatics overall yields (area %)
22.32
15.66
12.84
Mono-ring aromatic hydrocarbons selectivity
Double-ring aromatic hydrocarbons selectivity
83
Fig. 3.4 The composition of gaseous fraction with respect of catalytic temperature at the same
reactant to catalyst ratio of 3.
3.4.3 The analysis of gaseous fraction by Micro-GC
Gas was one of the major products of LDPE from packed-bed catalysis integrated with
microwave pyrolysis. The knowledge of the composition of the gas would be helpful to
investigate the reaction pathways in the packed bed reactor. The gas yield varied from 35.38
wt. % at the control to 75.59 wt. % at the temperature of 450 °C and the ratio of 2 as shown
in Table 3.1. The gaseous fraction was mainly comprised of ethylene, ethane, and small
quantities of hydrogen and methane. The composition of gaseous fraction with respect of
catalytic temperature is presented in Fig. 3.4. It was found that ethylene and ethane were the
dominating gaseous fractions (35.23 wt. % - 72.21 wt. % and 19.58 wt. % - 58.03 wt. %,
84
respectively). The amount of ethane was much higher than that of the control experiment
without catalyst added, while the yield of ethylene was generally higher than ethane in the
catalytic processes. The difference was ascribed to the shorter chains with a high branching
degree for LDPE, which caused the branched chains being cracked easily;19 consequently, the
thermal process alone facilitated a more rapid formation of ethane evolved from the branched
chains rather than ethylene from the dominant chains. It is stated that the amount of ethylene
was decreased as the catalytic was increased ranging from 375°C to 501°C. Although high
catalytic temperature favored the cracking of waxes, resulting in generating more ethylene;
more intermediate olefins in direct contact with the catalyst underwent a complete reforming
for the formation of aromatic hydrocarbons, offsetting this increasing tendency.1 It is worth
noting that the catalyst employed in the process enhances the production of hydrogen, which
is in agreement with the result reported elsewhere.7 That was attributed to the hydrogen
abstraction that occurs during the aromatization, oligomerization and hydrogen transfer
reactions, which was favored by ZSM-5 catalyst. The concentration of methane gradually
went up as the catalytic temperature was increased, which verifies that the cracking of short
branched chains could be promoted by the increasing catalytic temperature.
85
Fig. 3.5 The development of the gaseous composition as a function of reactant to catalyst
ratio at the same catalytic temperature (375°C).
The development of the gaseous composition as a function of reactant to catalyst ratio is
described in Fig. 3.5. The yield of ethylene was found to be increased in the most pronounced
way from 35.23 wt. % at the control to 82.17 wt. % at the ratio of 1.32, which suggests that
more catalyst added in the process accelerated the cracking of polyethylene. The yield of
hydrogen was first increased slightly as more catalyst was employed, albeit in a less
pronounced way than that of ethylene; it was thereafter decreased due to the cracking reaction
dominating the catalytic process. With the increasing addition of the catalyst, the amounts of
gases with thermodynamically stable molecule, such as ethane and methane, were generally
decreased.
86
3.4.4 The analysis of coke deposition on the catalyst
Given the coke deposited on the catalyst resulting in decrease of the active sites and
micropores blockage,21 the carbonaceous compounds were a crucial element to be taken into
consideration in the study of formation mechanism when zeolites are employed as catalysts.
Linear model equations for coke (Eq. 6) obtained are presented in the results of the
experiments as a function of catalytic temperature (X1, °C) and reactant to catalyst ratio (X2):
Y>1?@ = 0.61 − 1.09X − 0.25X + 0.20X X + 0.53X + 0.11X
(6)
The P-value of Eq. (6) was less than 0.0001, which was much smaller than α=0.05. Therefore,
the linear model could significantly predict the coke deposition on the catalyst. Moreover, the
coefficient of determination (R2) for Eq. (6) was 0.96, which indicates that the model
perfectly represents the relationships among the variables. The model term X1, X2, X1X2, X12
were significant apart from X22, since their P-values were <0.0001, 0.0016, 0.0339, <0.0001,
0.00885, respectively. The model of coke deposition yields as a function of catalytic
temperature and reactant to catalyst ratio is depicted in Fig. 3.6. According to Table 3.1, the
coke varied from 0.1 wt. % to 3.49 wt. %; it is noteworthy that deactivation is negligible
(below 1 wt. %) when operating at moderate conditions. Broadly speaking, the attractive
experimental results were primarily because of structures of ZSM-5 and the reaction
apparatus. Independent degradation of LDPE and subsequent online reforming were
conducive to reducing direct contact between waxes and catalyst 3. Furthermore, the ZSM-5
zeolite within the three-dimensional structure of interconnected channels had micropores; the
coke precursors formed were easily swept toward the outside of the crystals because of the
87
unique pore system.3,
12, 21
Consequently smaller coke retention was enhanced, which
mitigated the potential blockage of microporous entrance. As expected, the minor coke
deposition was also affected by short residence times set at 10mins, which abated the
formation of coke precursors.12 In detail, when catalytic temperature was fixed at 248.9 °C
and the ratio set at 3, the amount of coke accounted for 3.49 wt. %; the behavior observed
was attributed to incomplete conversion of waxes as main coke precursors.3 Formation and
deposition of coke was attenuated as the catalytic temperature was increased, declining to
below 1wt. % in major experiments. Besides the spent catalysts can reused or regenerated for
catalytic pyrolysis of waste plastic, which still have a decent performance claimed by the
study.12
Fig. 3.6 Effect of the interaction of the independent variables on coke yield (red pots
represent actual experimental values).
88
3.4.5 Mechanism analysis for aromatic hydrocarbons formation
It is observed that the aromatic hydrocarbon contents by catalytic pyrolysis were predominant
in the LDPE-derived oils. This phenomenon might be illuminated by olefinic intermediates
conversion and various endothermic reforming reactions occurring at high temperatures
which might be caused by the formation of hot spots because of microwave heating.25 Hence,
the degree of degradation during LDPE decomposition reaction was enhanced to some extent
via microwave heating.
A plausible reaction pathway for catalytic microwave pyrolysis of LDPE into aromatic
hydrocarbons over zeolite catalyst was presented in a two-step thermal and catalytic process
as shown in Fig. 3.7. Thermal degradation of polyethylene has been reported to take place
through two mechanisms when temperature was set below 500 °C: (1) random scission to
yield long-chain hydrocarbons; and (2) chain-end scission of the oligomers to yield
low-molecular weight products.3,
22
The two abovementioned mechanisms occurred
simultaneously, generating free radicals along the carbon chains; thereafter the cleavage of
the molecule caused the formation of a molecule with an unsaturated end and the other with a
terminal free radical.1 Consequently the radical fragments were transformed into straight
chain dienes, alkenes and alkanes through hydrogen transfer reactions.1 The volatiles
subsequently underwent catalytic cracking over ZSM-5 through two carbocationic
mechanisms: (1) the classical bimolecular mechanism (or β-scission); (2) the monomolecular
or proteolytic mechanism, through carbonium ions.3, 35 In the presence of LDPE, a large
amount of chain ends had access to the active sites inside the pores of ZSM-5 zeolite, in
89
which β-scission was the dominant reaction to form light olefins.14 When employing a
relatively high catalyst to reactant ratio and catalytic temperature, the volatiles were in direct
contact with the catalyst promoting the cracking and reforming reactions; otherwise the
catalytic degradation was not completed as shown in Run P-5, which resulted in forming high
concentration of intermediate olefins and sequentially yielding branching hydrocarbons.1
Lager intermediate olefins derived from the thermal degradation could not enter ZSM-5
micropores where the majority of the active acid sites were located. The reaction mechanism
for catalytic decomposition of macromolecules over zeolite is that they were degraded on the
active sites on the external surface of the zeolite crystallites; small fractions relatively
diffused into the zeolite micropores and further reacted on the internal sites.36 Therefore, the
porous structure of zeolite is significant for catalytic degradation of intermediate vapors.
R2
R2
n H 2C
CH
A
n H 2C
CH2
CH 2
CH
n H 2C
CH 2
R1
R1
HC
CH
CH
CH
CH2
R1 + R2
n
R
LDPE
n
HC
n H 2C
CH 2
CH
R
R
R
R
R
H2C
CH
CH
R
CH 2 +
H2
H2
Fig. 3.7 The overall reaction pathway for converting LDPE into aromatic hydrocarbons.
Moreover, the intermediate olefins as predominant products stepwise underwent
oligomerization, cyclation and hydrogen transfer reactions, contributing to forming aromatic
90
hydrocarbons; hence bimolecular reactions including alklation, cyclization and hydrogen
transfer reactions were enhanced in the micropores of ZSM-5, where steric hindrance of
bimolecular reaction caused the formation of aromatic hydrocarbons.14 Eventually upgraded
oils with high quality of aromatics were obtained when ZSM-5 zeolite was introduced. That
was attributed to the high number of Brönsted acid sites involved within this zeolite and
sharp selectivity of three dimensional microporous structure, promoting aromatization
reactions and lowering the probability of side reactions.1, 35
3.5 Conclusions
LDPE could be converted into valuable aromatics by catalytic microwave pyrolysis in
the presence of ZSM-5. The products distribution of the aromatic hydrocarbons was
correlated to catalytic temperature and reactant to catalyst ratio. A central composite
experimental design (CCD) was done to investigate the effects of catalytic temperature
and the ratio on the pyrolysis-oils composition and to achieve the maximum liquid
yield. GC-MS analysis showed gasoline-range hydrocarbons were enriched and
became the most abundant compounds which were approximately from 74.73% to
88.49% in upgraded pyrolysis-oils, depending on the catalytic pyrolysis conditions.
The optimized condition for maximizing amount of mono-ring aromatic hydrocarbons
was at 450°C and the ratio of 4. The syngas was mainly composed of light olefins,
paraffins, and hydrogen. Both high temperatures and high reactant to catalyst ratios
favoured the formation of mono-ring aromatics whilst high temperatures and low
ratios contributed to more undesirable polycyclic hydrocarbons. The primary
91
competing reaction with aromatic hydrocarbon production was the formation of coke
which was deposited on the surface of the catalyst leading to deactivate the acid sites,
accounting from 0.1 wt.% to 3.49 wt.%. Consequently a plausible reaction mechanism
was also proposed for detecting and understanding the overall catalytic microwave
pyrolysis of LDPE into aromatic hydrocarbons.
92
3.6 References
1.
D. P. Serrano, J. Aguado, J. M. Escola, J. M. Rodríguez and G. San Miguel, Journal
of Analytical and Applied Pyrolysis, 2005, 74, 370-378.
2.
M.-H. Cho, S.-H. Jung and J.-S. Kim, Energy & Fuels, 2010, 24, 1389-1395.
3.
M. Artetxe, G. Lopez, M. Amutio, G. Elordi, J. Bilbao and M. Olazar, Chemical
Engineering Journal, 2012, 207-208, 27-34.
4.
A. Undri, L. Rosi, M. Frediani and P. Frediani, Fuel, 2014, 116, 662-671.
5.
R. Bagri and P. T. Williams, Journal of Analytical and Applied Pyrolysis, 2002, 63,
29-41.
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95
CHAPTER FOUR
PRODUCTION OF RENEWABLE JET FUEL RANGE ALKANES AND
AROMATICS VIA INTEGRATED CATALYTIC PROCESSES OF
INTACT BIOMASS
4.1 Abstract
A novel pathway was investigated to produce jet fuel range paraffins (alkanes) and aromatics
by catalytic microwave-induced pyrolysis of intact biomass (Douglas fir) integrated with the
hydrotreating upgrading process. The proof-of-principle for the consecutive two-step process
for converting lignocellulosic biomass into jet fuel range paraffins and aromatics involves the
use of the well-promoted ZSM-5 in the process of catalytic microwave pyrolysis and Raney
nickel in the hydrotreating process.
The production of desired C8 – C15 aromatics was
achieved from catalytic depolymerization of intact biomass at 375 °C. Up to 12.63%
selectivity of C8 – C15 paraffins and 19.48% selectivity of hydro-aromatic hydrocarbons were
obtained from the hydrotreating of parent oil under a low-severity condition (for 2 h). A
central composite experimental design (CCD) was employed to investigate the effects of
reaction temperatures and initial pressures on the composition of hydrotreatd oils. We
observed that increasing reaction temperature and initial pressure with prolonged time could
enhance the hydrogenation and hydrogenolysis reactions to form jet fuel range paraffins and
aromatics. Gaseous fraction mainly consisted of unreacted hydrogen, carbon dioxide, and
96
methane.
Integrating catalytic processes of lignocellulosic biomass potentially paves a new
way for the development of jet fuels over inexpensive catalysts under the mild condition.
Keywords: Jet fuel range alkanes, aromatics, hydrotreating upgrading, catalytic microwave
pyrolysis; Raney nickel
4.2 Introduction
Growing concerns about global climate change and rapid diminishing fossil-based petroleum
reserves have spurred immense interest in the utilization of renewable resources for
developing new generation (hydrocarbon) biofuels, with a particular focus on green aviation
fuels
1, 2.
The commonly commercial aviation fuels designed for use in aircraft are jet fuels,
which are exclusively derived from petroleum refining 3. In order to reduce dependence on
fossil sources, the development of alternative fuels for aviation received more attention.
Lignocellulosic biomass is a ubiquitous and sustainable source of carbon that displays
promising potential in the manufacture of hydrocarbon transportation fuels and versatile
chemicals
4, 5.
Furthermore, renewable hydrocarbons based fuels will become drop-in
replacements for petroleum-derived jet fuels if current technological innovations contribute to
competitive production costs 6.
In fact, it is much tougher to develop renewable alternative fuels for utilization in aviation
than automobile application owing to the stringent specifications of jet fuels 7. To produce
green jet fuels (typically C8 - C15 hydrocarbons), Dumesic and co-workers have discovered a
97
new route by aldol condensation of furfural and acetone to synthesize the long carbon chain
intermediates of C8 - C16 8, 9. The hydrodeoxygenation process was stepwise introduced to
obtain the branched alkanes with high thermal stability and energy density, which could be
used as alternative aviation fuels 8. Moreover, other studies emphasized on biomass-derived
techniques to manufacture jet fuel range cyclic alkanes from the lignocellulose-derived
platform compounds
10-13.
However, these aforementioned techniques necessitate expensive
acetone or pinenes as the reactant and noble metal catalysts, resulting in infeasible
commercialization in biorefineries, comparing with commercial formulations based on
petroleum derived jet fuels. Other promising pathways such as hydrotreating of vegetable oils
and Fischer Tropsch synthesis have been intensively investigated with the aim of producing
bio-jet fuels 5. Although these technologies show promising potential in the manufacture of
jet fuel range alkanes, the synthetic feedstocks extracted from plant or animal oils are
commonly upgraded under high hydrogen pressure in the presence of noble metals catalysts
or sulfide-based catalysts 14.
Instead, fast pyrolysis has received special recognition because it is one of the viable process
options to convert lignocellulosic biomass to liquid product
15, 16.
The raw liquid product is
unable for use as a transportation fuel because of detrimental properties, including a much
lower energy density than petroleum fuels, poor thermal and chemical stabilities, and high
viscosity
17, 18.
To circumvent these problems, raw bio-oil has to be upgraded to eliminate
total or partial oxygenates and unsaturated degree prior to its practical application as
transportation fuel
19.
Among the upgrading approaches for jet fuels, hydrodeoxygenation
98
(HDO) is deemed to be a promising and effective process 20. This process normally removes
the oxygen content under high-pressure hydrogen with a catalyst
21,
22.
Most
hydrodeoxygenation researches have concentrated on sulfide Co–Mo and Ni–Mo supported
on γ-Al2O3 catalyst 19. Such conventional sulfide-based catalyst, however, are not satisfactory
for hydrotreating raw bio-oil since the introduction of sulfur-containing compounds (H2S or
thiophene) with the purpose of maintaining the catalysts at the active status, which gives rise
to increasing contamination risk of liquid products
23.
Alumina (Al2O3) support with the
catalysts is well known to be rapidly deactivated by coke deposition and potentially poisoned
by water exist in bio-oils
16, 24.
Noble-based catalysts can also be employed in the processes,
which significantly raise cost for catalysts.
In addition, the HDO process conditions are rather severe (300 – 400 °C, 80 - 300 bar H2
pressure)
17, 18.
Since advanced jet fuels from HDO process are expected to be enlarged in a
biorefinery scale, high capital costs caused by tolerance of severe condition and low
selectivity of liquid products render these processes uneconomical
25, 26.
Accordingly, the
production of jet fuels from renewable biomass resources calls for ideal technologies with
efficient solid-phase catalysts to make the processes economically feasible under a mild
reaction condition. It is more reasonable to use Ni-based catalysts for hydrotreating bio-oil
because of their high activity of hydrogenation and low cost. Raney-type nickel is widely
used as a versatile catalyst in such a hydrotreating process for reductive transformations of
organic compounds 27. Moreover, Raney-type nickel is utilized in other widespread fields for
fine chemical manufacturing owing to its good catalytic performance
99
18.
In addition to
catalyzing liquid products, costly hydrogen derived from aqueous-phase reforming could also
be obtained over a tin-promoted Raney-nickel catalyst by Huber and his colleagues 28.
Typically the commercial and military jet fuels are comprised of paraffins, naphthenes and
aromatics 2. Nonetheless, two main compositions of cyclic paraffins and aromatics in
commercial military jet fuels are hard to attain from well-developed hydrotreating of
vegetables or Fischer-Tropsch synthesis 14. On the other hand, it is widely known that oxygen
content of bio-oil can be partially or even completely eliminated by zeolite cracking during
fast pyrolysis of biomass 29. Carlson et al. have reported that biomass-derived carbohydrates
can be directly converted into aromatics with ZSM-5 as catalyst in a single catalytic pyrolysis
step
30.
Lei and his colleagues have focused on the production of aromatic hydrocarbons
through catalytic microwave-induced pyrolysis of Douglas fir sawdust
31.
Up to 92.60%
selectivity of the liquid organics obtained was jet fuel range (C8 – C15) aromatic hydrocarbons.
In this sense, these aromatics in the bio-oils with low oxygen content can be hydrogenated
into cyclic paraffins and olefins.
Based on aforementioned consideration, the directional conversion of lignocellulosic biomass
into jet fuels production by integrated processes are proposed: (1) lignocellulosic biomass is
firstly converted into jet fuel range (C8 – C15) aromatics by catalytic microwave pyrolysis
over well-promoted ZSM-5 and (2) the bio-oil derived from catalytic microwave pyrolysis is
hydrotreated into desired C8 – C15 hydrocarbons (including paraffins and minor aromatics )
which satisfy basic requirements of conventional jet fuels by using Raney nickel as the
100
catalyst. Since the coupling of two steps (catalytic microwave pyrolysis and downstream
bio-oil hydrotreating) over regular catalysts has not been previously studied, this study
demonstrates proof-of-principle of a novel consecutive two-step process to produce jet fuels
from intact biomass.
4.3 Experimental section
4.3.1 Materials
The feedstock used was Douglas fir sawdust pellets (Bear Mountain Forest Products Inc.,
USA) which were approximately 7 mm in diameter and 15 mm in length with a moisture
content of 8 wt.%. Proximate and elemental analysis of raw Doulgas fir sawdust pellets were
described in Table S1.
Parent ZSM-5(SiO2/Al2O3 Mole Ratio: 50) was purchased from
Zeolyst International, USA. Raney Ni 4200 (slurry in water) in an activated form was
supplied by Sigma-Aldrich Corporation (St. Louis, MO, USA). Phenol (99%), p-cresol (99%),
o-cresol (98%), guaiacol (98%), toluene (99.7%), p-xylene (99%), creosol (98%),
ethylbenzene (99%), 1H-indene (97%), indane (95%), naphthalene (99.6%), tetralin (97%),
1-methylnaphthalene (96%), n-undecane (99%), and n-dodecane (99%) were used as
purchased from Alfa
Aesar (Ward Hill, MA, USA).
Ethylcyclohexane
(99%),
1,2,4-trimethylbenzene (98%), 1,2,4-trimethylcyclohexane (97%), propylcyclohexane (99%),
and 1H-Indene,octahydro- (99%) were supplied by Sigma-Aldrich Corporation (St. Louis,
MO, USA).
101
4.3.2 Catalyst preparation
Parent ZSM-5 was oven dried at 105 °C for 12 h and calcined at 550°C for 5 h in a muffle
furnace. The main characteristics of the catalyst were reported in previous study
32.
Raney
nickel is notorious for its pyrophoricity, and it may ignite spontaneously when dried in air.
The Raney nickel slurry was thus dried at 60 °C till constant weight in the atmosphere of
nitrogen to avoid contact with air, prior to the subsequent catalytic test.
4.3.3 Catalytic microwave pyrolysis of lignocellulosic biomass
A Sineo MAS-II batch microwave oven (Shanghai, China) with a rated microwave power of
1000 W was employed for microwave pyrolysis. Detailed experimental setting was described
in our previous study 31, 33, 34. Fixed loading of Douglas fir sawdust pellets (40 g) were placed
in a 500 mL quartz flask inside the microwave oven. 0.05g activated carbon powder was used
as the absorber for the microwave-assisted pyrolysis. The system was purged with nitrogen
on a flow rate of 1000 mL/min for 15 min prior to pyrolysis reaction to maintain an
oxygen-free environment. The microwave pyrolysis was conducted at the temperature of
480 °C for 10 min. The pyrolysis volatile vapors from the flask passed through a packed bed
catalysis reactor which was filled with catalyst. The packed-bed reactor customized is
constructed of quartz and externally heated by a heating tape. A thermocouple was introduced
between the reactor and the heating tape to measure catalytic temperature. The catalytic
temperature was hold at 375 °C and catalyst (10 g) to reactant ratio was kept constant at 0.25.
The condensable liquid was collected as bio-oil. The non-condensable vapors escaped as gas
at the end of the condensers and were collected for analysis. The catalytic microwave
102
pyrolysis was duplicated for five times in order to gain abundant bio-oil which was stepwise
mixed together for hydrotreating.
4.3.4 Hydrotreatment of bio-oils from catalytic microwave pyrolysis
To produce the jet fuel range hydrocarbons, a closed reaction system with a stirred stainless
batch reactor of the 4592 micro stirred reactor (with a 50mL vessel) and a 4848 reactor
controller from Parr Instrument Company (Moline, IL, USA) is used. For each run, the
bio-oil (5 g) from catalytic microwave pyrolysis is loaded into the reactor together with 5 wt%
or 10 wt% Raney nickel catalysts (in terms of bio-oil). Then the reactor is sealed and vented
for five times with hydrogen to get rid of the air present in the vessel. Hydrogen is
subsequently adjusted to reach the set pressure. The automatic controller is employed to
control the temperature and the revolution of stirrer (300 rpm). The pressure inside the reactor
is recorded and the reactions proceed at set temperatures for the intended time. After the
experiment is finished, stirring is stopped and the reactor is rapidly cooled to ambient
temperature. Then, the gas is collected for analysis and the reactor is depressurized.
Consequently the liquid product is filtered to remove catalyst particles.
4.3.5 Analytical techniques
Elemental analysis (C, H and N) of feestocks, liquid samples, char, and coke deposited on
spent catalysts is conducted using a 2400 Series II CHN/O Elemental Analyzer (PerkinElmer,
USA). Oxygen content is calculated by difference.
103
The chemical composition of the bio-oils is characterized and semi-quantified by Agilent
7890A GC-MS (GC–MS; GC, Agilent 7890A; MS, Agilent 5975C) with a DB-5 capillary
column. The GC is first programmed to heat to 45°C for 3 min followed by heating to 300°C
at a rate of 10°C/min. The injection sample size is 1 μL. The flow rate of the carrier gas
(helium) is 0.6mL/min. The ion source temperature is 230°C for the mass selective detector.
Compounds are identified by comparing the spectral data with that in the NIST Mass Spectral
library. The area percent of compounds obtained from GC/MS results is utilized to predict
products selectivity. Each liquid sample is measured for 3 times to get the average. The
moisture content in the bio-oils is determined by a Karl Fischer (KF) compact titrator (V20
Compact Volumetric KF Titrator, Mettler-Toledo).
The gaseous product is collected in a 1L Tedlar gas bag and then offline analyzed by an
INFICON 3000 Micro-GC (INFICON Inc., Santa Clara, CA, USA) system with a thermal
conductivity detector (TCD). A standard gas mixture consisting of H2, N2, CH4, CO, CO2,
C2H4, C2H6, and C3H6 is used to calibrate the yield of non-condensable gas. Alkanes and
olefins (>C4) in gas samples are either not detected or negligible in this research. During the
stable period, the gas is injected into the Micro-GC every 5 min and each gas sample is also
measured for 3 times to get the average.
4.3.6 Experimental methods and data processing
The coke mass is determined by the difference before and after catalytic pyrolysis. The
weight of non-condensable gas is calculated using the following equation:
104
ℎ
=
−
− ℎ
−
(1)
A central composite experimental design (CCD) is used to optimize the process conditions
with respect to product distribution (Table 4.1). The reaction temperature (X , °C) and
reaction pressure (X
,
psi), are chosen as independent variables. Carbon yield of a specific
product are calculated based on the following equation:
B
C
B
C
D
C
=
H D=
E
C
C
C
× 100%
E
E
(2)
× 100%
(3)
Table 4.1 Experimental design and product yield distribution.a
Entry
Tb
Pc
(ºC)
(psi)
Categories (% in area)
Paraffins
Olefin
Hydro-aromatic
Hydro-cyclic
Aromatic
s
hydrocarbons
oxygenates
hydrocarbons
Phenols
Guaiacols
Other
oxygenates
1
200
76
0.56
1.02
3.30
7.66
20.26
34.46
19.20
13.54
2
200
924
4.39
5.60
20.15
16.67
7.91
8.56
22.84
13.88
3
129
500
0.67
0.57
6.59
6.73
16.51
25.53
29.57
13.83
4
271
500
9.58
2.41
21.64
6.08
11.80
16.87
19.14
12.48
5
150
200
1.94
0.47
9.47
6.08
13.29
24.8
28.45
15.50
6
150
800
1.87
1.17
14.37
19.98
6.97
16.05
26.60
12.99
7
250
200
4.83
1.17
23.18
7.66
22.29
15.19
14.53
11.15
8
250
800
6.86
3.98
13.94
8.77
17.1
14.19
24.70
10.46
9
200
500
2.26
0.54
20.00
14.69
9.74
13.18
29.30
10.29
10
200
500
2.58
1.54
17.55
13.16
12.67
13.78
25.43
13.29
11d
-
-
0
0.24
0.64
0.88
25.52
30.43
24.80
17.49
a
Reaction condition: Raney Nickel, 5 wt% with respect to reactant mass; Reaction time , 2 h.
b
T: Reaction temperature.
c
P: Initial pressure.
d
Entry-11 is the raw bio-oil from catalytic microwave pyrolysis.
105
Fig. 4.1 Product distribution of organics based on carbon number and various functional
groups.
4.4 Results and discussion
4.4.1 Catalytic transformation of lignocellulosic biomass into aromatics
To produce jet fuel range hydrocarbons, the transformation of biomass into jet fuel range (C8
106
– C15) aromatics is primarily required. It was found that the yield of bio-oil was 33.09±2.01
wt%; the water content of bio-oils obtained was close to 50 wt%, which was derived from
both the moisture of raw biomass and the subsequent dehydration reaction during biomass
pyrolysis. However the carbon yield of the bio-oils was merely 22.93±2.00% due to the share
of the water.
The product distribution of organics based on carbon number ranged from C4
to C16 As shown in Fig. 4.1 A. It is noteworthy that the C8 – C15 were found to be the
dominant lumps of chemical compounds, occupying approximately 60%. It was also noted
that the selectivity of undesired carbon number range (C4 – C7) was up to 40%, originating
from uncompleted catalysis of pyrolysis volatiles evolved from decomposition of intact
biomass. The chemical compounds were categorized on the basis of various functional
groups as depicted in Fig 4.1 B. Generally the typical compositions have been categorized
into the lumps of aromatic hydrocarbons, phenols, guaiacols, furans, etc. Aromatic
hydrocarbons (25.52%) except toluene belong to jet fuel range aromatics, including
mono-cyclic and polycyclic aromatic hydrocarbons. The small amount of poly-cyclic
aromatic hydrocarbons like naphthalene was evolved from oligomerization reactions of
mono-cyclic aromatic hydrocarbons. Total amounts of phenols and lignin-derived guaiacols
were the most abundant compounds, most of which corresponding to carbon number
exceeded C8. Other small amount of oxygenates, e.g. furans (10.97%) and esters (1.59%),
mainly comprised of C4 – C7 compounds. Such oxygenates with low carbon numbers is
attributed to decomposition of cellulose and hemicellulose from intact biomass.
107
4.4.2 Hydrotreatment of bio-oils derived from catalytic microwave pyrolysis
4.4.2.1 Chemical composition of hytrotreated oils
Since the bio-oils produced by catalytic microwave pyrolysis of intact biomass principally
consist of C8 - C15 range aromatics, the controllable adjustment of aromatics with 8 - 15
carbon numbers are considered as precursors of jet fuels. In the hydrotreating process, the
directional production of C8 - C15 aromatics was hydrotreated by using Raney nickel under a
low-severity condition. The loss of the hydrotreated bio-oils was ignored if considering the
recovery (more than 95 wt%) of the bio-oils could be achieved after the hydrotreating
experiments. The product distributions are summarized as a function of reaction temperature
and initial pressure in Table 4.1 and selectivity of paraffins is visualized in accordance to the
response surface and contour line in Fig. 4.2. It was observed that paraffins as significant
composition of jet fuels were in the range from 0 to 9.58% depending on alterations of
reaction conditions. Quadratic model equations for selectivity of paraffins (Eq. (4)) as a
function of reaction temperature (X1, °C) and initial pressure (X2) is present the results of the
experiments:
Y.;0;II%J3 = 2.42 + 2.56X + 0.84X + 0.53X X + 1.33X + 0.14X
(4)
The model for yield prediction was reduced by using backward statistical analysis, and
parameters were sequentially removed based on the coefficient’s P-value until all remaining
were significant (P-value < 0.05). For ANOVA analysis, the P-value of Eq. (4) was 0.0002,
which is less than α = 0.05, suggesting that Eq. (4) is significant to present the relationship
between selectivity of paraffins and reaction conditions. The coefficient of determination (R2)
108
for Eq. (4) is 0.95, evidencing that the mode fairly represent the relationships among the
independent variables.
Fig. 4.2 Effect of the interaction of the independent variables on selectivity of paraffins.
Unsaturated hydrocarbons (hydro-aromatic hydrocarbons and olefins) partially hydrotreated
in the process accounted for 0.64 – 23.18% and 0.24 – 5.60%, respectively. Reaction
conditions also had a crucial influence on the selectivity of aromatic hydrocarbons,
occupying from 6.97 to 25.52%. Oxygenated compositions (phenols, guaiacols, and other
oxygenates) showed remarkable variations influenced by reaction conditions, varying from
8.56 to 30.43%, 14.53 to 29.57%, and 10.29 to 17.49%, respectively. These results indicate
that hydrogenation and hydrogenolysis reactions impacted by reaction conditions
simultaneously determined the selectivity of these compositions in hydrotreated oils.
109
4.4.2.2 The effect of reaction temperature on the chemical composition of hydrotreated
oils
In order to further understand chemical reactions in the process and obtain more insight into
hydrotreated oils, the chemical compounds of hydrotreated oils are elucidated as a function of
reaction temperature in Table 4.2. It can be seen that the reaction temperature had a
significant effect on the hydrotreated reactions. In comparison with the bio-oil obtained by
catalytic microwave pyrolysis of biomass, the amount of total paraffins progressively
increased in hydrotreatd oils. It was noticed that paraffins were in the range from 0.67% at
129 °C to 10.90% as reaction temperature increased to 300 °C, implying that a higher
reaction temperature accelerated the formation of paraffins. Paraffins obtained in jet fuel
range
were
predominantly
comprised
of
ethylcyclohexane,
propylcyclohexane,
octahydro-1H-indene, trimethylcyclohexane, and their derivatives. It was observed that the
carbon selectivity of specific paraffins (e.g. ethylcyclohexane) was improved with the
increase of the reaction temperature.
Nonetheless, it is well known that Raney nickel is very
hydrophilic, and the presence of water exert a deteriorating effect on the catalytic activity of
Raney nickel.35, 36 Owing to water content (close to 50 wt%) in parent oils, water molecule
was easily fixed to the Raney nickel surface by oxygen bonding, thereby inhibiting the
adsorption of organic substrates to the metallic surface.36 Therefore, the catalytic activity was
limited, reducing the degree of hydrotreated reactions to produce more paraffins. We have
determined that when the organic phase and aqueous phase were separated by a separating
funnel, the selectivity of paraffins significantly increased (results not shown) after the
hydrotreating reaction of the organic phase at 200 °C, comparing to the result from the
110
hydrotreatment of mixed bio-oils conducted at the same reaction condition. In order to
improve the yield of paraffins, our further efforts should be directed toward lowering water
content in parent oils for hydrotreating upgrading.
Table 4.2 Major paraffins and olefins in hydrotreated oils as a function of reaction
temperaturea
Reaction temperature ( ºC)
Raw
129
200
271
300
oil
Paraffins
Cyclohexane, ethyl-
-
-
0.18±0.01 1.35±0.23 1.40±0.22
Pentane, 2,3,4-trimethyl-
-
-
Cyclohexane, propyl-
-
-
Cyclohexane, (1-methylethyl)-
-
-
-
1.07±0.06 1.17±0.10
Cyclohexane, 1,1,2-trimethyl-
-
-
-
0.31±0.02 1.27±0.13
Hexane, 2,3,4-trimethyl-
-
-
-
0.78±0.08 0.49±0.04
1H-Indene, octahydro-, cis-
-
-
1-Ethyl-2,2,6-trimethylcyclohexane
-
-
Undecane
-
Dodecane
-
-
Other paraffins
-
-
0.76
0.89
1.81
Total (% in area)
0
0.67
2.58
9.58
10.90
1,4-Pentadiene
-
-
-
2-Pentene
-
-
0.35±0.03 0.46±0.05 0.32±0.03
2,4,4-Trimethyl-1-hexene
-
-
0.48±0.04 0.53±0.03 0.77±0.08
Cyclohexane, ethenyl-
-
0.19±0.01
-
-
0.85±0.10
0.24
0.38
0.71
0.77
-
-
0.47±0.05 0.71±0.08
0.34±0.04 1.24±0.13 0.51±0.04
0.44±0.04 1.53±0.09 1.46±0.15
-
0.67±0.08 0.43±0.03
0.67±0.02 0.49±0.04 1.02±0.12 1.13±0.11
0.37±0.05 0.25±0.03 0.52±0.05
Olefins
Other olefins
111
0.65±0.08 0.35±0.02
Total (% in area)
aReaction
0.24
0.57
1.54
2.41
2.29
condition: Initial pressure, 500 psi; Raney Nickel, 5 wt% with respect to reactant
mass; Reaction time, 2 h.
Furthermore, there is a minor selectivity for olefins, such as vinylcyclohexane, increasing
from 0.24 to 2.29% alongside the increment of reaction temperatures. Hydro-aromatic
hydrocarbons were mainly derived from polycyclic aromatic hydrocarbons, which
experienced a raised tendency (0.64 – 23.94%) as the reaction temperature went up (Table
4.3); whereas aromatic hydrocarbons displayed a decreased trend as reaction temperatures
increased. It was noteworthy that neither double-cyclic nor polycyclic aromatic hydrocarbons
were detected under the conditions conducted at 271 and 300 °C. These tendencies imply that
hydro-aromatic compounds like tetralin were evolved from hydrogenation reactions of
naphthalene and its derivatives. However, the inadequate ring reduction of polycyclic
aromatic hydrocarbons is probably because of a conformational effect.27 Two aromatic rings
of naphthalene were coplanar and one aromatic ring was prone to be hydrogenated into a
saturated ring; however the aromatic ring and cyclic alkane ring of tetralin was not coplanar
after hydrogenation. The other aromatic ring was thus directed away from the catalyst surface,
resulting in lack of reduction.27 Although polycyclic aromatic hydrocarbons were thoroughly
converted into hydro-aromatic hydrocarbons, there were mono-cyclic aromatic hydrocarbons
including ethylbenzene and p-xylene existing in hydrotreated oils even at 300 °C. It is
affirmed that the single ring hydrogenation of aromatic hydrocarbons appeared to occur much
slower, which is probably attributed to steric hindrance caused by alkyl groups.27
112
Table 4.3 Major chemical compounds in hydrotreated oils as a function of reaction
temperature.a
Reaction temperature ( ºC)
Raw oil
129
200
271
300
4-Methylindane
-
1.84±0.18
2.21±0.23
2.01±0.20
2.26±0.23
Tetralin
-
1.38±0.12
2.65±0.27
3.02±0.30
2.50±0.26
4,7-Dimethylindan
-
-
1.12±0.09
1.83±0.20
1.93±0.20
2-Methyltetralin
-
1.17±0.13
1.78±0.18
0.98±0.09
1.70±0.14
5,6-Dimethylindan
-
0.10±0.01
0.16±0.01
0.34±0.02
0.27±0.03
6-Methyltetralin
-
-
-
2.58±0.26
2.44±0.20
5-Methyltetralin
-
0.80±0.09
3.13±0.26
0.45±0.05
1.06±0.09
2,7-Dimethyltetralin
-
-
2.41±0.23
3.21±0.32
2.62±0.27
1,8-Dimethyltetralin
-
-
-
0.87±0.06
1.43±0.12
1,4-Dimethyltetralin
-
-
0.22±0.02
0.32±0.02
0.55±0.04
6-Ethyltetralin
-
0.25±0.03
0.90±0.07
1.21±0.18
1.26±0.13
5,6-Dimethyltetralin
-
-
0.34±0.03
0.45±0.05
6-propyltetralin
-
-
-
0.21±0.02
0.19±0.22
1,1-Dimethyltetralin
-
-
-
0.54±0.06
0.28±0.05
0.64±0.07
0.40±0.05
-
0.09±0.02
0.14±0.03
-
0.65
2.97
3.64
4.82
0.64
6.59
17.55
21.64
23.9
Cyclopentanone
-
1.55±0.16
-
0.58±0.06
-
2-Methylcyclopentanone
-
0.61±0.06
0.29±0.02
0.32±0.04
-
Cyclohexanol
-
0.65±0.07
3.59±0.35
1.53±0.18
3.67±0.30
Hydro-aromatic hydrocarbons
2,5,8-Trimethyltetralin
Others
Total (% in area)
Hydro-cyclic oxygenates
113
2-Methylcyclohexanol
-
-
0.58±0.06
0.32±0.04
0.85±0.07
3-Methylcyclohexanol
-
-
1.06±0.10
0.25±0.02
1.28±0.14
γ-Valerolactone
-
0.32±0.03
0.68±0.07
-
0.89±0.09
1,2-Cyclohexanediol
-
0.21±0.02
2.13±0.23
0.65±0.06
2.77±0.28
Others
0.88
3.39
4.83
2.43
3.04
Total (% in area)
0.88
6.73
13.16
6.08
12.5
Toluene
0.31±0.03
-
0.68±0.07
0.54±0.04
-
Ethylbenzene
0.30±0.02
0.21±0.02
0.46±0.05
0.34±0.03
0.39±0.03
p-Xylene
2.71±0.30
1.29±0.13
1.70±0.14
1.67±0.18
1.72±0.18
1,2,3-Trimethylbenzene
2.36±0.21
1.30±0.10
2.18±0.19
1.32±0.15
0.89±0.07
Indane
0.78±0.06
2.11±0.21
-
1.24±0.15
1.58±0.14
Indene
1.17±0.10
-
-
-
-
Naphthalene
3.56±0.36
2.49±0.19
-
-
-
1-Methylnaphtalene
2.39±0.20
0.70±0.08
-
-
-
2,7-Dimethylnaphthalene
3.59±0.30
3.55±0.28
0.30±0.02
-
-
Others
8.35
4.86
7.35
6.69
3.14
Total (% in area)
25.52
16.51
12.67
11.80
7.66
Aromatic hydrocarbons
a Reaction
condition: Initial pressure, 500 psi; Raney Nickel, 5 wt% with respect to reactant
mass; Reaction time, 2 h.
114
OH
OH
O
OH
O
CH3
H2
O
H2
OH
Hy
dr o
g
H2
H2
en
a tio
n
OH
Hydrogenation
Hydrogenolysis
Hydrogenation
Hydro genoly sis
H2
OH
CH3
H2
Hydrogenation
OH
O
OH
OH
OH
H2
H2
OH
Hydrogenation
Fig. 4.3 Proposed reaction pathway for converting phenolic monomers (phenol, guaiacol,
catechol) into cyclohexanol and 1, 2-cyclohexanediol.
Reaction temperatures also had a significant influence on the selectivity of oxygenated
compounds. The total amount of hydro-cyclic oxygenates including cyclohexanol and 1,
2-cyclohexanediol increased as the reaction temperature was raised in Table 4.4; whilst the
total amount of phenols (e.g. phenol, phenol, 2-methyl, phenol, 2, 5- dimethyl) revealed an
opposite trend steadily declining from 30.43 to 6.91%.
Likewise, there were distinct
downward tendency of guaiacols (such as guaiacol, creosol) gradually decreased from 24.80
to 19.79%.
It was noteworthy that a maximum selectivity of hydro-cyclic oxygenates
(12.50%) was gained at the catalytic temperature of 300 °C; whereas both phenols (6.91%)
115
and guaiacols (19.79%) showed low selectivity at this condition. These observations are the
key point to propose the reaction mechanisms of product conversion among these compounds.
Proposed reaction pathway for converting the most representative phenolic monomers
(phenol, guaiacol, and catechol) into cyclohexanol and 1, 2-cyclohexanediol is described in
Fig. 4.3. Phenol was either directly hydrogenated to cyclohexanol in the presence of Raney
nickel or hydrogenated to cyclohexanone acting as intermediate compounds. Cyclohexanone
was progressively hydrogenated to form the final product (cyclohexanol). Catechol derived
from lignin depolymerization has two adjacent hydroxyl groups attached on the aromatic ring.
Generally it was hydrogenated at the aromatic ring producing the intermediate compound
(2-hydroxycyclohexanone), which was further converted into 1, 2-cyclohexanediol via
hydrogenation.37 Meanwhile, the elevated reaction temperature also possibly favored the
scission of hydroxyl group from phenyl to form phenol through hydrogenolysis reactions.
Guaiacol with adjacent methoxy and hydroxyl functional groups could readily be
hydrogenated at the aromatic ring to saturated intermediates.26 Methoxy functional group of
the intermediate was hydrogenated to hydroxyl group, finally forming 1, 2-cyclohexanediol.
Other parallel reaction pathways of guaiacol conversion (directional hydrogenation and
hydrogenlysis of guaiacol methoxy functional groups) generated catechol and phenol,
respectively. However, the total amount of other oxygenates fluctuated in a tight range
(~15%), including propylene glycol and ethyl butanoate.
116
Table 4.4 Major chemical compounds in hydrotreated oils as a function of reaction
temperature.a
Reaction temperature ( ºC)
Raw oil
129
200
271
300
Phenol
3.94±0.35
4.51±0.45
-
-
-
2-Methylphenol
2.53±0.26
3.18±0.30 1.94±0.20
-
-
3-Methylphenol
-
-
4-Methylphenol
3.85±0.35
4.84±0.50
2,5-Dimethylphenol
3.42±0.32
3.23±0.30 3.58±0.36 3.41±0.30 3.71±0.30
3-Ethylphenol
1.20±0.10
1.10±0.08
Phenols
2.26±0.23 1.25±0.12 0.39±0.05
-
-
3.58±0.36
-
1.13±0.08
-
0.50±0.04 0.69±0.06 0.20±0.02
-
p-Cresol, 2-ethyl-
-
m-Cresol, 4-ethyl-
1.24±0.15
0.39±0.05
-
-
Other phenols
14.25
7.78
4.18
5.18
2.18
Total (% in area)
30.43
25.53
13.78
16.87
6.91
2-Ethyl-5-methylphenol
-
1.14±0.10 0.37±0.03
1.13±0.10 0.98±0.10 0.26±0.03
Guaiacols
Guaiacol
4.33±0.40
5.54±0.60 4.08±0.40 2.15±0.20 1.97±0.20
2-Methoxy-3-methylphenol
0.36±0.04
0.43±0.05
Creosol
6.36±0.55
7.96±0.85 7.39±0.70 6.21±0.55 6.02±0.51
p-Ethylguaiacol
3.51±0.28
4.73±0.50 6.36±0.61 5.24±0.55 5.23±0.35
4-Propylguaiacol
3.55±0.32
4.95±0.35 5.82±0.50 4.84±0.40 4.87±0.42
Other guaiacols
6.69
5.96
1.78
0.70
1.70
Total (% in area)
24.80
29.57
25.43
19.14
19.79
-
-
-
Other oxygenates
Propylene glycol
-
1.84±0.20 2.36±0.20 2.25±0.20 2.14±0.23
Ethyl butanoate
0.91±0.08
1.74±0.14 1.97±0.20 2.01±0.18 2.29±0.32
Butyl acetate
0.17±0.02
0.30±0.02 0.39±0.30 0.35±0.05 0.37±0.04
Others
16.41
9.95
117
8.57
7.87
11.25
Total (% in area)
aReaction
17.49
13.83
13.29
12.48
16.05
condition: Initial pressure, 500 psi; Raney Nickel, 5 wt% with respect to reactant
mass; Reaction time, 2 h.
Fig. 4.4 Chemical composition of hydrotreated oils on the basis of initial pressure. Reaction
condition: Initial pressure, 500 psi; Raney Nickel, 5 wt% with respect to reactant mass;
Reaction time, 2 h.
4.4.2.3 The effect of initial pressure on the chemical composition of hydrotreated oils
The chemical composition of hydrotreated oils on the basis of initial pressure is shown in Fig.
4.4. Compared with the parent oils from catalytic microwave pyrolysis without any paraffins
(alkanes) detected, this content increased gradually from 0.56 to 5.73% with the increase of
initial pressure from 76 to 1000 psi.
The olefins fraction was fluctuant caused by initial
pressure, which first increased from 0.24% (the parent oil) to 5.6% (initial pressure of 924 psi)
and then declined to 1.12% with the increment of initial pressure to 1000 psi. As such, the
118
content of hydro-cyclic oxygenates had the same tendency as olefins from 0.88 to 9.63%. It
was observed that the selectivity of hydro-aromatic hydrocarbons remarkably went up from
0.64 to 21.22% as the initial pressure increased. On the contrary, the high amount of phenols
showed a considerably declined trend from 30.43 to 9.58% in the increasing range of initial
pressure. Unlike the phenols, guaiacols and other oxygenates steadily fluctuated, which
suggests that the selectivity of these contents could not be impacted by initial pressure. These
outcomes imply that a high initial pressure, which means that more volume of hydrogen was
provided on the Raney nickel surface, contributed to more hydrogenated reactions occurring
in the reactor.
Fig. 4.5 Chemical composition of hydrotreated oils as a function of reaction time. Reaction
condition: Reaction temperature, 200 ºC; Initial pressure, 500 psi; Raney Nickel, 5 wt% with
respect to reactant mass.
119
4.4.2.4 The effect of reaction time on the chemical composition of hydrotreated oils
Apart from reaction temperature and initial pressure, reaction time was another crucial factor
that influenced product distribution as shown in Fig. 4.5. Especially, the product distribution
in the resulting product prominently shifted towards the paraffins (0 – 12.63%) with
increasing reaction time, suggesting that hydrogenating and hydrogenolysis reactions were
simultaneously enhanced for longer reaction duration. The amount of olefins first appeared to
increase slightly, and then dramatically decreased, finally increased to 3.2% at prolonged
reaction time. The total amount of hydro-aromatic hydrocarbons kept at high selectivity,
ranging from 12.78 to 25.60%; while the selectivity of hydro-cyclic oxygenates experienced a
downward trend from 17.49 to 4.85%. As prolonged reaction time could give rise to ring
reduction of aromatic hydrocarbons, which resulted in the decrease of aromatic hydrocarbons
to form hydro-aromatic hydrocarbons. Owing to more time provided for the reaction,
hydrogenolysis was likely facilitated by means of the scission of hydroxyl and methoxy
groups on the aromatic ring. Wherefore these aromatic hydrocarbons produced from the
process of hydrogenolysis could offset the consumption for aromatic hydrocarbon
hydrogenation, thereby keeping the selectivity at approximately 11%. It was noted that the
total amount of phenols and guaiacols was impacted by prolonged reaction time especially
going from 10 to 12h. There were totally 19.81% of phenols and guaiacols existing in
hydrotreated oils when reaction time was conducted for 12 h; whilst other non-aromatic
oxygenates significantly increased to 29.13% at this condition. These non-aromatic
oxygenates probably came from hydrogenolysis of hydro-cyclic oxygenates after
120
hydrogenation of phenols and guaiacols. These results indicate that both hydrogenolysis and
hydrogenation reactions jointly occurred on the surface of Raney nickel in the reactor.
Fig. 4.6 Chemical composition of hydrotreated oils on the basis of catalyst loading. Reaction
condition: Reaction temperature, 200 ºC; Initial pressure, 500 psi; Reaction time, 2h.
4.4.2.5 The effect of catalyst loading on the chemical composition of hydrotreated oils
The effect of catalyst loading (5 and 10 wt% with respect to reactant mass) was determined
by two separate experiments as explained in Fig. 4.6. Catalyst loading did not have a vital
influence on the selectivity of paraffins and olefins, which were slightly fluctuant at 2.5% and
1.2%, respectively. It is obvious that the selectivity of both hydro-aromatic hydrocarbons and
hydro-cyclic oxygenates increased when more Raney nickel was added in the reactor. In
contrast, the total amount of phenols, guaiacols, and other oxygenates gradually declined as
121
more catalysts were employed. These results indicate that more catalyst loading, which
means that more active sites were offered on the Raney nickel surface, gave rise to more
hydrogenation and hydrogenolysis reactions.
4.4.3 Analysis of gaseous fraction
The composition of gaseous fraction were detected and quantified by Micro-GC, which
would be helpful to determine the reaction pathways in the reactor. The composition of
gaseous phase as a function of reaction temperature, reaction time, and catalyst loading were
presented in Fig 4.7 A, B, C, respectively. As for all experiments, unreacted hydrogen was
detected at the end of reactions, implying that the reactions were not carried out under
hydrogen starved conditions.
The main component generated in terms of the reactions was
carbon dioxide due to various decarboxylation reactions. There was no carbon monoxide
detected in the gaseous fraction, which may be caused by the water-gas-shift reaction.20, 38
The formation of carbon monoxide from decarbonylation reaction with water content was
thereafter converted into carbon dioxide and hydrogen in the presence of Raney nickel.
Others (e.g. ethane, ethylene, propane) were also present in the reactions, which could be
formed from hydrogenolysis of lager hydrocarbons as intermediates produced during the
process.
The composition of gaseous fraction with respect to catalytic temperature is depicted in Fig.
4.7 A. Hydrogen was the dominant composition from 98.01 to 87.29 vol.%, which suggests
that hydrogenation and hydrogenolysis reactions consuming more volume of hydrogen were
122
enhanced by increasing reaction temperature. The amount of methane and carbon dioxide
gradually went up as reaction temperatures increased, which verifies that decarboxylation and
oligomerization reactions on the catalyst were promoted by increasing reaction temperature.
The development of the gaseous composition as a function of reaction time is illustrated in
Fig. 4.7 B. The yield of carbon dioxide was found to increase in the most pronounced way
from 3.67 to 14.26 vol.% in the range of 2 to 12 h, which indicates that prolonged time
performed in the process accelerated the decarboxylation reactions. The concentration of
hydrogen progressively decreased from 94.78 to 83.51 vol.% due to the hydrogen-consumed
reactions. The volume of methane saw a slight increasing trend with increasing reaction time.
Therefore, prolonged time could facilitate the degree of hydrogenation and hydrogenolysis
reactions. The composition of gaseous fraction on the basis of catalyst loading is shown in
Fig 4.7 C. It was observed that with the addition of more Raney nickel, trends of the results
were identical to those of reaction temperatures and reaction time. The amount of hydrogen
showed a downward tendency from 94.78 to 91.52 vol.%; whereas both carbon dioxide and
methane predicted upward tendencies. It is reaffirmed that more catalysts provided for the
process could improve the degree of hydrogenation and hydrogenolysis reactions.
123
Fig. 4.7 Effects of Reaction temperature (A), Reaction time (B), catalyst loading (C) on the
composition of gaseous fraction at same initial pressure of 500 psi.
124
4.5 Conclusions
This study for proof-of-concept of coupling catalytic microwave pyrolysis with downstream
hydrotreating upgrading using modified ZSM-5 and Raney nickel has revealed that the
integrated processes are potentially profound approach when targeting production of jet fuel
range paraffins and aromatics from lignocellulosic biomass. The integral processes detailed
here were demonstrated to deliver up to 12.63% selectivity of jet fuel range alkanes and
19.48% selectivity of hydro-aromatic hydrocarbons for 12 h hydrotreating reaction.
Increasing reaction temperature, initial pressure, and prolonged time could enhance the
hydrogenation and hydrogenolysis reactions to form jet fuel range paraffins and aromatics.
The gaseous fraction was mainly composed of hydrogen, carbon dioxide, and methane. The
integrated processes within inexpensive catalysts under the low-severity condition provide a
new route that specifically targets jet fuel range paraffins and aromatics from intact biomass.
125
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128
CHAPTER FIVE
FROM LIGNOCELLULOSIC BIOMASS TO RENEWABLE
CYCLOALKANES FOR JET FUELS
5.1 Abstract
A novel pathway was investigated to produce jet fuel range cycloalkanes from intact biomass.
The consecutive processes for converting lignocellulosic biomass into jet fuel range
cycloalkanes principally involved the use of the well-promoted ZSM-5 in the process of
catalytic microwave-induced pyrolysis and Raney nickel catalysts in the hydrogen saving
process. Up to 24.68% carbon yield of desired C8 – C16 aromatics was achieved from
catalytic microwave pyrolysis at 500 °C. We observed that solvents could assist in the
hydrogenation reaction of naphthalene; and the optimum result for maximizing the carbon
selectivity (99.9%) of decalin was from the reaction conducted in the n-heptane medium. The
recovery of organics could reach ~94 wt. % after the extracting process. These aromatics in
the n-heptane medium were eventually hydrogenated into jet fuel range cycloalkanes. Various
factors were employed to determine the optimal result under mild conditions. Increasing
catalyst loading, reaction temperature, and prolonged time could enhance the hydrogenation
reactions to improve the selectivity of jet fuel range cycloalkanes. Three types of
hydrogenation catalysts (NP Ni, Raney-Ni 4200, home-made Raney Ni) were chosen to
evaluate the catalytic performance. Results indicated that the home-made Raney nickel was
the optimal catalyst to obtain the highest selectivity (84.59%) towards jet fuel range
129
cycloalkanes. These cycloalkanes obtained can be directly used as additives to make the
desired jet fuels by blending with other hydrocarbons. Hence integrating catalytic processes
and conversion of lignocellulosic biomass pave a new avenue for the development of green
bio-jet fuels over inexpensive catalysts under mild conditions.
Keywords: Bio-jet fuels; renewable cycloalkanes, Raney Ni catalyst; hydrogenation;
lignocellulosic biomass
5.2 Introduction
Growing concerns about global climate change and rapid diminishing fossil-based
petroleum reserves have spurred immense interest in the utilization of renewable
resources for developing new generation (hydrocarbon) biofuels, with a particular
focus on green aviation fuels.1,
2
The development of alternative aviation fuels
received more attention in order to reduce dependence on fossil sources.
Lignocellulosic biomass is a ubiquitous and sustainable source of carbon that displays
promising potential in the manufacture of hydrocarbon transportation fuels and
versatile chemicals.3, 4
The commercial aviation fuels designed for use in aircrafts are jet fuels, which are
exclusively derived from petroleum refining.5 The current jet fuels originated from
fossil resources are rich in straight-chain alkanes.6,
7
It was found that both
straight-chain alkanes and branched-chain alkanes are highly susceptible to pyrolytic
130
cracking, resulting in coking in a jet engine, increasing deposition and fouling of fuel
control units and injectors, and increasing operating costs due to the increased
maintenance.8, 9 On the other hand, the thermal stability of a jet fuel can be greatly
enhanced by utilizing liquids rich in cycloalkanes; because cycloalkanes are compact
molecules within robust ring strain and contribute to a more dense jet fuel and burn
cleanly with high heats of combustion and low freezing points, comparing with
straight-chain alkanes and branched-chain alkanes with lower densities (~ 0.76-0.78
g/mL).8,
10
However, commercial petroleum fuels normally do not have very high
cycloalkane contents and it is difficult to concentrate the cycloalkanes on a
commercial scale.11 To fulfill the demand of increasing the density and volumetric
heating, it is imperative for cycloalkanes to be synthesized and added into jet fuels
(e.g. Jet A and JP-8).7 Therefore, most efforts to produce jet fuels have focused on
increasing the cycloalkanes content, for example JP-5 navy fuel contains 52.8%
cycloalkanes.12
To produce renewable drop-in fuels for aviation, Dumesic group13, 14 and Huber et al.6,
15
have discovered a new route by aldol condensation of furfural and acetone to
synthesize the long carbon chain intermediates of C8 - C15. The hydrodeoxygenation
process was stepwise introduced to obtain the branched alkanes with high thermal
stability and energy density, which could be used as drop-in fuels for aviation. In
general, most studies emphasized on manufacturing jet fuel range (C8 - C16)
straight-chain alkanes and/or branched-chain alkanes from the lignocellulose-derived
131
platform compounds, which cannot meet the specification of cycloalkanes in jet
fuels.16-18 In addition, these multiple processes necessitate expensive reactants and
noble metal catalysts, resulting in infeasible commercialization in biorefineries
comparing with commercial formulations based on petroleum derived jet fuels.
In addition, fast pyrolysis has also received special recognition and it is one of the
viable process options to convert lignocellulosic biomass to pyrolysis oil.19, 20 The
pyrolysis oil is unable for use as a transportation fuel because of detrimental
properties, including a much lower energy density than petroleum fuels, poor thermal
and chemical stabilities, and high viscosity.21, 22 To circumvent these problems, raw
pyrolysis oil has to be upgraded to eliminate total or partial oxygenates and
unsaturated degree prior to its practical application as transportation fuel.23 Among the
upgrading approaches for jet fuels, hydrodeoxygenation (HDO) is deemed to be a
promising and effective process.7 The HDO process conditions are rather severe (300
– 400 °C, 80 - 300 bar H2 pressure).21, 22 It is widely known that the oxygen content in
the fast pyrolysis oils was removed through HDO process in the form of H2O. In this
regard, large amount of costly hydrogen was consumed by the dehydration reaction,
which significantly reduces the hydrogenating efficiency. Since advanced jet fuels
from HDO process are expected to be enlarged in a biorefinery scale, high capital
costs caused by tolerance of severe condition and low selectivity of liquid products
render these processes uneconomical.24,
25
Accordingly, the production of jet fuels
from renewable biomass resources calls for ideal technologies with efficient
132
solid-phase catalysts to make the processes economically feasible under a mild,
environmentally friendly reaction condition.
In contrast, oxygen content of bio-oil can be partially or even completely eliminated
by zeolite cracking during fast pyrolysis of biomass.26 Carlson et al. have reported that
biomass-derived feedstock can be directly converted into aromatics with ZSM-5 as
catalyst in a single catalytic pyrolysis step.27, 28 Lei and his colleagues have focused on
the production of aromatics through catalytic microwave-induced pyrolysis of
biomass-derived feedstock.29,
30
Up to 92.60% selectivity towards aromatic
hydrocarbons was obtained, belonging to jet fuel range (C8 – C16) aromatics. In this
sense, the certain range aromatics with low oxygen content are prone to be
hydrogenated into cycloalkanes together with minor aromatics under low-severity
conditions. The hydrogenated process is regarded as a hydrogen saving process, in
which hydrogen is converted in the form of cycloalkanes rather than water.
It is more reasonable to use Ni-based catalysts for hydrotreating bio-oil because of
their high activity of hydrogenation and low cost. Raney-type nickel is widely used as
a versatile catalyst in such a hydrotreating process for reductive transformations of
organic compounds.31 Moreover, Raney-type nickel is utilized in other widespread
fields for fine chemical manufacturing owing to its good catalytic performance.22 In
addition to catalyzing liquid products, costly hydrogen derived from aqueous-phase
reforming could also be obtained over a tin-promoted Raney-nickel catalyst by Huber
133
and his colleagues.32 It was found that water molecule was easily fixed to the Raney
nickel surface by oxygen bonding, thereby inhibiting the adsorption of organic
substrates to the metallic surface.33 Hence the removal of water in bio-oils is an
essential step to enhance the catalytic performance in the hydrogen saving process.
Some solvents (e.g. ethyl acetate) can be used to extract the organic mixture from
pyrolysis oils. However, far too little attention has been paid to evaluate the impact of
the solvents as part of the catalytic system.34 Unlike the organic synthesis literature, in
which the influence of solvents has been well-demonstrated and understood; while the
effect of solvents for hydrogenation and hydrogenolysis reactions are commonly
neglected and/or rarely comprehended.35 Thus advance towards understanding the
effect of solvents on the hydrogenation of bio-oils can lay the foundations for
valorization of bio-oils.
According to aforementioned consideration, the direct conversion of lignocellulosic
biomass into jet fuels by integrated processes was conducted. Lignocellulosic biomass
was firstly converted into jet fuel range (C8 – C16) aromatics by catalytic microwave
pyrolysis over well-promoted ZSM-5. The organic phase of the bio-oils derived from
catalytic microwave pyrolysis was thereafter extracted by the optimum solvent after
evaluating the solvent influence on the hydrogenation reaction of model compounds.
In the final step, the organic mixture was hydrogenated into desired C8 – C16
hydrocarbons (including cycloalkanes and minor aromatics) which satisfy basic
requirements of conventional jet fuels by using Raney nickel as the catalysts. Since the
134
coupling of three main steps (catalytic microwave pyrolysis, liquid-liquid extraction,
and downstream hydrogen saving process) over regular catalysts has not been
previously reported on the direct use of lignocellulosic biomass, this study
demonstrates proof-of-principle of a novel consecutive three-step process to produce
cycloalkanes for jet fuels.
5.3 Experimental
5.3.1 Materials
The feedstock used was Douglas fir sawdust pellets (Bear Mountain Forest Products
Inc., USA) which were approximately 7 mm in diameter and 15 mm in length. It was
evidenced that pelleted biomass could be maximally decomposed into pyrolytic
volatile vapors by microwave heating, which could be thereafter catalyzed into
targeted products.29,
30
Parent ZSM-5 (SiO2/Al2O3 Mole Ratio: 50) was purchased
from Zeolyst International, USA. Methanol (99.8%), 2-propanol (99.7%), ethyl acetate
(99.5%), benzene (99.5%), n-hexane (99%), cyclohexane (99%), toluene (99.7%),
n-heptane (99%), methylcyclohexane (99%), p-xylene (99%), ethylbenzene (99%),
1H-indene (97%), indane (95%), naphthalene (99.6%), 1,2,3,4-tetrahydronaphthalene
(97%),
decahydronaphthalene
(98%),
1-methylnaphthalene
(96%),
2-methylnaphthalene (97%), n-dodecane (99%), anthracene (99%), pyrene (98%),
Nickel-Aluminum alloy powder in an non-activated type were used as purchased from
Alfa
Aesar
(Ward
Hill,
MA,
USA).
1,2-dimethylcyclohexane
(99%),
1,3-dimethylcyclohexane (99%), 1,4-dimethylcyclohexane (99%), ethylcyclohexane
135
(99%),
1,2,4-trimethylbenzene
propylcyclohexane
(99%),
(98%),
1,2,4-trimethylcyclohexane
hexahydroindan
(99%),
bicyclohexyl
(97%),
(99%),
perhydrofluorene (97%), tetradecahydroanthracene (99%), Raney Ni 4200 (slurry in
water) in an activated form were supplied by Sigma-Aldrich Corporation (St. Louis,
MO, USA).
5.3.2 Catalyst preparation
The activity of parent ZSM-5 was improved by suffering both hydrothermal and
calcined treatments. Under the gentle stirring, parent ZSM-5 powder was added into
deionized water (mass ratio=1) at 60 °C. After addition, the mixture was kept on
stirring for 2 h under this condition. The slurry was then dried at 105 °C till constant
weight. The sequential process was the catalyst calcination: hydrothermally treated
ZSM-5 was calcined at 550 °C for 5 h in a muffle furnace. The catalysts were
pelletized and sieved to 20 – 40 mesh. The main characteristics of the catalyst were
reported in our previous study.36
Non-pyrophoric Raney nickel (referred as NP Ni) was prepared from metallic alloy
powders. Under gentle stirring, Ni-Al alloy powders (2 g) were slowly impregnated
with 1.2 wt. % NaOH aqueous solution (20 mL) at room temperature. After addition,
the temperature was elevated from room temperature to 80 °C and hold at the
temperature for 30 min. Additional 2 mL of 12 wt. % NaOH aqueous solution was
added to the slurry and stirred gently at 80 °C for 30 min for further alkali leaching.
136
Subsequently the sample was washed to neutrality using distilled water and reserved in
water for future catalytic use.
The home-made Raney nickel catalyst was developed using 20 wt. % NaOH aqueous
solution to remove Al in the following procedure. 2 g of the above Ni-Al was slowly
added into 20 mL NaOH aqueous solution under gentle stirring. After addition, the
slurry was kept on stirring at 80 °C for 1 h. The excess of sodium hydroxide was
finally washed with distilled water until nearly neutral pH was reached. The obtained
Raney nickel catalyst was stored in water. In order to test the catalytic performance of
Raney nickel manufactured from the process, Raney nickel 4200 (referred as
Raney-Ni 4200) purchased in the activated form was used as criterion for catalyzing
the hydrogenation reaction of bio-oils. Raney nickel is notorious for its pyrophoricity,
and it may ignite spontaneously when dried in air. The Raney-Ni 4200 slurry and
home-made Raney nickel catalyst were thus dried at 60 °C till constant weight in the
atmosphere of nitrogen to avoid contact with air, prior to the subsequent catalytic test.
5.3.3 Catalytic microwave-induced pyrolysis of lignocellulosic biomass
A Sineo MAS-II batch microwave oven (Shanghai, China) with a rated microwave power of
1000 W was employed for microwave pyrolysis. Detailed experimental setting was described
in our previous studies.29, 37 Fixed loading of Douglas fir sawdust pellets (25 g) for each run
were placed in a 500 mL quartz flask inside the microwave oven. 0.05g of activated carbon
powder was used as the absorber for the microwave-assisted pyrolysis. The system was
137
purged with nitrogen on a flow rate of 1000 mL/min for 15 min prior to pyrolysis reaction to
maintain an oxygen-free environment. The microwave pyrolysis was conducted at the
temperature of 480 °C for 10 min, which could maximize the pyrolytic volatiles yield as
reported in our previous studies.29, 38, 39 The pyrolysis volatile vapors from the flask passed
through a packed bed catalysis reactor which was filled with catalyst. The packed-bed reactor
customized is constructed of quartz and externally heated by a heating tape. A thermocouple
was introduced between the reactor and the heating tape to measure catalytic temperatures.
As previous work reported, the optimal condition to maximize the composition of jet fuel
range aromatics was set at 500 °C.29 In this regard, the catalytic temperature was hold at
500 °C and catalyst (7.5 g) to biomass ratio was kept consistently at 0.3. The condensable
liquid was collected as bio-oil. The non-condensable vapors escaped as gas at the end of the
condensers and were collected for analysis. The catalytic microwave pyrolysis was duplicated
for two times in order to gain abundant bio-oils which were stepwise mixed together for the
hydrogenated process.
5.3.4 Hydrogenation of model compounds in diverse solvents
To evaluate the effect of diverse solvents on the hydrogenation of model compounds, a closed
reaction system with a stirred stainless batch reactor of the 4592 micro stirred reactor (with a
50mL vessel) and a 4848 reactor controller from Parr Instrument Company (Moline, IL, USA)
was used. In our previous study, naphthalene and its derivatives were the most abundant
compositions in the bio-oils from catalytic microwave pyrolysis of Doulgas fir sawdust
pellets;29 thus it is more appropriate to use naphthalene as the model compound for the
138
hydrogenation reaction. Naphthalene (1 g), Raney-Ni 4200 (0.1 g), solvent (7 g) were placed
in the reactor. Solvents used were methanol, 2-propanol, ethyl acetate, benzene, n-hexane,
cyclohexane, toluene, n-heptane, methylcyclohexane, decalin, and n-dodecane. Then the
reactor was sealed and vented for five times with hydrogen to get rid of the air present in the
vessel. Hydrogen was subsequently adjusted to reach the set pressure (500 psi). The
automatic controller was employed to control the temperature and the revolution of stirrer
(300 rpm). The pressure inside the reactor was recorded and the reactions proceeded at a set
temperature (200 °C) for 2 h. After the experiment finished, stirring was stopped and the
reactor was rapidly cooled to ambient temperature. Then, the gas was collected for analysis
and the reactor was depressurized. Consequently the liquid product was filtered to remove
catalyst particles.
5.3.5 Hydrotreatment of bio-oil derived from catalytic microwave pyrolysis
The combined bio-oils evolved from catalytic microwave pyrolysis were extracted by
the optimal solvent after determining the effects of all solvents in the hydrogenated
system. To produce the jet fuel range cycloalkanes, the mixture of organics and the
optimum solvent was loaded into the reactor together with 10 wt. % or 20 wt. %
catalysts (in terms of the reactants). The reactor was flushed with hydrogen and then
tightly closed. The experiments for each run were performed at 150, 200, or 250 °C
under stirring (300 rpm) for the intended time. The final step conducted was the same
as aforementioned hydrogenation of model compounds.
139
5.3.6 Analytical techniques
Elemental analysis (C, H, N, and O) of Douglas fir sawdust pellets, liquid samples,
char, and coke deposited on spent catalysts was conducted using a 2400 Series II
CHN/O Elemental Analyzer (PerkinElmer, USA).
The textural properties of the catalyst were determined by means of N2
adsorption–desorption (Micromeritics TriStar II 3020 Automatic Physisorption
Analyzer). Fresh smaples were degassed in vacuum at 300 °C for 1 h. The
Brunauer–Emmett–Teller equation was applied to calculate the specific surface area
using adsorption data at p/po= 0.05–0.25. The pore volume was evaluated by using the
Barrett–Joyner–Halenda (BJH) method.
Powder X-ray diffraction (XRD) patterns were executed on a Rigaku Smartlab X-ray
diffractometer equipped with a Cu Kα X-ray source operating at 40 kV and 40 mA.
The scattering angle 2θ was varied from 10° to 80°.
The particle size distribution and surface morphology of the samples were measured
with a scanning electron microscope (SEM, FEI Quanta 200 F).
The chemical composition of the bio-oils was characterized and qualified by Agilent
7890A GC-MS (GC–MS; GC, Agilent 7890A; MS, Agilent 5975C) with a DB-5
capillary column. The GC was first programmed to heat to 45°C for 3 min followed by
140
heating to 300°C at a rate of 10°C/min. The injection sample size was 1 μL. The flow
rate of the carrier gas (helium) was 0.6mL/min. The ion source temperature was 230°C
for the mass selective detector. Compounds were identified by comparing the spectral
data with that in the NIST Mass Spectral library. The area percent of changed
concentrations of model compounds obtained from GC/MS results was utilized to
predict product concentration in bio-oils.
The moisture content in the bio-oils was determined by a Karl Fischer (KF) compact
titrator (V20 Compact Volumetric KF Titrator, Mettler-Toledo).
The gaseous product was collected in a 1L Tedlar gas bag and then offline analyzed by an
INFICON 3000 Micro-GC (INFICON Inc., Santa Clara, CA, USA) system with a thermal
conductivity detector (TCD). A standard gas mixture consisting of H2, N2, CH4, CO, CO2,
C2H4, C2H6, and C3H6 was used to calibrate the yield of non-condensable gas. Alkanes and
olefins (>C4) in gas samples were either not detected or negligible in this research.
5.3.7 Data evaluation
The coke mass was determined by the difference before and after catalytic pyrolysis.
The weight of non-condensable gas was calculated using the following equation:
L
=C
−
− ℎ
141
−
(1)
Conversion of naphthalene, overall carbon yields of the liquid, gas, and solid products,
carbon yield of a specific product were calculated based on the following equations.
B
H
B
C
B
C
Eℎ ℎ
D
=
H D=
=
C
E
C
C
C
Eℎ ℎ
Eℎ ℎ
× 100% (2)
× 100%
E
E
(3)
× 100% (4)
Fig. 5.1 Overall carbon yield distribution (A) and main aromatics carbon selectivity
(B) for extracted bio-oils from catalytic microwave pyrolysis of Douglas fir sawdust
pellets
142
5.4 Results and discussion
5.4.1 Catalytic transformation of lignocellulosic biomass into C8- C16 aromatics
To produce jet fuel range cycloalkanes, the transformation of lignocellulosic biomass
into jet fuel range (C8 – C16) aromatics is primarily required. The water content of
bio-oils obtained was close to 20 wt% (based on biomass loading), which was derived
from both the moisture of raw biomass and the subsequent dehydration reaction during
biomass pyrolysis. The yield of organics (7.52 g) in the bio-oils accounted for 15.04
wt. %, prior to extracting process. The overall carbon yield distribution from catalytic
microwave pyrolysis of biomass is shown in Fig. 5.1(A). It is noteworthy that the char
were found to be the dominant composition, occupying 37.37%. It was also noted that
the carbon yield (34.57%) of non-condensable gas was calculated by difference. The
gas was mainly comprised of hydrogen, carbon dioxide, carbon monoxide, methane,
originating from decarboxylation, decarbonylation, cracking, and oligomerization
reactions. In addition, 24.68% carbon yield of aromatics was obtained after mixing all
bio-oils from catalytic microwave pyrolysis of Doulas fir sawdust pellets. In contrast,
the carbon yield (3.38%) of catalyst coke was much lower than other studies reported
elsewhere.27, 28, 40, 41
143
Table 5.1 Effect of various solvents on the hydrogenation reaction of naphthalene.a
+
Naphthalene
Decalin
Tetralin
Carbon selectivity (C mol%)
Entry
Solvent
Conv. (%)
Tetralin
Decalin
1
Methanol
22.6
15.1
7.5
2
2-propanol
99.9
52.9
47.0
3
Ethyl acetate
100
77.1
22.9
4
Benzene
95.2
86.1
9.1
5
n-hexane
99.9
7.1
92.8
6
Cyclohexane
99.9
13.4
86.5
7
Toluene
98.9
86.8
12.1
8
n-heptane
100
0.1
99.9
9
Methylcyclohexane 99.9
30.4
69.5
10
Decalin
100
0.5
99.5
11
n-dodecane
100
32.6
67.4
a
Reaction condition: naphthalene (0.0078 mol), solvent (7 g), Raney-Ni 4200 (0.1 g,
10 wt. % with respect to reactant mass), 500 psi H2 (initial pressure), 200 ºC for 2 h.
The carbon selectivity of main aromatics from catalytic microwave pyrolysis of
Douglas fir sawdust pellets is depicted in Fig 5.1(B). Generally the typical
compositions belong to the lumps of jet fuel range aromatic hydrocarbons, including
mono-cyclic and poly-cyclic aromatic hydrocarbons. Carbon selectivity towards
xylenes was 12.54%, which was the dominant share in the liquid organics. Other
mono-cyclic aromatic hydrocarbons, e.g. indene (5.06% carbon selectivity) and indane
(2.99% carbon selectivity), were also obtained. The large amount of poly-cyclic
144
aromatic hydrocarbons including naphthalene and its derivatives was evolved from
oligomerization reactions of mono-cyclic aromatic hydrocarbons.42,
43
Carbon
selectivity of naphthalene and 2-methylnaphthalene were 10.77% and 12.30%,
respectively. The results imply that total amounts of poly-cyclic aromatic
hydrocarbons were the most abundant compounds. Other poly-cyclic aromatic
hydrocarbons with three rings had low carbon selectivity, most of which
corresponding to carbon number did not exceed C16.
5.4.2 Solvents effect on the hydrogenation of naphthalene
Aiming to ultimately determine the solvents effect on the hydrogenation of
aromatics-enriched bio-oils evolved from catalytic microwave pyrolysis, we first
explored the influence of various solvents in the hydrogenation reaction of a model
compound (naphthalene). Table 5.1 summarize the results from hydrogenation of
naphthalene with Raney-Ni 4200 performed in several solvents under 500 psi H2
(initial pressure) at 200 °C for 2 h. As for all experiments, only unreacted hydrogen
was detected at the end of reactions, implying that the reactions were not carried out
under hydrogen starved conditions. The solvents effect on the hydrogenation of
naphthalene demonstrates that the reaction medium could promote the catalytic
transformation. The experiment conducted in methanol achieved only 22.6%
conversion of naphthalene (entry 1). Trace amount of decalin (7.5% toward carbon
selectivity) was detected in this solvent. The results imply that Raney-Ni 4200 could
be poisoned when hydrogenation was conducted in neat methanol, which is identical
145
to the study reported elsewhere.35 Unlike the inhibited effect of methanol, both
2-propanol and ethyl acetate led to almost full conversion of the substrate. 47.0%
carbon selectivity of decalin in 2-propanol medium was obtained from the
hydrogenation reaction of naphthalene, which is remarkably superior to 22.90%
carbon selectivity conducted in the medium of ethyl acetate. It is well developed that
2-propanol could serve as H-donor in the hydrogen transfer during hydrogenation.44
Hence, 2-propanol as the reaction medium could facilitate the hydrogenation reaction
of tetralin to decalin.
Among the solvents of aromatic hydrocarbon (benzene and toluene), over 95%
naphthalene was converted; however poor conversions were gained in decalin (9.1%
and 12.1%, respectively). It was probably attributed to the unsaturated solvents which
could first attach to the surface of Raney-Ni 4200 and then reacted with hydrogen
absorbed at the surface. The competitive reaction resulted in insufficient hydrogen and
active sites of catalyst for the hydrogenation reaction of tetralin to decalin. Conducting
the reaction in alkanes (entry 5-6, 8-11), however, contributed to very high to full
conversion:
n-hexane
(99.9%),
cyclohexane
(99.9%),
n-heptane
(100%),
methylcyclohexane (99.9%), decalin (100%), n-dodecane (100%). For maximizing the
hydrogenation of naphthalene to decalin, n-heptane showed the best conversion
(99.9% carbon selectivity), comparing with n-hexane (92.8%), cyclohexane (86.5%),
methylcyclohexane (69.5%), decalin (99.5%), n-dodecane (67.4%). As n-heptane does
not dissolve with water but display good extracting performance,35 therefore,
146
n-heptane is the desired solvent that should be used for extracting the
aromatics-enriched bio-oils. Moreover n-heptane is an essential compound in JP-4 and
Jet A,2 the hydrogenated mixture including cycloalkanes and n-heptane can be directly
utilized as jet fuels or drop-in fuels. Meanwhile, pure cycloalkanes can be separated
with n-heptane through vaporization. Accordingly, results revealed that the
hydrogenation reaction heavily depended on the solvents employed; n-heptane was the
optimal solvent, which was then used for extracting aromatics-enriched bio-oils and
hydrogenation reactions.
5.4.3 In suit hydrogenation of extracted bio-oils
5.4.3.1 The effect of catalyst loading on the chemical composition of hydrogenated
bio-oils
Given the water influence in the hydrogenation process, the mixed bio-oils from
catalytic microwave pyrolysis were extracted by the optimum solvent. 7.02 g of liquid
organics (~14 wt. %, based on biomass mass) were extracted using n-heptane (40 g).
The loss of organics (0.5 g) in water phase can be neglected if considering that the
recovery of organics could reach ~94 wt. %. Since the bio-oils produced by catalytic
microwave pyrolysis of intact biomass principally consisted of C8 - C16 range
aromatics, the controllable adjustment of aromatics with 8 - 16 carbon numbers are
considered as precursors of jet fuels. According to the hydrogenation reaction of
model compound in the medium of n-heptane, naphthalene of 99.9% carbon selectivity
was converted into saturated decalin. It is well known that decalin is the most
147
abundant cycloalkane in jet fuels.45 These conditions set a standard for designing the
following hydrogenation reaction of extracted bio-oils, thus the ratio of reactant to
solvent was set at 1:7. In this respect, additional solvent (9.14 g) was added into the
extracted bio-oils, maintaining the ratio at 1:7. Subsequently, the extracted bio-oils (C8
- C16 aromatics) were hydrogenated by several variables under low-severity
conditions.
The products distribution is summarized as a function of catalyst loading (10 wt. %
and 20 wt. % with respect to reactant mass) in Fig 5.2 (A). It was observed that
cycloalkanes as significant composition of jet fuels were in the range from 32.22 to
58.90% based on the GC/MS peak area, depending on alterations of catalyst loadings.
More than half of the aromatic intermediates were converted into jet fuel range
cycloalkanes using 20 wt. % Raney-Ni 4200, whose amount in hydrogenated bio-oils
was equal to that in JP-5. Unsaturated hydrocarbons (hydro-aromatic hydrocarbons
and cycloolefins) partially hydrogenated in the process accounted for 36.81 – 11.01%
and 9.04 – 10.31%, respectively. The variation of hydro-aromatic hydrocarbons (e.g.
tetralin) was significantly impacted by the catalyst loading. Adequate catalyst loading
could impel the hydro-aromatic hydrocarbons to be hydrogenated into cycloalkanes.
Catalyst loading also had a crucial influence on the selectivity of hydro-cyclic
alcohols, occupying from 4.74 to 11.74%, which suggests that phenols were
completely hydrogenated into hydro-cyclic alcohols under the condition of sufficient
catalyst loading. The amount of aromatic hydrocarbons decreased as the catalyst
148
loading was increased, implying that more active sites were offered for catalyzing the
aromatic intermediates to hydro-aromatic hydrocarbons or cycloalkanes. Other
oxygenated aromatics showed slight variations influenced by catalyst loading, varying
from 7.50 to 5.77%. These results indicate that hydrogenated reactions were
remarkably impacted by the catalyst loading.
The effect of catalyst loadings on partial cycloalkanes carbon selectivity is explained
in Fig. 5.2 (B). Catalyst loading did not have a vital influence on the carbon selectivity
of mono-cyclic cycloalkanes. These results indicate that mono-cyclic aromatic
hydrocarbons were prone to be hydrogenated into mono-cyclic alkanes even in the
presence of low catalyst loading. Two aromatic rings of naphthalenes were coplanar
and one aromatic ring was readily hydrogenated into a saturated ring; however the
aromatic ring and cyclic alkane ring of tetralin was no longer coplanar after
hydrogenation.31 The other aromatic ring was thus directed away from the catalyst
surface, resulting in lack of reduction. If more adequate catalysts were provided, the
other aromatic ring could also be hydrogenated at the spare active sites. It is obvious
that the carbon selectivity of both hexahydroindan (C9H16) and decalin (C10H18)
increased when more Raney-Ni 4200 catalyst were added in the reactor. In addition,
the carbon selectivity of bicyclohexyl (C12H22), perhydrofluorene (C13H22), and
tetradecahydroanthracene (C14H24) gradually increased as more catalysts were
employed. These results reveal that more catalyst loading, which means that more
149
active sites were offered on the Raney-Ni 4200 surface, gave rise to more
hydrogenated reactions.
Fig. 5.2 Overall products distribution (A) and patial aromatics carbon selectivity (B)
with resptect with catalyst loading. Reaction condition: Reaction temperature, 200 ºC;
Initial pressure, 500 psi; Reaction time, 2h.
150
Table 5.2 Products distribution and partial cycloalkanes carbon selectivity as a
function of reaction temperature.a
Temperature ( °C)
150
200
250
Cycloalkanes
47.43
58.90
63.34
Cycloolefins
9.48
10.31
5.20
Hydro-aromatic hydrocarbons
20.76
11.07
16.09
Hydro-cyclic
alcohols
8.34
11.74
3.78
Aromatic hydrocarbons
8.32
2.20
6.88
Other oxygenated aromatics
5.67
5.77
4.71
1,4-dimethylcyclohexane
1.06
1.18
1.35
1,3-dimethylcyclohexane
1.23
-
1.80
1,2-dimethylcyclohexane
-
0.53
0.40
Ethylcyclohexane
0.87
0.95
3.90
1,2,4-trimethylcyclohexane
0.23
0.29
0.35
Propylcyclohexane
0.68
0.79
1.71
Hexahydroindan
9.43
10.84
10.66
Decalin
12.67
15.13
15.63
Bicyclohexyl
0.67
0.71
1.14
Perhydrofluorene
1.04
1.13
-
Tetradecahydroanthracene
2.01
2.90
2.49
Overall selectivity (% in area)
Cycloalkanes selectivity (C mol%)
aReaction
condition: Initial pressure, 500 psi; Raney-Ni 4200, 20 wt. % with respect
to reactant mass; Reaction time, 2 h.
151
5.4.3.2 The effect of reaction temperature on the chemical composition of hydrogenated
bio-oils
Based on the previous result, catalyst loading of 20 wt. % was feasible for the
following experiments. In order to further understand chemical reactions in the process
and obtain more insight into hydrogenated bio-oils, the chemical compounds of
hydrogenated bio-oils are elucidated as a function of reaction temperature in Table 5.2.
It can be seen that the reaction temperature had a significant effect on the in situ
hydrogenation reactions. The amount of total cycloalkanes progressively increased
with the increasing of reaction temperatures. It was noticed that cycloalkanes were in
the range from 47.43% at 150 °C to 63.34% as reaction temperature increased to 250
°C, implying that a high reaction temperature favored the hydrogenation reactions.
Furthermore, there is a minor selectivity for cycloolefins, generally decreased from
9.48 to 5.20% alongside the increment of reaction temperatures. Therefore, elevated
temperature could enhance the hydrogenation reaction of olefins into cycloalkanes.
Hydro-aromatic hydrocarbons and aromatic hydrocarbons experienced declined
tendencies (20.76 – 11.07% and 8.32 to 2.20%, respectively) as the reaction
temperature went up to 200 °C; whereas the two compositions displayed an increased
trend as reaction temperatures increased from 200 °C to 250 °C. That is because the
elevated reaction temperature possibly favored the scission of hydroxyl group from
phenyl to form aromatic hydrocarbons through hydrogenolysis reactions.46 It was
found that the amount of hydro-cyclic alcohols was much lower at 250 °C than that at
200 °C. It can also be seen that other oxygenated aromatics was also declined to
152
4.71% at 250 °C. In the gas fraction, total amount (less than 1 vol. %) of small
hydrocarbons (such as methane, ethane, propane) were also detected at 250 °C, which
could be produced from hydrocracking of lager hydrocarbons. These results suggest
that hydrogenolysis and hydrocracking except hydrogenation reactions could take
place at high reaction temperatures.
Reaction temperatures also had a significant influence on the carbon selectivity of
specific cycloalkanes. It was noteworthy that mono-cyclic alkanes significantly
increased as the reaction temperature increased, especially from 200 to 250 °C.
Meanwhile, ethylcyclohexane (C8H16) dramatically increased from 0.95 to 3.90%,
which suggests that high reaction temperatures improved hydrocracking reactions of
poly-cyclic
alkanes
after
being
hydrogenated.
The
carbon
selectivity
of
hexahydroindan was slightly influenced by the changed reaction temperature when it
was set in the range of 200 - 250 °C. It showed that all indane and indene in the
extracted bio-oils were completely hydrogenated into hexahydroindan under these
conditions. Likewise, there were small upward tendency of decalin carbon selectivity,
which gradually increased from 12.67 to 15.63%. Carbon selectivity of other
poly-cyclic alkanes showed a steady increment, which implies that the hydrogenated
reactions of polycyclic aromatic intermediates could also be facilitated by increasing
reaction temperatures.
153
5.4.3.3 The effect of reaction time on the chemical composition of hydrogenated bio-oils
Apart from catalyst loading and reaction temperature, reaction time was another
crucial factor that influenced product distribution and cycloalkanes carbon selectivity
as shown in Table 5.3. Especially, the product distribution in the resulting product
prominently shifted towards the cycloalkane (39.62 – 71.92%) with increasing
reaction time, suggesting that hydrogenation reactions were enhanced because of
longer reaction duration. The amount of cycloolefins appeared to increase slightly to
11.15% at prolonged reaction time (3 h). The total amount of hydro-aromatic
hydrocarbons distantly declined, ranging from 29.24 to 0.51%. Likewise, the
selectivity of aromatic hydrocarbons experienced a downward trend from 9.67 to
1.82%. Prolonged reaction time could give rise to ring reduction of aromatic
hydrocarbons, which resulted in the decrease of aromatic hydrocarbons to form
hydro-aromatic hydrocarbons. Owing to more time provided for the reaction, the
production of cycloalkanes was likely facilitated by means of hydrogenating related
hydro-cyclic aromatic hydrocarbons. It was noted that the total amount of hydro-cyclic
alcohols was significantly impacted by prolonged reaction time especially going from
1 to 2 h.
154
Table 5.3 Products distribution and partial cycloalkanes carbon selectivity on the basis
of reaction time.a
Reaction time (h)
1
2
3
Cycloalkanes
39.62
58.90
71.92
Cycloolefins
9.27
10.31
11.15
Hydro-aromatic hydrocarbons
29.24
11.07
0.51
Hydro-cyclic
alcohols
5.22
11.74
10.06
Aromatic hydrocarbons
9.67
2.20
1.82
Other oxygenated aromatics
6.98
5.77
4.55
1,4-dimethylcyclohexane
1.02
1.18
-
1,3-dimethylcyclohexane
1.11
-
1.47
1,2-dimethylcyclohexane
0.23
0.53
-
Ethylcyclohexane
0.87
0.95
3.04
1,2,4-trimethylcyclohexane
0.34
0.29
1.00
Propylcyclohexane
0.76
0.79
0.94
Hexahydroindan
5.02
10.84
10.26
Decalin
14.15
15.13
14.51
Bicyclohexyl
0.64
0.71
0.77
Perhydrofluorene
0.67
1.13
0.77
Tetradecahydroanthracene
2.22
2.90
3.23
Overall selectivity (% in area)
Cycloalkanes selectivity (C mol%)
a Reaction
condition: Initial pressure, 500 psi; Raney-Ni 4200, 20 wt. % with respect to
reactant mass; Reaction temperature, 200
The
carbon
selectivity
for
°C.
mono-cyclic
alkanes
increased,
especially
for
ethylcyclohexane (0.87-3.04%), as the reaction time was elevated; whilst the carbon
155
selectivity of hexahydroindan and decalin were decreased from the period of 2 - 3 h.
Aromatic intermediates have already been converted into corresponding cycloalkanes
under the condition of 2 h. Prolonged time (3 h) could render the cycloalkanes to be
hydrocracked and oligomerized, forming mono-cyclic alkanes or derivatives. For the
gas result from the reaction of 3 h, trace volume of hydrocarbons was found,
suggesting that the hydrocracking and oligomerization reactions have taken place in
the process.
In contrast, polycyclic aromatic hydrocarbons with three rings are very
difficult to be hydrogenated due to their structures. Prolonging reaction time to until 3
h could make these polycyclic aromatic intermediates be totally hydrogenated, thereby
improving the carbon selectivity of poly-cyclic alkanes. Wherefore these result
indicate that prolonging reaction can principally enhance the hydrogenation reactions;
hydrocracking and oligomerization reactions may also occur jointly in the process
after 2 h reaction time.
5.4.3.4 The effect of catalysts selection on the chemical composition of hydrogenated
bio-oils
The chemical composition of hydrogenated bio-oils in the presence of various
catalysts (NP Ni, Raney-Ni 4200, and home-made Raney Ni) is depicted in Table 5.4.
Compared with the overall product distribution over Raney-Ni 4200 in the purchased
form, the result over NP Ni was similar under the same condition; while the result for
producing cycloalkanes in the presence of home-made Raney nickel was more
superior to the others. It was observed that 84.59% selectivity towards cycloalkanes
156
was achieved, which was even better than the result using Raney-Ni 4200 when the
reaction time was 3 h. The total amount (less than 5%) of cycloolefins was acquired in
the presence of home-made Raney Ni, which amount meets the specifications of
commercial jet fuels. Small amounts of hydro-aromatic hydrocarbons and aromatic
hydrocarbons were detected when using home-made Raney Ni, suggesting that
aromatic intermediates were thoroughly hydrogenated into cycloalkanes. As such, the
content of hydro-cyclic alcohols had the same tendency as hydro-aromatic
hydrocarbons and aromatic hydrocarbons, which indicates that partial hydro-cyclic
alcohols could be converted into cycloalkanes via the scission of hydroxyl group.
There was the lowest amount of other oxygenated aromatics in the hydrogenated
bio-oils using home-made Raney Ni as the catalyst, implying that hydrogenolysis
reactions also took place in the process. Small hydrocarbons detected in the gas
revealed that these hydrocarbons with small volume were from the hydrocracking and
hydrogenolysis of larger hydrocarbons.
Table 5.4 Products distribution
and partial cycloalkanes carbon selectivity in the
presence of various catalysts.a
Catalyst categories
NP Ni
Raney-Ni
Home-made
4200
Raney Ni
Overall selectivity (% in area)
Cycloalkanes
59.51
58.90
84.59
Cycloolefins
8.94
10.31
2.83
Hydro-aromatic hydrocarbons
12.33
11.07
0.60
157
Hydro-cyclic
alcohols
9.93
11.74
6.84
Aromatic hydrocarbons
3.09
2.20
0.99
Other oxygenated aromatics
6.20
5.77
4.15
1,4-dimethylcyclohexane
-
1.18
-
1,3-dimethylcyclohexane
3.80
-
4.10
1,2-dimethylcyclohexane
0.36
0.53
0.66
Ethylcyclohexane
0.99
0.95
4.18
1,2,4-trimethylcyclohexane
0.39
0.29
1.37
Propylcyclohexane
1.09
0.79
1.65
Hexahydroindan
11.05
10.84
8.09
Decalin
14.29
15.13
16.69
Bicyclohexyl
0.69
0.71
1.13
Perhydrofluorene
0.64
1.13
1.04
Tetradecahydroanthracene
2.47
2.90
3.12
Cycloalkanes selectivity (C mol%)
a
Reaction condition: Initial pressure, 500 psi; Catalyst, 20 wt. % with respect to
reactant mass; Reaction temperature, 200 °C ; Reaction time, 2 h.
It was found that the carbon selectivity of both dimethylcyclohexane and ethylcyclohexane
were maximum in the presence of home-made Raney Ni. Other mono-cyclic alkanes such as
propylcyclohexane and trimethylcyclohexane were also shown the largest carbon selectivity.
Therefore home-made Raney Ni was the optimal catalyst for the production of mono-cyclic
alkanes. Of the catalysts, home-made Raney Ni displayed the lowest carbon selectivity of
hexahydroindan. These results indicate that hexahydroindan formed was possibly
hydrocracked over the catalyst, thereby forming more mono-cyclic alkanes. Unlike the
hexahydroindan, the carbon selectivity for decalin, bicyclohexyl, tetradecahydroanthracene
was all the highest when using the home-made Raney Ni as the catalyst. These outcomes
158
imply that employing home-made Raney Ni as the catalyst could improve the overall
selectivity of cycloalkanes, and obtain the highest carbon selectivity of mono-cyclic alkanes.
Table 5.5 Textural properties of Ni-Al alloy, NP Ni, Raney-Ni 4200 and home-made
Raney Ni catalysts.a
SBET
Vpore
Spore
dpore
(m2/g)
(cm3/g)
(m2/g)
nm
Ni-Al alloy
0.35
0
0
0
NP Ni
157.7
0.048
38.2
5.0
Raney-Ni 4200
38.1
0.112
43.0
10.4
Home-made Raney Ni
52.4
0.034
35.8
3.8
a
SBET: BET surface area;
Vpore: pore volume;
Spore: pore surface area;
dpore:
average pore size.
5.4.4 Catalyst characterization
During alkali leaching, Al in the Ni-Al alloy powder reacted with NaOH solution. In
this respect, it is expected that Raney Ni is primarily comprised of metallic nickel.
When the Al/NaOH stoichiometry is more than 1, insoluble Al(OH)3 was also formed
by the reaction between Al and the water.47 Therefore NP-Ni produced with
insufficient amount of NaOH is regarded as a Ni-Al(OH)3 catalyst. However, there is
also small amount of residual hydrated alumina absorbed in the spongy structure of
Ranye Ni, as illuminated by Zhu et al.47 The textural properties of NP-Ni, Raney-Ni
4200 and home-made Raney Ni, compared with the parent Ni-Al alloy powder, are
illustrated in Table 5.5. BET surface area, pore volume, and pore surface area of all
catalysts were significantly improved by the alkali treatments. After dissolution of Al
159
component, the BET surface area of home-made Raney Ni dramatically increased
from 0.35 to 52.4 m2/g, which is higher than that of purchased Raney-Ni 4200 (38.1
m2/g) and other Raney Ni catalyst.48 The decent BET surface area of home-made
Raney Ni assisted the adsorption of hydrogen on the surface of the catalyst for the
hydrogenation reaction. The catalytic activity of NP-Ni was probably inhibited by the
large amount of hydrated alumina adsorbed in the structure, although the NP-Ni has
the highest BET surface area. Besides 0.034 cm3/g of pore volume and 35.8 m2/g of
pore surface area for home-made Raney Ni were also generated. The average pore size
of home-made Raney Ni is 3.8 nm. The suitable average size is close to the
mono-cyclic aromatics diameter, thus mono-cyclic aromatics are prone to be adsorbed
in the pores.49
Fig. 5.3 The XRD patterns of the Ni-Al alloy powder, NP-Ni, Raney-Ni 4200, and
home-made Raney Ni.
160
A
B
C
D
Fig. 5.4 SEM images of the Ni-Al alloy powder (A), NP-Ni (B), Raney-Ni 4200 (C),
and home-made Raney Ni (D).
Fig. 5.3 outlines the XRD patterns of the Ni-Al alloy, NP Ni, Raney-Ni 4200 and
home-made Raney Ni. The XRD patterns of Ni-Al alloy indicates that it consists of the
two kinds of Ni3Al2 and Ni3Al domains. After dissolving Al component by 20 wt. %
NaOH solution, the diffractions with respect to metallic Ni were observed as
amorphous nature for both Raney-Ni 4200 and home-made Raney Ni. They mainly
displayed diagnostic (111), (200), and (220) diffractions of fcc Ni at 2θ of 44.5, 51.8,
161
and 76.3°, respectively.50 With regarding to NP Ni, the large content of Al(OH)3 was
formed and identified according to the peaks assignable to gibbsite and bayerite.
The SEM images of Ni-Al alloy powder, NP-Ni, Raney-Ni 4200 and the home-made
Raney Ni are shown in Fig. 5.4. The morphological differences between Ni-Al alloy
powder and the other three catalysts are readily visible from Fig. 5.4 (A), (B), (C), and
(D). It was observed that both Raney-Ni 4200 and home-made Raney Ni are
constituted by the typical fractured and angular particles, which are in line with other
research;51 while Ni-Al alloy powder shows the intact metallic structure. The small
particles of home-made Raney Ni are more dispersive in comparison with Raney-Ni
4200, confirming a higher BET surface. For NP Ni, it is composed of nickel angular
particles and irregularly oriented aluminum crystal-like particles.47
5.4.5 Reaction pathway for the conversion of lignocellulosic biomass into jet fuel range
cycloalkanes
These observations are the key point to propose the reaction pathway for the
conversion of lignocellulosic biomass into jet fuel range cycloalkanes. Based on the
quantified products distribution in this study, and related results from direct
microwave pyrolysis of woody biomass and catalytic microwave pyrolysis of
cellulose,36, 38 the overall reactions network (including catalytic microwave pyrolysis
and hydrogen saving process) is summarized and shown in Fig. 5.5. In the primary
route from lignocellulosic biomass to jet fuel range cycloalkanes, cellulose underwent
162
a series of dehydration, decarboxylation, and decarbonylation to form furan
compounds during catalytic pyrolysis.36,
52
These furans then went through
decarbonylation, aromatization, and oligomerization reactions inside the pores of
well-promoted ZSM-5 to form aromatic hydrocarbons. Likewise, we observed that
hemicellulose was decomposed into furan compounds,38 which was identical to the
result using cellulose as feedstock.36 These furan compounds also diffused into the
well-promoted ZSM-5 pores, which underwent dehydration, decarboxylation,
decarbonylation, and oligomerization reactions to generate aromatic hydrocarbons.
Monomeric phenolic compounds were primarily generated from decomposition of
lignin in lignocellulosic biomass, which were thereafter catalyzed into aromatic
hydrocarbons through a series of dehydration, cracking, and oligomerization
reactions.28, 52
Single or multiple aromatic rings were transformed into single or multiple alkane rings by
hydrogenation reactions in the presence of Raney nickel catalysts. As a result, the achieved
aromatic hydrocarbons in the jet fuel range could be converted into jet fuel range
cycloalkanes from the hydrogen saving process under very mild conditions. Trace volume of
small hydrocarbons was also produced by hydrocracking reactions. The cycloalkanes
(together with minor aromatics and cycloolefins) can be used as additives to make the desired
jet fuels by blending with other hydrocarbons. The cycloalkanes produced in the hydrogen
saving process can be further upgraded through the hydrocracking and hydroisomerization
process to from both straight and branched alkanes.53, 54 The straight and branched alkanes
163
blend with cycloalkanes according to the proportions, which can produce different
commercial and military jet fuels.
H
H
Fig. 5.5 Proposed reaction pathways for the conversion of lignocellulosic biomass into
jet fuel range cycloalkanes.
5.5 Conclusions
164
This study demonstrated that the integrated processes were potentially profound
approaches when targeting production of jet fuel range cycloalkanes from
lignocellulosic biomass. The integral processes detailed here were illustrated to deliver
up to 84.59% selectivity of jet fuel range cycloalkanes from intact biomass under very
mild conditions. In the first step, the overall carbon yield of jet fuel range aromatics
from catalytic microwave pyrolysis was 24.68%. The bio-oils were extracted by
n-heptane and the recovery of organics could reach ~ 94 wt. %. These aromatics in
n-heptane medium were hydrogenated into jet fuel range cycloalkanes. Compared with
the conventional HDO process in the presence of noble catalyst, the process is a
hydrogen saving process, resulting in the costly hydrogen being transformed in the
form of cycloalkanes, rather than water. The home-made Raney nickel was the
optimum catalyst to obtain the highest selectivity of jet fuel range cycloalkanes. It is
more likely that these high-yield cycloalkanes mixing with minor aromatics and
cycloolefins can be potentially used as additives in jet fuels. From this perspective,
these integrated processes by using inexpensive catalysts and intact biomass under the
mild conditions deliver a novel and feasible route, specifically targeting jet fuel range
cycloalkanes. Future efforts direct towards improving the carbon yield of aromatics
from catalytic microwave pyrolysis and implement an economic analysis in a
biorefinery based on the integral process.
165
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169
CHAPTER SIX
DEVELOPMENT OF A CATALYTICALLY GREEN ROUTE FROM
DIVERSE LIGNOCELLULOSIC BIOMASS TO HIGH-DENSITY
CYCLOALKANES FOR JET FUELS
6.1 Abstract
This study reports a novel route to manufacture high-density cycloalkanes for jet fuels from
diverse lignocellulosic biomasses. The consecutive processes for manufacturing high-density
cycloalkanes primarily included the catalytic microwave-induced pyrolysis of diverse
lignocellulosic biomasses (hybrid poplar, loblolly pine, and Douglas fir) over a
well-promoted ZSM-5 and hydrogenation process in the present of the Raney nickel catalyst.
Two variables (catalytic temperature and catalyst to biomass ratio) were employed to
determine the optimal condition for the production of C8 – C16 aromatics in the catalytic
microwave pyrolysis. The maximum carbon yield of desired aromatics was 24.76%, which
was achieved from catalytic microwave pyrolysis of hybrid poplar at 500 °C with the catalyst
to biomass ratio of 0.25. We observed the aromatics derived from catalytic microwave
pyrolysis in the n-heptane medium were completely hydrogenated into renewable
high-density cycloalkanes for jet fuels. In the hydrogenation process, increasing catalyst
loading and reaction temperature could promote the selectivity of high-density cycloalkanes.
Results indicated that the hybrid poplar was the optimal feedstock to obtain the highest
selectivity (95.20%) towards high-density cycloalkanes. The maximum carbon yield of
170
cycloalkanes-enriched hyrogenated organics based on hybrid poplar was 22.11%. These
high-density cycloalkanes with high selectivity can be directly used as additives in jet fuels,
such as JP-5, JP-10, and RJ-5.
Keywords: High-density cycloalkanes, bio-jet fuels; diverse lignocellulosic biomasses;
hydrogenation; home-made Raney Ni catalyst
6.2 Introduction
There are immense interests in converting renewable resources and developing new
generation (hydrocarbon) biofuels, with a particular focus on green aviation fuels.1, 2
Lignocellulosic biomass is considered as a ubiquitous and sustainable source of carbon
that displays promising potential in the manufacture of hydrocarbon transportation
fuels and versatile chemicals.3, 4 The agricultural and forest resources reserves in U.S.
potentially satisfy one-third of the country’s petroleum demand.5 Additionally,
demand for diesel and jet fuels in the United States is expected to continue increasing
by 27% in the following years as opposed to gasoline 6; thus it is essential to shift
biofuel production derived from lignocellulosic biomass towards distillate-range liquid
alkanes in the future.
The current jet fuels originated from fossil resources are principally comprised of
linear-chain and branched-chain alkanes.7,
8
However, lower densities (~0.76-0.78
g/mL) of linear-chain alkanes and branched-chain alkanes have to be blended with
171
high-density hydrocarbons to meet the specifications of jet fuels.9 Conversely,
cycloalkanes (especially the polycyclic alkanes) are compact molecules within robust
ring strain and contribute to a more dense jet fuel and burn cleanly with high heats of
combustion and low freezing points. 10-12 To make up the shortage of chain alkanes, jet
fuel range cycloalkanes or aromatic hydrocarbons should be synthesized and added
into commercial jet fuels (e.g. Jet A and JP-8).8, 13 Therefore, most efforts to substitute
petroleum-based high-density tactical fuels, such as JP-5, JP-10 and RJ-5 have focused
on increasing polycyclic high-density hydrocarbons.
Several technologies have been evaluated for the production of renewable drop-in
fuels for aviation; Dumesic group14, 15 and Huber et al.7, 16 have pioneered a new route
to synthesize the long carbon chain alkanes using the lignocellulose-derived platform
chemicals. Nevertheless, these multiple steps necessitate expensive reactants and noble
metal
catalysts
under
severe
reaction
conditions,
resulting
in
infeasible
commercialization in biorefineries. In the current work for producing high-density
biofuels, Harvey group17 and Zou et al.18 have successfully developed liquid fuels with
a density of 0.94 g/mL via the dimerization of pinene followed by hydrogenation.
Unfortunately the sources of some specific woods and plants used for the production
of pinene are limited. Accordingly the production of high-density alkanes for jet fuels
using cheap and abundant sources call for ideal technologies with efficient solid-phase
catalysts to make the processes economically feasible.
172
It is widely known that oxygen content of bio-oil could be partially or even completely
eliminated by zeolite cracking during fast pyrolysis of biomass.19 Biomass-derived
feedstock has been directly converted into aromatics with ZSM-5 as the catalyst in a
single
catalytic
pyrolysis
step.20,
21
Besides,
Lei
group
have
found
that
microwave-induced pyrolysis could enhance the selectivity of aromatics in the
upgrading bio-oil.22-24 Yet, the evaluation of how the biomass-derived feedstock
affects the chemical composition (especially for aromatic hydrocarbons) of bio-oil
from catalytic pyrolysis and its impact as a precursor of cycloalkanes for jet fuels are
often overlooked.25 Wherefore understanding how biomass feedstock influences the
chemical composition of the resulting bio-oil is vital to obtain specific range aromatic
hydrocarbons for jet fuels.26 For lignocellulosic biomass sources, hybrid poplar,
loblolly pine, and Douglas fir have emerged as front-runners in U.S.5 These
fastest-growing and abundant biomasses in North America and are well suited for a
variety of applications such as biofuels production and bio-based chemicals, providing
flexibility for future biorefineries.
Furthermore, it was noticed that high density aromatics (C8 - C16) are prone to be
hydrogenated into cycloalkanes.27,
28
To this end, diverse lignocellulosic biomasses were
comparably converted into aromatics by catalytic microwave pyrolysis over well-promoted
ZSM-5. The organic phase of the bio-oils derived from catalytic microwave pyrolysis was
thereafter extracted by the optimum solvent. In the final step, the organic mixture was
hydrogenated into C8 – C16 high-density hydrocarbons (including cycloalkanes and minor
173
aromatics) by using Raney nickel as the catalysts, which satisfy basic specifications of
conventional jet fuels.
6.3 Experimental
6.3.1 Materials
The feedstock used was Douglas fir sawdust pellets (Bear Mountain Forest Products
Inc., USA) which were approximately 7 mm in diameter and 15 mm in length. Hybrid
poplar and loblolly pine samples were supplied by the Composite Materials and
Engineering Center (CMEC) at Washington State University. Parent ZSM-5
(SiO2/Al2O3 Mole Ratio: 50) was purchased from Zeolyst International, USA.
Nickel-Aluminum alloy powder in a non-activated type was used as purchased from
Alfa Aesar (Ward Hill, MA, USA).
6.3.2 Catalyst preparation
The activity of parent ZSM-5 was improved by suffering both hydrothermal and
calcined treatments. Under the gentle stirring, parent ZSM-5 powder was added into
deionized water (mass ratio=1) at 60 °C. After addition, the mixture was kept on
stirring for 2 h under this condition. The slurry was then dried at 105 °C till constant
weight. The sequential process was the catalyst calcination: hydrothermally treated
ZSM-5 was calcined at 550°C for 5 h in a muffle furnace. The catalysts were
pelletized and sieved to 20 – 40 mesh. The main characteristics of the catalyst were
reported in our previous study.23, 24
174
The Raney Ni catalyst was developed using a 20 wt% NaOH aqueous solution to remove Al
in the following procedure. 10 g of the above Ni-Al was slowly added into 100 mL NaOH
aqueous solution under gentle stirring. After addition, the slurry was kept on stirring at 80 °C
for 1 h. The excess of sodium hydroxide was finally washed with distilled water until nearly
neutral pH was reached. The obtained Raney Ni catalyst was stored in water. Raney Ni is
notorious for its pyrophoricity, and it may ignite spontaneously when dried in air. The Raney
Ni catalyst was thus dried at 60 °C till constant weight in the atmosphere of nitrogen to avoid
contact with air, prior to the subsequent catalytic test.
6.3.3 Catalytic microwave-induced pyrolysis of diverse ligonocellulosic biomasses
Detailed experimental setting was described in our previous studies.24, 29 The feedstock
was air dried at 105 °C for 24 h to remove the physically bound moisture, prior to
conducting the experiments. Fixed loading (20 g) of hybrid poplar, loblolly pine, and
Douglas fir for each run were placed in a 500 mL quartz flask which was placed inside
a Sineo MAS-II batch microwave oven (Shanghai, China) by a constant microwave
power setting (700 W). 0.05 g of activated carbon powder was used as the absorber for
the microwave-assisted pyrolysis. All reactions of microwave pyrolysis were
conducted at the temperature of 480 °C for 10 min. The pyrolysis volatile vapors from
the flask passed through a packed bed catalysis reactor which was filled with
well-promoted ZSM-5. The packed-bed reactor customized was constructed of quartz
and externally heated by a heating tape. A thermocouple was introduced between the
175
reactor and the heating tape to measure catalytic temperatures. The condensable liquid
was collected as bio-oils. The non-condensable vapors escaped as gas at the end of the
condensers and were collected for analysis.
6.3.4 Hydrotreatment of liquid organics derived from catalytic microwave
pyrolysis
Based on the solvents effects in the hydrogenation of a model compound
(naphthalene) in our previous study,27 the combined bio-oils evolved from catalytic
microwave pyrolysis were extracted by the optimal solvent (n-heptane). To produce
the high-density cycloalkanes for jet fuels, a closed reaction system with a stirred
stainless batch reactor of the 4592 micro stirred reactor (with a 50 mL vessel) and a
4848 reactor controller from Parr Instrument Company (Moline, IL, USA) was used.
The mixture of organics and the n-heptane medium was loaded into the reactor
together with 10 wt% or 20 wt% Raney Ni catalyst (in terms of the reactants). Then
the reactor was sealed and vented for five times with hydrogen to get rid of the air
present in the vessel. Hydrogen was subsequently adjusted to reach the set pressure
(500 psi). The automatic controller was employed to control the temperature and the
revolution of stirrer (300 rpm). The pressure inside the reactor was recorded and the
reactions proceeded at a set temperature (see Table 6.3) for the intended time. After
the experiment finished, stirring was stopped and the reactor was rapidly cooled to
ambient temperature. Then, the gas was collected for analysis and the reactor was
176
depressurized. Consequently the liquid product was filtered to remove catalyst
particles.
6.3.5 Analytical techniques
Elemental analysis (C, H and N) of diverse biomass sources, liquid samples, char, and
coke deposited on spent catalysts was conducted using a 2400 Series II CHN/O
Elemental Analyzer (PerkinElmer, USA).
The textural properties of the catalyst were determined by means of N2
adsorption–desorption (Micromeritics TriStar II 3020 Automatic Physisorption
Analyzer). Fresh catalysts were degassed in vacuum at 300 °C for 1 h. The
Brunauer–Emmett–Teller equation was applied to calculate the specific surface area
using adsorption data at p/po= 0.05–0.25. The pore volume was evaluated by using the
Barrett–Joyner–Halenda (BJH) method.
Powder X-ray diffraction (XRD) patterns were executed on a Rigaku Smartlab X-ray
diffractometer equipped with a Cu Kα X-ray source operating at 40 kV and 40 mA.
The scattering angle 2θ was varied from 10° to 80°.
The particle size distribution and surface morphology of the samples were measured
with a scanning electron microscope (SEM, FEI Quanta 200 F).
177
The chemical composition of the bio-oils was characterized and qualified by Agilent
7890A GC-MS (GC–MS; GC, Agilent 7890A; MS, Agilent 5975C) with a DB-5
capillary column. The GC was first programmed to heat to 45°C for 3 min followed by
heating to 300°C at a rate of 10°C/min. The injection sample size was 1 μL. The flow
rate of the carrier gas (helium) was 0.6mL/min. The ion source temperature was 230°C
for the mass selective detector. Compounds were identified by comparing the spectral
data with that in the NIST Mass Spectral library. The area percent of changed
concentrations of model compounds obtained from GC/MS results was utilized to
predict product concentration in bio-oils.
The moisture content in the bio-oils was determined by a Karl Fischer (KF) compact
titrator (V20 Compact Volumetric KF Titrator, Mettler-Toledo).
The gaseous product was collected in a 1L Tedlar gas bag and then offline analyzed by
an INFICON 3000 Micro-GC (INFICON Inc., Santa Clara, CA, USA) system with a
thermal conductivity detector (TCD). A standard gas mixture consisting of H2, N2,
CH4, CO, CO2, C2H4, C2H6, and C3H6 was used to calibrate the yield of
non-condensable gas. Alkanes and olefins (>C4) in gas samples were either not
detected or negligible in this research.
6.3.6 Experimental methods and data evaluation
178
A central composite experimental design (CCD) was employed to optimize the process
conditions and liquid organics yield from hybrid poplar, loblolly pine, Douglas fir.
The catalytic temperature (X1, °C) and catalyst to biomass ratio (X2) were chosen as
independent variables. In these experiments based on CCD, the mass of catalyst varied
from 2.2 to 7.8 g, while packed bed catalytic temperature ranged from 269 to 481 °C.
Two additional experiments conducted at 500 °C or with the catalyst loading of 1 g
were compared.
The coke mass was determined by the difference before and after catalytic pyrolysis.
After the experiment of microwave-induced pyrolysis, the char was left in the quartz
flask of pyrolysis reactor; while the coke was formed on the ZSM-5 catalyst in the
packed-bed catalysis reactor. The weight of non-condensable gas was calculated using
the following equation:
L
=C
−
− ℎ
−
(1)
Overall carbon yields of the liquid, gas, and solid products and carbon selectivity of a
specific product were calculated based on the following equations.
B
C
B
C
D
=
H D=
C
E
C
C
C
179
× 100%
E
E
(3)
× 100% (4)
Table 6.1 Textural properties of Ni-Al alloy powder and home-made Raney Ni catalyst.a
SBET
Vpore
Spore
dpore
(m2/g)
(cm3/g)
(m2/g)
nm
Ni-Al alloy powder
0.35
0
0
0
Home-made Raney Ni
52.4
0.034
35.8
3.8
aS
BET:
BET surface area; Vpore: pore volume;
Spore: pore surface area;
dpore: average
pore size
6.4 Results and discussion
6.4.1 Catalyst characterization of as-prepared Raney nickel
During alkali leaching, Al in the Ni-Al alloy powder reacted with NaOH solution. As
expected, the Raney Ni catalyst is primarily composed of metallic nickel. The textural
properties of as-prepared Raney Ni catalyst and the parent Ni-Al alloy powder were
demonstrated in Table 6.1. After the alkali treatment, BET surface area, pore volume,
and pore surface area were significantly improved. The BET surface area of the Raney
Ni catalyst dramatically increased by dissolution of Al component from 0.35 to 52.4
m2/g. Moreover, pore volume and pore surface area progressively went up to 0.034
cm3/g and of 35.8 m2/g, respectively. The decent BET surface area of the Raney Ni
catalyst assisted the adsorption of hydrogen on the surface of the catalyst for the
hydrogenation reaction. The average pore size of the Raney Ni catalyst is 3.8 nm,
which is close to the mono-cyclic aromatics diameter, thus mono-cyclic aromatics are
prone to be adsorbed in the pores.30
180
Fig. 6.1 outlines the XRD patterns of the Ni-Al alloy powder and the as-prepared
Raney Ni catalyst. The XRD patterns of Ni-Al alloy shows that it contains Ni3Al2 and
Ni3Al domains. After dissolving Al component by NaOH solution, the diffractions of
metallic Ni were observed as amorphous nature. It mainly displayed diagnostic (111),
(200), and (220) diffractions of fcc Ni at 2θ of 44.5, 51.8, and 76.3°, respectively.
Fig. 6.1 The XRD patterns of the Ni-Al alloy powder and home-made Raney Ni
catalyst.
The SEM images of Ni-Al alloy powder and the Raney Ni catalyst are illustrated in
Fig. 6.2. The morphological differences between Ni-Al alloy powder and the
home-made Raney Ni catalyst are readily visible from Fig. 6.2 (A) and (B). According
to SEM images of Ni-Al alloy powder and home-made Raney Ni catalyst, the small
particles of home-made Raney Ni are more dispersive, confirming a high BET surface.
It was also observed that the Raney Ni catalyst was constituted by the typical fractured
181
and angular particles, which are in line with other research;31 while Ni-Al alloy
powder shows the intact metallic structure. Therefore, the Raney Ni catalyst treated by
20 wt% NaOH solution could provide the excellent catalytic activity for the
hydrogenation reaction.
B
A
Fig. 6.2 SEM images of the Ni-Al alloy powder (A) and the as-prepared Raney Ni
catalyst (B).
6.4.2 Optimization of liquid organics from catalytic microwave pyrolysis
6.4.2.1 Product yield distributions from diverse lignocellulosic biomasses
The product yield distribution from hybrid poplar, loblolly pine, and Douglas fir are
summarized as a function of catalytic temperature and catalyst to biomass ratio.
It
was found that the yields of liquid organics and non-condensable gas were in the range
from 14.44 to 22.07 wt% versus 39.20 to 50.34 wt%, 13.56 to 21.02 wt% versus 40.98
to 52.02 wt%, 13.88 to 20.88 wt% versus 38.18 to 49.66 wt%, respectively. The
182
carbon yields of the liquid organics were compared in the following section. These
results indicate that catalytic temperature had significant effects on both yields of
liquids and gas. The yields of liquid organics declined with the increase of the catalytic
temperature; while the yield of non-condensable gas increased with the similar
increasing tendency of catalytic temperatures. These outcomes indicate that the
elevated catalytic temperature could facilitate catalytic cracking rate of large molecule
compounds into small molecules, thereby reducing the liquid organics yields.
Due to
the scissions of the C-O and C-C bonds at high catalytic temperature, the yields of
gases fractions (CO, CO2, and small hydrocarbons) were enhanced. The yields of
liquid organics were also visualized in accordance to the response surface and contour
line in Fig. 6.3 (A) for hybrid poplar, (B) loblolly pine, and (C) for Douglas fir. For
ANOVA analysis, the P-value of the three models are all less than 0.05, suggesting the
models are significant to present the relationships between product yields and reaction
conditions. The coefficient of determination (R2) for the three models are all larger
than 0.95, evidencing that these models fairly represent the relationships with regard to
the independent variables. Accordingly, these models can be used to predict the
maximum mass yields of liquid organics according to the two variables. It is worth
noting that the product yields were also affected by the catalyst to biomass ratio.
Increasing the ratio contributed to improving the yield of gas, whereas the yield of
liquid organics gradually decreased. It was found that when more catalysts were
loaded; more spare active sites would be provided for catalytic cracking reactions in
the fixed time, resulting in the varied yields of the products. Of the three biomasses,
183
the optimal biomass species for maximizing the organic yields was hybrid poplar.
When pyrolysis without catalyst; the liquid organic yield of hybrid poplar went up to
22.07 wt%, exceeding those from loblolly pine and Douglas fir. Similar phenomenon
was observed by other studies as a result of catalytic temperature and catalyst to
biomass ratio affecting product yields.23, 29
A
B
184
C
Fig. 6.3 Effect of the interaction of the independent variables on organic yield from
hybrid poplar (A), loblolly pine (B) and Douglas fir (C) yields (red pots represent
actual experimental values).
Solid carbonaceous residue (char and coke) from pyrolytic and catalytic processes can
be distinguished due to the ex-situ catalysis. The char yield evolved from diverse
biomass species (hybrid poplar, loblolly pine, Douglas fir) remained virtually constant
at 22.50 wt%, 20.20 wt%, and 21.50 wt%, respectively. It was claimed that high ash
and lignin contents could incur the high formation of solid residues including char and
coke in the pyrolytic process.26, 32 Hence high ash content in hybrid poplar33 and high
lignin content in Douglas fir22 caused the higher formation of char. Given the coke
deposited on the catalyst leading to deactivation of the active sites and micropores
blockage, the carbonaceous compounds were a crucial element to be taken into
consideration in the study of formation mechanism when zeolites are employed as
catalysts. The coke deposition yields of hybrid poplar, loblolly pine, and Douglas fir as
a function of catalytic temperature and catalyst to biomass ratio varied from 1.27 to
185
3.57 wt.%, 1.18 to 3.99 wt%, 1.95 to 4.02 wt%, respectively. The less formation of
coke from hybrid poplar was mainly attributed to less lignin content,33 comparing to
higher contents in loblolly pine and Douglas fir.
6.4.2.2 Chemical composition of liquid organics from catalytic microwave
pyrolysis of diverse biomass sources
As previous work reported, catalyst temperature had a significant effect on the
chemical composition of bio-oils by using ZSM-5 as the catalyst.23, 29 In this sense,
catalytic temperature was chosen as the variable to investigate and even optimize the
quality of liquid organics. In comparison with the liquid organics obtained by
non-catalytic pyrolysis of hybrid poplar, loblolly pine, and Douglas fir, the selectivity
towards aromatic hydrocarbons significantly increased, which was attributed to
microporous structure of ZSM-5 catalyst and active sites generated in the micropores.
It was found that the selectivity of aromatic hydrocarbons ranged from 17.95 to
94.33%, 23.75 to 93.16%, 2.72 to 92.62% as catalytic temperature increased to 500
°C, implying that the elevated catalytic temperature accelerated reaction rates of
cracking, dehydration, decarbonylation, decarboxylation, and aromatization in the
short time, giving rise to high selectivity towards aromatic hydrocarbons. It was
reported that cellulose as the feedstock for catalytic pyrolysis produced higher amount
of aromatic hydrocarbons than hemicellulose and lignin.34 It could also be inferred
from these results that higher amount of aromatic hydrocarbons from catalytic
pyrolysis of hybrid poplar was due to higher cellulose content.
186
Phenols and guaiacols were the most abundant compounds in raw organics without
using the catalyst, which was mainly derived from the direct decomposition of lignin;
whereas these compounds dramatically declined at 500 °C, and guaiacols were even
not detected from all biomass sources. These results indicate that high catalytic
temperature favored the scission of methoxy and hydroxyl groups from phenyl to form
aromatic hydrocarbons. Other minor oxygenated aromatics could be ignored when
catalytic temperature from the three biomass sources were set at 500 °C.
Scheme 6.1 Proposed reaction pathways for the conversion of lignocellulosic biomass
into high-density cycloalkanes for jet fuels
6.4.2.3 Catalytic transformation of lignocellulosic biomass into C8 - C16 aromatics
It was assured that the optimal condition to maximize the amount of aromatics was set
at 500 °C. To produce jet fuel range cycloalkanes, the novel synthetic route was
187
proposed in Scheme 6.1. Therefore, the transformation of lignocellulosic biomass into
C8 – C16 aromatics is primarily required. In this regard, the catalytic temperature was
held at 500 °C and catalyst to biomass ratio was kept constant at 0.25. The catalytic
microwave pyrolysis was duplicated for three times for all biomass species in order to
gain abundant liquid organics which were stepwise mixed together for the
hydrogenation process. The overall carbon yield distribution and partial aromatics
carbon selectivity from catalytic microwave pyrolysis of diverse biomass species are
shown in Table 6.2. It was noteworthy that the non-condensable gas and char were
found to be the dominant compositions from all biomasses. The non-condensable gas
was mainly comprised of hydrogen, carbon dioxide, carbon monoxide, methane.
Table 6.2 Products distribution
and main aromatics carbon selectivity from diverse
lignocellulosic biomasses at 500 °C.a
Biomass sources
Hybrid poplar
Loblolly pine
Douglas fir
Overall carbon selectivity (C mol%)
Gas
42.34
43.08
40.53
Char
30.63
31.73
34.29
Coke
2.27
2.28
3.66
Aromatics
24.76
22.91
21.52
Aromatics carbon selectivity (C mol%)
Toluene
4.77
5.18
4.32
Ethylbenzene
2.60
2.39
1.79
p-xylene/m-xylene
13.34
12.03
12.54
Trimethylbenzene
5.22
4.05
4.37
188
Indane
2.49
2.08
2.99
Indane
5.86
4.05
5.06
Phenol
1.34
1.86
1.52
p-cresol/m-cresol
1.59
1.55
2.03
Naphthalene
17.83
13.59
10.77
1-methylnaphthalene
2.87
3.72
1.68
2-methylnaphthalene
14.20
16.70
12.30
Anthracene
0.78
1.59
1.25
Pyrene
1.20
0.26
2.40
a Reaction
condition: Catalyst, 25 wt% with respect to biomass; Reaction temperature,
480 °C; Reaction time, 10 min.
In addition, It was also noted that 24.76% carbon yield of aromatics was obtained from
catalytic microwave pyrolysis of hybrid poplar, which was higher than those of
loblolly pine (22.91%) and Douglas fir (21.92%).
In contrast, the carbon yields of
catalyst coke were all much lower than other studies reported elsewhere.20, 21, 35, 36
Since the ex-situ catalysis could make the indirect contact between biomass and
catalyst, the catalyst could not be poisoned by the inorganic elements in the biomass
and solid carbonaceous char could not deposit on the catalyst, contributing to the low
carbonaceous residue formed on the catalyst. In general, hybrid poplar could be
considered as the better resource for the production of aromatics.
The main aromatics carbon selectivity from catalytic microwave pyrolysis of diverse
biomass sources is also depicted in Table 2. Generally the typical compositions belong
to the lumps of C8 – C16 aromatics, including mono-cyclic and poly-cyclic aromatic
189
hydrocarbons. Carbon selectivity towards xylenes from all biomass sources was
comparable between 12 and 14%. Other mono-cyclic aromatic hydrocarbons among
the three biomasses, e.g. toluene, phenol, indane, and indene, were also comparable.
Two-ring aromatic hydrocarbons were the dominant compositions including
naphthalene and its derivatives, originating from oligomerization reactions of
mono-cyclic aromatic hydrocarbons. As the pore diameter of the well-promoted
ZSM-5 catalyst is similar to the diameters of two-ring aromatic hydrocarbons,
two-ring aromatic hydrocarbons formed could be prone to diffusing out. Carbon
selectivity of naphthalene and 2-methylnaphthalene from hybrid poplar were 17.83%
and 14.20 %, respectively. The carbon selectivity was comparable with those of
loblolly pine, yet a little higher than that of Doulas fir (10.77% and 12.30%,
respectively). Other poly-cyclic aromatic hydrocarbons with three rings all had low
carbon selectivity, most of which corresponding to carbon number did not exceed C16.
6.4.3 Hydrotreatment of extracted organics
6.4.3.1 The effect of catalytic temperature on the chemical composition of
hydrogenated organics
Given the water influence in the hydrogenation process, the bio-oils from the catalytic
pyrolysis step were extracted by the optimum solvent (n-heptane). The liquid organics
together with n-heptane from each biomass were separated and weighed to measure
the loss of organics in the extracting step. The loss of organics in water phase can be
neglected if considering that all recoveries of organics exacted by n-heptane could
190
reach ~94 wt%, most of which were aromatic hydrocarbons. Since the liquid organics
produced by catalytic microwave pyrolysis of intact biomasses principally consisted of
C8 - C16 aromatics, the controllable adjustment of aromatics with 8 - 16 carbon
numbers are considered as precursors of jet fuels. According to hydrogenation of the
model compound in the medium of n-heptane, naphthalene with 99% carbon
selectivity was transformed into saturated decalin.27 Thus the mass ratio of reactant to
solvent was set at 1:7 in light of hydrogenation reaction of the model compound. The
extracted organics within C8 - C16 aromatics were stepwise hydrogenated in terms of
several variables.
In order to further understand the effect of the Raney Ni catalyst on the chemical
reactions and obtain more insight into hydrogenated organics, the chemical
compounds of hydrogenated organics from Douglas fir are elucidated as a function of
reaction temperature in Table 6.3. The selection of Douglas fir is because Douglas fir
was the typical feedstock, obtaining appropriate ratio mono-ring and poly-ring
aromatic hydrocarbons from catalytic microwave pyrolysis, which can be more readily
investigated into the conversions of the aromatic hydrocarbons.
It is noticeable that
the reaction temperature significantly influenced the hydrogenation reactions.
Cycloalkanes was formed through hydrocycloaddition of corresponding aromatic
hydrocarbons or other ring-unsaturated hydrocarbons. The gross amount of
cycloalkanes progressively increased with the increment of reaction temperatures. It
can be seen that cycloalkanes were in the range from 54.82% at 150 °C to 83.34% as
191
reaction temperatures went up to 250 °C, suggesting that increasing the reaction
temperature could facilitate the reaction rate of hydrogenation, resulting in
enhancement of cycloalkanes during the fixed reaction time. In this regard, more than
half of the aromatic intermediates were converted into jet fuel range cycloalkanes,
whose amount in hydrogenated bio-oils was equal to that in JP-5 37. Moreover, there is
a minor selectivity for cycloolefins, generally declining from 8.32 to 3.20% along with
the increasing of reaction temperatures. Therefore, elevated temperature could
facilitate the hydrogenation of cyclic olefins into cycloalkanes. Hydro-aromatic
hydrocarbons and aromatic hydrocarbons show decreased trend (20.07 to 6.32% and
8.19 to 1.88%, respectively). It was found that the amount of hydro-cyclic alcohols
was much lower at 250 °C than that at 200 °C, and other oxygenated aromatics was
also declined at 250 °C. These outcomes imply that the higher reaction temperature
possibly favored the scission of hydroxyl group from phenyl to form aromatic
hydrocarbons through hydrogenolysis reactions.38 These aromatic hydrocarbons could
be further transformed into ring-saturated hydrocarbons. In addition, less than 1 vol%
small hydrocarbons (such as methane, ethane, and propane) were produced at 250 °C,
which
were
possibly
from
hydrocracking
of
larger
hydrocarbons.
Hence,
hydrogenolysis and hydrocracking reactions could simultaneously occur at high
reaction temperatures in the hydrogenation system.
Reaction temperatures also had a significant effect on the carbon selectivity of specific
cycloalkanes. It was noteworthy that the carbon selectivity of mono-cyclic alkanes
192
significantly increased as the reaction temperature increased, especially from 150 to
200 °C. Meanwhile, 1, 3-dimethylcyclohexane (C8H16) and ethylcyclohexane (C8H16)
gradually increased from 2.27 to 5.04% and 2.34 to 4.67%, which indicates that high
reaction temperatures presumably accelerated hydrocracking reactions of poly-cyclic
aromatic hydrocarbons after being partially hydrogenated. Because aromatic and
alkane ring in hydro-aromatic hydrocarbons are not coplanar, the hydrogenated alkane
ring could be continuously hydrocracked on the Raney Ni catalyst and finally opened
to form branched chains attached on the aromatic ring. The carbon selectivity of
hexahydroindan (C9H16) was slightly impacted by alterations of reaction temperature.
Likewise, there were small upward tendency of decalin (C10H18) carbon selectivity,
which gradually increased from 16.03 to 16.28% when it was set in the range of
200-250 °C. It showed that all naphthalene was completely hydrogenated into decalin
under these conditions. Carbon selectivity of other poly-cyclic alkanes showed a
steady increment, implying that the hydrogenation rate of polycyclic aromatic
intermediates could also be facilitated by increasing reaction temperatures.
193
Table 6.3 Products distribution
and partial cycloalkanes carbon selectivity from
Douglas fir as a function of reaction temperature.a
Temperature ( °C)
150
200
250
Cycloalkanes
54.82
70.83
83.34
Cycloolefins
8.32
7.32
3.20
Hydro-aromatic hydrocarbons
20.07
10.22
6.32
Hydro-cyclic
alcohols
5.34
5.19
3.78
Aromatic hydrocarbons
8.19
3.20
1.88
Other oxygenated aromatics
3.26
3.24
1.48
1,4-dimethylcyclohexane
1.32
1.18
2.68
1,3-dimethylcyclohexane
2.27
4.54
5.04
1,2-dimethylcyclohexane
1.45
2.57
2.23
Ethylcyclohexane
2.34
3.07
4.67
1,2,4-trimethylcyclohexane
0.78
0.69
1.36
Propylcyclohexane
1.06
1.49
1.90
Hexahydroindan
9.86
10.40
11.64
Decalin
14.78
16.03
16.28
Bicyclohexyl
1.01
1.74
1.56
Perhydrofluorene
1.45
1.54
1.34
Tetradecahydroanthracene
2.67
2.89
3.02
Overall selectivity (% in area)
Cycloalkanes selectivity (C mol%)
aReaction
condition: Initial pressure, 500 psi; Raney Ni catalyst, 10 wt% with respect
to reactant mass; Reaction time, 2 h.
194
Fig. 6.4 Overall products distribution (A) and partial cycloalkanes carbon selectivity
(B) with respect with catalyst loading. Reaction condition: Reaction temperature, 200
°C; Initial pressure, 500 psi; Reaction time, 2h.
6.4.3.2 The effect of catalyst loading on the chemical composition of hydrogenated
organics
Yet the C8-C16 aromatics were not thoroughly converted into jet fuel range
cycloalkanes, even when the reaction temperature was carried out at 250 °C.
195
Other
effects such as catalyst loading were further investigated into the chemical
compositions of hydrogenated organics from Douglas fir. The products distribution is
summarized as a function of catalyst loading (10 wt% and 20 wt% with respect to
reactant mass) in Fig 6.4 (A). It was observed that the selectivity towards cycloalkanes
were 70.83% and 84.59% depending on alterations of catalyst loadings. The more
catalysts loaded into the reactor could provide more reactive sites for the
hydrocycloaddition of unsaturated hydrocarbons, thus improving the hydrogenated
efficiency. Unsaturated hydrocarbons (hydro-aromatic hydrocarbons and cycloolefins)
partially hydrogenated in the process accounted for 10.22 – 4.60% and 7.32 – 2.83%,
respectively. The variation of hydro-aromatic hydrocarbons (e.g. tetralin) was
significantly impacted by the catalyst loading. Adequate catalyst loading could impel
the hydro-aromatic hydrocarbons to be hydrogenated into cycloalkanes. The decreases
of the unsaturated hydrocarbons, as more catalysts were employed, were mainly due to
the transformations into corresponding cycloalkanes. Catalyst loading also had an
influence in the selectivity of hydro-cyclic alcohols, occupying from 5.19 to 3.84%.
The amount of aromatic hydrocarbons decreased with the increasing of catalyst
loading, implying that more active sites were provided for converting the aromatic
intermediates into hydro-aromatic hydrocarbons or cycloalkanes. Other oxygenated
aromatics showed slight variations, varying from 3.24 to 2.15%. These results indicate
that the chemical compositions of hydrogenated organics were essentially influenced
by the catalyst loading.
196
The effect of catalyst loadings on partial cycloalkanes carbon selectivity is explained
in Fig. 6.4 (B). Catalyst loading had a slight influence on the carbon selectivity of
mono-cyclic cycloalkanes, including dimethylcyclohexanes and ethylcyclohexane.
These results indicate that mono-cyclic aromatic hydrocarbons were prone to being
hydrogenated into mono-cyclic alkanes even in the presence of low catalyst loading.
Two aromatic rings of naphthalenes were coplanar and one aromatic ring was readily
converted into a saturated ring, however the aromatic ring and cyclic alkane ring of
tetralin was no longer coplanar after hydrogenation.39 Thus the opposite aromatic ring
was directed away from the catalyst surface. Once more adequate catalysts were
applied; the other aromatic ring was still hydrogenated at the spare active sites. It is
obvious that the carbon selectivity of both hexahydroindan and decalin increased since
more Raney Ni catalyst were introduced. In addition, the carbon selectivity of
bicyclohexyl (C12H22), perhydrofluorene (C13H22), and tetradecahydroanthracene
(C14H24) gradually increased as more catalysts were employed. Consequently, more
catalyst loading offered more active sites on the Raney Ni catalyst surface and inside
the structure, giving rise to more hydrocycloaddition reactions.
197
Table 6.4 Products distribution
and partial cycloalkanes carbon selectivity as a
function of diverse biomass sources.a
Biomass sources
Hybrid
Loblolly
poplar
pine
Douglas fir
Overall selectivity (% in area)
Cycloalkanes
95.20
94.06
92.70
Cycloolefins
1.01
0.89
1.20
Hydro-aromatic hydrocarbons
0.23
0.22
0.34
Hydro-cyclic
alcohols
2.67
3.12
3.56
Aromatic hydrocarbons
0.67
1.20
1.05
Other oxygenated aromatics
0.22
0.51
1.15
1,4-dimethylcyclohexane
4.23
5.48
1.52
1,3-dimethylcyclohexane
5.20
3.27
5.23
1,2-dimethylcyclohexane
2.34
1.53
2.11
Ethylcyclohexane
5.87
4.59
5.03
1,2,4-trimethylcyclohexane
3.06
2.29
1.36
Propylcyclohexane
1.32
0.88
1.79
Hexahydroindan
10.34
11.23
10.84
Decalin
16.76
16.34
17.05
Bicyclohexyl
1.23
1.65
1.32
Perhydrofluorene
1.89
2.05
1.34
Tetradecahydroanthracene
3.02
3.76
3.49
Cycloalkanes selectivity (C mol%)
aReaction
condition: Initial pressure, 500 psi; Raney Ni catalyst, 20 wt% with respect
to reactant mass;
Reaction temperature, 250
°C; Reaction time, 2 h.
6.4.3.3 The effect of diverse biomass sources on the chemical composition of
hydrogenated organics
198
Apart from reaction temperature and catalyst loading, biomass source was another
crucial factor that influenced product distribution and carbon selectivity of the
cycloalkanes. The chemical composition of hydrogenated organics from diverse
biomass sources (hybrid poplar, loblolly pine, Douglas fir) is depicted in Table 6.4.
After the hydrogenation process, the recovery of the hydrogenated organics from the
diverse biomass sources were all over 95 wt%. Compared with the overall product
distribution from loblolly pine and Douglas fir, the result for the production of jet fuel
range cycloalkanes from hybrid poplar was more superior to the others under the same
condition. As mentioned above, the higher compositions of cellulose in hybrid poplar
contributed to the shares of aromatics derive from catalytic microwave pyrolysis,
therefore giving rise to higher content of cycloalkanes via hydrocycloadditions of
aromatics. It was observed that 95.20% selectivity towards cycloalkanes was achieved
from hybrid poplar. From the optimal result of hybrid poplar, approximately 0.028
mole of hydrogen were consumed by the parent aromatics (1 g) by the hydrogenation
reactions. The minor amount of hydrogen consumed in the process was affirmed that
hydrogen preferentially acted as the H-donor to undergo the hydrocycloaddition of
ring-unsaturated
hydrocarbons
to
generate
cycloalkanes,
rather
than
the
hydrodeoxygenation of oxygenates to form water. Because of the high content of the
cycloalkanes in the hydrogenated organics, the hydrogenated organics can be directly
regarded as the high density jet fuels. As a result, the carbon yields of the high density
jet fuel based on the three lignocellulosic biomasses (hybrid poplar, loblolly pine, and
Douglas fir) were 22.11%, 20.46%, and 19.22% respectively.
199
The gross amount (~1%) of cycloolefins was comparable, which amount meets the
specifications (less than 5%) of commercial jet fuels. Small amounts of
hydro-aromatic hydrocarbons and aromatic hydrocarbons were detected in the
hydrogenated organics from all biomass sources, suggesting that aromatic
intermediates were thoroughly hydrogenated into cycloalkanes. As such, the content of
hydro-cyclic alcohols derived from phenols were much lower under this condition ,
which indicates that partial hydro-cyclic alcohols could be converted into cycloalkanes
via the scission of hydroxyl group substituted on the phenyl and subsequent
hydrocycloaddition of the received aromatic hydrocarbons. There was trace amount of
other oxygenated aromatics in the hydrogenated organics from all biomasses sources,
implying that hydrogenolysis reactions also took place in the process. Small
hydrocarbons detected in the gas revealed that these hydrocarbons with small volume
were from the hydrocracking and hydrogenolysis of larger hydrocarbons.
It was found that the carbon selectivity of both dimethylcyclohexanes and
ethylcyclohexane were maximum from hybrid poplar. Other mono-cyclic alkanes such
as trimethylcyclohexane were also shown the largest carbon selectivity. Therefore
hybrid poplar was the optimal biomass source for the production of mono-cyclic
alkanes. Of the biomass sources, hybrid poplar displayed the lowest carbon selectivity
of hexahydroindan. Unlike the hexahydroindan, the carbon selectivity towards decalin,
bicyclohexyl, perhydrofluorene, and tetradecahydroanthracene was all comparable
200
from three biomasses. Accordingly, these outcomes imply that employing hybrid
poplar as the feedstock in the integral steps could finally improve the overall yield of
cycloalkanes, and obtain the highest carbon selectivity of mono-cyclic alkanes.
6.5 Conclusions
In summary, the microwave-induced pyrolysis followed by hydrogenation is a
profound approach for the production of renewable high-density cycloalkanes from
lignocellulosic biomass. Diverse lignocellulosic biomasses (hybrid poplar, loblolly
pine, and Douglas fir) were taken consideration to manufacture liquid organics in the
catalytic microwave pyrolysis. The optimal condition to maximize the C8 - C16
aromatics in the step of catalytic microwave pyrolysis was at 500 °C with the catalyst
to biomass ratio of 0.25. It was observed that the maximum carbon yield (24.76%) of
aromatics and highest selectivity (94.33%) towards aromatic hydrocarbons in liquid
organics were both from catalytic microwave pyrolysis of hybrid poplar.
These aromatics in n-heptane medium were thoroughly hydrogenated into jet fuel
range cycloalkanes at 250 °C in the presence of 20 wt% (based on the mass of liquid
aromatics) the Raney Ni catalyst. The hybrid poplar derived products obtained the
highest selectivity (95.20%) towards high-density cycloalkanes for jet fuels under very
mild conditions. From lignocellulosic biomass to cycloalkanes-enriched hydrogenated
organics through the integral processes, the maximized carbon yield of the
hydrogenated organics from hybrid poplar was 22.11%. Increasing the reaction
201
temperature and catalyst loading could facilitate the hydrogenation reaction to improve
the production of cycloalkanes. It is more likely that in the future these cycloalkanes
can be directly used as additives in jet fuels or hydrocracked into linear and branched
alkanes in the jet fuel range through hydrocracking and isomerization reactions.
202
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205
CHAPTER SEVEN
SYNTHESIS OF HIGH-DENSITY JET FUEL FROM PLASTICS VIA
CATALYTICALLY INTEGRAL PROCESSES
7.1 Abstract
The present study was aimed at synthesizing JP-5 navy fuel from plastics through a novel
pathway. The consecutive processes for manufacturing JP-5 navy fuel principally included
the catalytic microwave-induced degradation of low-density polyethylene (a model
compound of waste plastics) and the hydrotreatment of obtained liquid organics. The catalytic
microwave degradtation was conducted at the catalytic temperature of 375 °C and catalyst to
feed ratio of 0.1. The carbon yield of the liquid organics from the catalytic microwave
degradation was 66.18%, mainly consisting of a mixture of aromatic hydrocarbons and
aliphatic olefins. Several variables, such as initial pressure and catalyst to reactant ratio, were
employed to determine the optimal condition for the production of alternative jet fuels in the
hydrotreating process. We observed that the aromatic hydrocarbons and aliphatic olefins as
the precursors of jet fuels could be converted into jet fuel range aliphatic alkanes and
cycloalkanes. The hydrotreated organics from the experiment conducted at the reaction
temperature of 250 °C for 2 h, including 31.23% selectivity towards aliphatic alkanes, 53.06%
selectivity towards cycloalkanes, and 15% selectivity towards remaining aromatic
hydrocarbons, which were consistent with the specifications of JP-5 navy fuel. In this regard,
the catalytic microwave degradation of plastics and the hydrotreatment of obtained liquid
206
organics can be regarded as a clear breakthrough to produce alternative jet fuels. From a
commercial point of view, the catalytically integrated processes could be the most feasible for
synthesizing advanced jet fuels (e.g. JP-5 navy fuel).
Keywords: High-density jet fuels; cycloalkanes; aliphatic alkanes; LDPE; hydrogenation
7.2 Introduction
The consumption of virgin plastics has increased exponentially over the past decades,
since they are acting as an indispensable ingredient of utmost importance in daily life.
Global plastic consumption is predicted to continuously rise by 4% in the forthcoming
years,1 leading in parallel to growing waste plastics generated. Most waste plastics
disposed in landfills caused a serious danger towards the environment owing to
plastics degradation and subsequent contaminant generation.2 Plastic that pollutes
oceans and waterways has severe impacts on environment and economy. A new study
in Science indicated the horrifying numbers: In 2010, the study found between 4.8 and
12.7 million metric tons (that was about 10.5 billion to 28 billion pounds) of plastic
entered the oceans. Plastic pollution has an incalculably lethal effect on everything
from plankton to whales.3
The current fate of these waste plastics indicates
approximately 50% were disposed as they cannot be recovered.4 Landfills and
incineration for disposal of waste plastics are commonly utilized among the
conventional existing methods.5, 6 Energy recovery by means of energetic valorization
(incineration) stimulates the release of harmful compounds such as acid gases, dioxins,
207
and furans into the atmosphere together with heavy metals causing damage towards
the environment and human health.7
These inherent issues of plastics recycling impel us to seek alternative valorization
technologies for the production of high value-added chemicals or fuels from waste
plastics. Hence conversion of waste plastics into valuable chemicals and fuels has
attracted crucial interest worldwide. The state of the art on tertiary recycling
technologies like depolymerization offers a promising alternative to plastic recycling.6,
8
Waste plastics valorization approached by thermal degradation give rise to a
heterogeneous hydrocarbon mixture of paraffins and olefins over a wide range of
molecular weights.9 It is noted that a reactively broad spectrum of products is
generated from the thermal degradation of macromolecules to small molecules
complicating their utilization on an industrial scale at present.10 Therefore, the
preliminary products have to undergo downstream catalytic upgrading prior to being
used as transportation fuels.
Obtaining highly valuable fuels from waste plastics by means of simple technologies
is extremely tough because of strict quality standard for fuel substitutes with regard to
the presence of unsaturated and aromatic compounds.11 Catalytic cracking of waste
plastics towards transportation fuels (e.g. gasoline and diesel) are more feasible.12 The
application of a catalyst in a thermal degradation process presents remarkable merits
because products in the desired range of carbon atom number can be enhanced and
208
lower operating temperatures are attained.13 Nonetheless the in-situ catalysis, which
means the direct contact between the catalysts and waste plastics, noticeably resulted
in poor conversion rates and fast catalyst deactivation.14 Compared to conventional
waste management strategies, the ex-situ catalysis (combination of sequential
pyrolysis and catalytic reforming) for waste plastics valorization has evidenced to
entail sound energy and environmental benefits.15 A downstream packed-bed reactor
introduced in the studies was conducive to an upward tendency in contents of aromatic
hydrocarbons and light olefins.16, 17 Suitable catalysts and reactors can in principle
control the product yield, product distribution as well as reduce the reaction
temperature.1 In light of these premises, the conversion of waste plastics into fuels or
valuable chemicals by means of catalytic degradation is regarded as the sustainable
way to valorize the waste plastics.
In the context of increasing price of crude oils, the demand for diesel and jet fuels in
the United States is expected to continue increasing by 27% in the following years as
opposed to gasoline;18 Therefore it is essential to shift alternative fuels towards
distillate-range liquid alkanes in the future. The current jet fuels originated from fossil
resources are principally comprised of aliphatic alkanes (paraffins) and cycloalkanes
(naphthenes).19, 20 These C8 - C16 alkanes are hydrogen saturated, clean burning, and
chemically stable. Nevertheless, linear-chain alkanes and branched-chain alkanes own
lower densities (~0.76-0.78 g/mL), which cannot satisfy the specifications of jet
fuels.21 To overcome the shortage of aliphatic alkanes, jet fuel range cycloalkanes or
209
aromatic hydrocarbons should be synthesized and added into commercial jet fuels (e.g.
Jet A and JP-8).20, 22 Most efforts to produce military jet fuels have concentrated on
increasing the cycloalkanes content; for example, JP-5 navy fuel contains 52.8% of
cycloalkanes, 30.8% of aliphatic alkanes, and 15.9% of aromatic hydrocarbons.23
Notwithstanding, waste plastics can be implemented to produce advanced products via
thermal degradation and catalytic reforming, the combined processes from waste
plastics to high valuable jet fuels
have not paved a feasible route.
In our previous work, Lei and his co-workers have found that microwave-induced
degradation could enhance the selectivity of aromatic hydrocarbons.17, 24, 25 It was also
obtained that the microwave-induced degradation of low-density polyethylene (a
model of waste plastics) generated significant amounts of aromatic and aliphatic
hydrocarbons.17 These unsaturated hydrocarbons belonged to the jet fuel range
compounds, which provide an indication to manufacture jet fuel range alkanes from
these preliminary products. From an industrial implementation point of view, the
catalysts introduced in a process condition favor the economy of the process
commercialized in a continuous operating plant.6 As most of waste plastics like
polyethylene and polypropylene possess approximately 14 wt% hydrogen content
(H/Ceff = 2), these hydrogen-efficient feedstock could reduce or even eliminate the
coke formation in the catalytic degradation process.26 The formation of coke deposited
on the ZSM-5 catalyst used in the aforementioned study was as low as 0.1 wt%, which
could prolong the life time of the catalyst and contribute to lowering the catalyst cost.
210
In addition, it was noticed that jet fuel range aromatic hydrocarbons (C8 - C16) are prone to be
hydrogenated into cycloalkanes under mild conditions.27, 28 Toward this end, the plastics
(low-density polyethylene) were converted into jet fuel range aromatic and aliphatic
hydrocarbons over well-modified ZSM-5 catalyst, according to the optimal condition for
maximizing the carbon yield of aromatic hydrocarbons. The unsaturated hydrocarbons were
stepwise hydrogenated into C8 – C16 hydrocarbons (including cycloalkanes, aliphatic alkanes
and aromatic hydrocarbons) by using Raney nickel as the catalysts, which meet basic
specifications of conventional jet fuels.
7.3 Experimental
7.3.1 Materials
Low density polyethylene (LDPE) (CAS number 9002-88-4) in the form of pellets was
purchased from Sigma-Aldrich Corporation (St. Louis, MO, USA). The density and melting
point of LDPE are 0.925g/cm3 at ambient temperature and 116 °C, respectively. Parent
ZSM-5 (SiO2/Al2O3 Mole Ratio: 50) was purchased from Zeolyst International, USA. Raney
Ni 4200 (slurry in water) in an activated form was supplied by Sigma-Aldrich Corporation
(St. Louis, MO, USA). Toluene (99.7%), p-xylene (99%), ethylbenzene (99%), n-octane
(98%), n-propylbenzene (98%), indane (95%), n-nonane (99%), naphthalene (99.6%),
1,2,3,4-tetrahydronaphthalene (97%), decahydronaphthalene (98%), n-decane (99%),
1-methylnaphthalene (96%), 2-methylnaphthalene (97%), n-undecane (99%), n-dodecane
(99%), anthracene (99%), n-tridecane (98%), n-tetradecane (99%), n-pentadecane (99%), and
211
n-hexadecane (99%) were used as purchased from Alfa Aesar (Ward Hill, MA, USA).
1,2-dimethylcyclohexane (99%), 1,3-dimethylcyclohexane (99%), 1,4-dimethylcyclohexane
(99%), ethylcyclohexane (99%), 1,2,4-trimethylbenzene (98%), 1,2,4-trimethylcyclohexane
(97%), propylcyclohexane (99%), and hexahydroindan (99%) were supplied by
Sigma-Aldrich Corporation (St. Louis, MO, USA).
7.3.2 Catalyst preparation
The activity of parent ZSM-5 was improved by suffering both hydrothermal and
calcined treatments. Under the gentle stirring, parent ZSM-5 powder was added into
deionized water (mass ratio=1) at 60 °C. After addition, the mixture was kept on
stirring for 2 h under this condition. The slurry was then dried at 105 °C till constant
weight. The sequential process was the catalyst calcination: hydrothermally treated
ZSM-5 was calcined at 550°C for 5 h in a muffle furnace. The catalysts were
pelletized and sieved to 20 – 40 mesh. The main characteristics of the catalyst were
reported in our previous study. 17, 25
Raney nickel is notorious for its pyrophoricity, and it may ignite spontaneously when dried in
air. The Raney nickel 4200 (slurry in water) was thus dried at 60 °C till constant weight in the
atmosphere of nitrogen to avoid contact with air, prior to the subsequent catalytic test. The
textural properties of as-received Raney Ni 4200 catalyst and the SEM image are outlined in
Table 7.1 and Fig. 7.1, respectively.
212
Table 7.1 Textural properties of as-received Raney-Ni 4200 catalysts.a
Raney-Ni catalyst
a
SBET: BET surface area;
SBET
Vpore
Spore
dpore
(m2/g)
(cm3/g)
(m2/g)
nm
38.1
0.112
43.0
10.4
Vpore: pore volume;
Spore: pore surface area;
dpore:
average pore size.
Fig. 7.1 SEM image of the as-received Raney Ni catalyst.
Raney nickel is notorious for its pyrophoricity, and it may ignite spontaneously when dried in
air. The Raney nickel 4200 (slurry in water) was thus dried at 60 °C till constant weight in the
atmosphere of nitrogen to avoid contact with air, prior to the subsequent catalytic test. The
textural properties of as-received Raney Ni 4200 catalyst and the SEM image are outlined in
213
Table 1 and Fig. 1, respectively.
7.3.3 Catalytic microwave-assisted degradation of low-density polyethylene (LDPE)
Detailed experimental setting was described in our previous studies.17, 29 Fixed loading
of low-density polyethylene pellets (100 g) for each run were placed in a 500 mL
quartz flask inside the microwave oven (Shanghai, China) by a constant microwave
power setting (700 W). 0.5 g of activated carbon powder was used as the absorber for
the microwave-assisted degradation. All reactions of microwave degradation were
conducted at the temperature of 350 °C for 20 min until all LDPE pellets were
completely vaporized. The pyrolytic volatile vapors from the flask passed through a
packed bed catalysis reactor which was filled with catalyst. As previous work
reported, the optimal condition to maximize the liquid yield was set at 375 °C.17 In this
regard, the catalytic temperature was hold at 375 °C and catalyst (10 g) to feed ratio
was kept constantly at 0.1. The catalytic microwave degradation was duplicated for
four times in order to gain abundant liquid organics which were thereafter combined
together for the hydrogenated process.
7.3.4 Hydrotreatment of liquid organics derived from catalytic microwave degradation
To saturate the liquid organics evolved from catalytic microwave degradation, a closed
reaction system with a stirred stainless batch reactor of the 4592 micro stirred reactor (with a
50 mL vessel) and a 4848 reactor controller from Parr Instrument Company (Moline, IL,
USA) was used (Fig. S2). The liquid organics (8 g) and the catalysts with intended ratio were
214
loaded into the reactor. Then the reactor was sealed and vented for five times with hydrogen
to get rid of the air present in the vessel. Hydrogen was subsequently adjusted to reach the set
pressure. The automatic controller was employed to control the temperature and the
revolution of stirrer (100 rpm). The pressure inside the reactor was recorded and the reactions
proceeded at a set temperature for the intended time. After the experiment finished, stirring
was stopped and the reactor was rapidly cooled to ambient temperature. Then, the gas was
collected for analysis and the reactor was depressurized. Consequently the liquid product was
filtered to remove catalyst particles.
7.3.5 Analytical techniques
Elemental analysis of LDPE pellets, liquid samples, char, and coke deposited on spent
catalysts was conducted using a 2400 Series II CHN/O Elemental Analyzer
(PerkinElmer, USA).
The chemical composition of the bio-oils was characterized and qualified by Agilent
7890A GC-MS (GC–MS; GC, Agilent 7890A; MS, Agilent 5975C) with a DB-5
capillary column. The GC was first programmed to heat to 45°C for 3 min followed by
heating to 300°C at a rate of 10°C/min. The injection sample size was 1 μL. The flow
rate of the carrier gas (helium) was 0.6mL/min. The ion source temperature was 230°C
for the mass selective detector. Compounds were identified by comparing the spectral
data with that in the NIST Mass Spectral library. The area percent of changed
215
concentrations of model compounds obtained from GC/MS results was utilized to
predict product concentration in bio-oils.
The gaseous product was collected in a 1L Tedlar gas bag and then offline analyzed by an
INFICON 3000 Micro-GC (INFICON Inc., Santa Clara, CA, USA) system with a thermal
conductivity detector (TCD). A standard gas mixture consisting of H2, N2, CH4, CO, CO2,
C2H4, C2H6, and C3H6 was used to calibrate the yield of non-condensable gas. Alkanes and
olefins (>C4) in gas samples were either not detected or negligible in this research.
216
Table 7.2 Experimental design and product yield distribution of hydrotreated
organics.a
Yield (area in %)b
Initial
Catalyst
to
pressure (psi)
reactant ratio AA
Entry
CA
HAH
AH
Others
H-1
200
0.02
17.13
2.61
12.05
67.05 1.16
H-2
200
0.08
18.00
3.05
11.17
67.05 0.73
H-3
800
0.02
23.19
12.70
16.80
45.80 1.51
H-4
800
0.08
23.09
27.23
11.76
36.74 1.18
H-5
76
0.05
18.01
3.76
10.96
66.48 0.79
H-6
924
0.05
21.68
16.57
9.56
50.56 1.63
H-7
500
0.05
18.58
6.94
13.81
59.37 1.30
H-8
500
0.05
18.69
7.81
14.37
58.12 1.01
H-9
500
0.05
19.11
7.23
14.18
58.51 0.97
H-10
500
0.05
18.99
7.15
14.47
58.46 0.93
H-11
500
0.05
19.02
7.03
14.22
58.65 1.08
H-12
500
0.01
18.10
5.78
14.09
60.20 1.83
H-13
500
0.09
19.56
9.91
13.94
55.01 1.58
a
b
Reaction conditions: reaction temperature, 150 °C ;
AA: Aliphatic alkanes; CA: Cyclic alkanes;
reaction time, 1 h.
HAH: Hydro-aromatic
hydrocarbons; AH: Aromatic hydrocarbons.
7.3.6 Experimental methods and data evaluation
A central composite experimental design (CCD) was employed to optimize the process
conditions and the yield of jet fuel range alkanes. The initial pressure of hydrogen (X1,
psi) and catalyst to reactant ratio (X2) were chosen as independent variables as
illustrated in Table 7.2. In these experiments, the mass of catalyst varied from 0.08 to
217
0.72 g (with respect to the ratio from 0.01 to 0.09), while the initial pressure ranged
from 76 to 924 psi.
The coke mass was determined by the difference before and after catalytic
degradation. The weight of non-condensable gas was calculated using the following
equation:
L
= !"#
−
− ℎ
−
(1)
Overall carbon yields of the liquid, gas, and solid products and carbon selectivity of a
specific product were calculated based on the following equations.
MNOPQR STUVW =
XQVUY QZ [NOPQR TR N \OQW][^
× _``%
XQVUY QZ [NOPQR ZUW TR
MNOPQR YUVU^TbT^S =
XQVUY QZ [NOPQR TR N \OQW][^
× _``%
XQVUY QZ [NOPQR TR TWUR^TZTUW \OQW][^Y
(a)
(c)
7.4 Results and discussion
7.4.1 Catalytic transformation of LDPE into liquid organics
In order to manufacture advanced jet fuels, the transformation of LDPE into the
precursors of jet fuels in the carbon number range from 8 to 16 is primarily required.
Based on our previous study for catalytic microwave degradation of LDPE,17 optimal
catalytic temperature set at 375 °C could maximize the yield of liquid organics.
Henceforth, the LDPE was subjected to the optimized catalytic temperature. It was
found that the liquid organics from catalytic microwave degradation of LDPE was
64.41±5.20 wt%. There have never been any results could achieve such high yield
218
organics.4, 12 The mild catalytic temperature and sufficient catalyst loading could favor
the conversion rate of waxes from direct thermal degradation into the lumps of liquid
organics through catalytic cracking, cyclization, and oligomerization reactions. The
overall carbon yield distribution from catalytic microwave degradation of LDPE is
shown in Fig. 7.2(A). It is noteworthy that the share of liquid organics accounted for
the dominant composition, occupying 66.18%; while both the minor carbon yield of
coke and char were approximately 1%, which is as low as our aforementioned result.17
It is also noted that the low carbon yield (32.01%) of non-condensable gas was
calculated by difference. The gas consisted of ethylene, ethane, hydrogen, and
methane, originating from the large extent of cracking and oligomerization reactions.
Comparing with the feedstock using lignocellulosic biomasses in the system of
catalytic pyrolysis,29,
30
the carbon yield of non-condensable gas derived from the
feedstock of plastics was relatively low.
219
Fig. 7.2 Overall carbon yield distribution (C mol%), (A) and carbon selectivity of main
chemical compounds (C mol%), (B) in the liquid organics from catalytic microwave
degradation of LDPE.
The carbon selectivity of main chemical compounds in the liquid organics from
catalytic microwave degradation of LDPE is depicted in Fig 7.2 (B). In general, the
typical compositions have been categorized into aromatic hydrocarbons, olefins,
alkanes; and partial chemicals in the organics were quantified. Carbon selectivity
towards xylenes (C8H10) was 18.34 %, which was the dominant part in the liquid
220
organics. Likewise, the large amount of propylbeneze (C9H12) was also obtained,
displaying remarkable carbon selectivity (14.04%).
Consequently, mono-cyclic
aromatic hydrocarbons in the jet fuel range were preferentially formed over the
well-promoted ZSM-5 catalyst with regard to aliphatic hydrocarbons, which can be
ascribed to the fast diffusion of the intermediates toward the active sites inside the
micropores. A small amount of double-cyclic aromatic hydrocarbons including
naphthalene and its derivatives was evolved from the oligomerization and
polymerization reactions of mono-cyclic aromatic hydrocarbons.31 Catalytic cracking
of waxes from thermal degradation of LDPE also gave rise to small amounts of
aliphatic hydrocarbons. The results imply that the products gained after catalytic
microwave degradation of LDPE at 375 °C is mostly made up of hydrocarbons in the
jet fuel range (C8 – C16).
7.4.2 Hydrotreatment of liquid organics derived from catalytic microwave degradation
7.4.2.1 Chemical composition of hytrotreated organics
Since the liquid organics produced by catalytic microwave degradation of LDPE were
principally comprised of C8 - C16 range unsaturated hydrocarbon, the controllable
adjustment of unsaturated hydrocarbon with 8 - 16 carbon numbers are considered as
precursors of jet fuels. In the hydrotreating process, the directional production of C8 C16 unsaturated hydrocarbon was hydrotreated by using Raney Ni catalyst under
low-severity conditions. The loss of the liquid organics was ignored if considering the
221
recovery (more than 95 wt%) of the liquid organics could be achieved after the
hydrotreating experiments.
The experimental design and product yield distribution of hydrotreated liquid organics are
summarized as a function of initial pressure and catalyst to reactant ratio in Table 7.2. It was
observed that remaining aromatic hydrocarbons were in the range from 36.74 to 67.05%
depending on alterations of reaction conditions. The high amounts of aromatic hydrocarbons
retained are due in part to the mild reaction conditions, such as low reaction temperature and
reaction time. These results also indicate that the initial pressure and catalyst to reactant ratio
had a slight influence in the total amounts of aliphatic alkanes; while the enhancement in the
two variables led toward a parallel increase in the yields of cycloalkanes. The optimal
condition for maximum yields of saturated hydrocarbons was found to be the initial pressure
of 800 psi and the catalyst to reactant ratio of 0.08. The corresponding amount of
cycloalkanes reached at the highest selectivity (27.23%), whereas the yield of aromatic
hydrocarbon dropped to 36.74%. Thus, these augments of cycloalkanes took place mostly at
the expense of the aromatic hydrocarbon because its share decreased. As such, increasing the
initial pressure and catalyst to reactant ratio contributed to improving the yield of
hydro-aromatic hydrocarbons. These results are consistent with hydrogenation reactions
predominantly occurring under low-severity conditions.32 Since these high amounts of
aromatic hydrocarbon remained in the hydrotreated organics cannot be straightforward
utilized as jet fuels, further efforts for converting aromatic hydrocarbon into saturated
hydrocarbons containing aliphatic and cyclic alkanes should be made.
222
Fig. 7.3 Chemical composition of hydrotreated organics on the basis of initial
pressure. Key: Aliphatic alkanes (red); Cycloalkanes (green); Hydro-aromatic
hydrocarbons (blue); Aromatic hydrocarbons (sky blue); Others (pink). Reaction
condition: Reaction temperature, 150 °C; Raney Ni catalyst, 5 wt% with respect to
reactant mass; Reaction time, 1 h.
7.4.2.2 The effect of initial pressure on the chemical composition of hydrotreated oils
To further understand the effect of initial pressure on the hydrotreated organics, the
reaction implemented at the initial pressure of 1200 psi was compared to previous
results. The initial pressure within the 76 – 1200 psi range was investigated working at
150 °C for 1 h and using the catalyst to reactant ratio of 0.05. Fig. 7.3 shows the yields
by categories obtained in these hydrotreated reactions. Seemingly, the arguments of
initial pressure improved the extent of hydrogenation reactions toward cycloalkanes,
223
reaching the highest yield in this fraction (25.06%) at 1200 psi of hydrogen pressure.
At 1200 psi of initial pressure, the most remarkable result is a drop in the total
aromatic hydrocarbon contents. The enhancement of initial pressure is supposed to
hydrogenate the aromatic hydrocarbons to cycloalkanes and hydro-aromatic
hydrocarbons. However, the yield of hydro-aromatic hydrocarbons remained fairly
low (around 10%), likely because of hydro-aromatic hydrocarbons acting as the
intermediates, which could be thereafter transformed into saturated hydrocarbons.
Hence the increase in cycloalkanes took place at the expense of aromatic
hydrocarbons. Unlike the cycloalkanes, it is noteworthy to indicate that, even at the
low initial pressure (76 psi), the aliphatic alkanes grew to approximately 18%, which
is indicative of its high hydrogenation activity of aliphatic olefins. Besides no aliphatic
olefins were detected regardless of the initial pressure; therefore these conditions
allowed for the removal of them entirely from jet fuels. These outcomes imply that a
high initial pressure, which means that more volume of hydrogen was provided on the
Raney nickel surface, contributed to more hydrogenation reactions occurring in the
fixed time.
224
Fig. 7.4 Chemical composition of hydrotreated organics with respect to catalyst to
reactant ratio. Key: Aliphatic alkanes (red); Cycloalkanes (green); Hydro-aromatic
hydrocarbons (blue); Aromatic hydrocarbons (sky blue); Others (pink).
Reaction
condition: Reaction temperature, 150 °C; Initial pressure, 500 psi; Reaction time, 1 h.
7.4.2.3 The effect of catalyst loading on the chemical composition of hydrotreated
organics
Fig. 7.4 shows representative results of hydrotreated organics regarding the chemical
compounds on the basis of catalyst to reactant ratio at the same initial pressure of 500
psi and reaction temperature of 150 °C for 1 h. The catalyst to reactant ratio did not
have a vital influence in the yield of aliphatic alkanes, which slightly increased. Even
though more catalysts were employed, other reactions except the hydrogenation of
aliphatic olefins did not take place due to the mild reaction conditions.
The increase
in the amount of catalyst loading led toward an improvement in the extent of
225
hydrogenation reactions, augmenting the share of cycloalkanes and decreasing those of
aromatic hydrocarbons especially when a ratio of 0.13 was used. Nevertheless, the
increase in catalyst to reactant ratio was less effective than initial pressure in
augmenting the extent of aromatic hydrocarbons hydrogenation. Over half of
hydrotreated organics remained belonged to aromatic hydrocarbons, dramatically
exceeding specification of jet fuels. It is also can be seen that the total amount of
hydro-aromatic hydrocarbons were slightly fluctuant at 14%. These results indicate
that high catalyst to reactant ratio, which means that more active sites were offered on
the Raney nickel surface, gave rise to more hydrogenation reactions.
7.4.2.4 The effect of reaction temperature on the chemical composition of hydrotreated
organics
Based on the previous comparison, the initial pressure of 250 psi and catalyst to
reactant ratio of 0.05 were feasible for the following experiments. In order to further
understand chemical reactions in the process and obtain more insight into hydrotreated
organics, the chemical compounds of hydrotreated organics are elucidated as a function of
reaction temperature at the same reaction time of 1 h in Table 7.3. It was found that the
reaction temperature had a significant effect on the hydrotreated reactions. The total amounts
of both aliphatic alkanes and cycloalkanes progressively increased with the increasing
reaction temperatures. It was noticed that the share of cycloalkanes grew to 29.37% as the
reaction temperature increased to 200 °C, implying that elevated reaction temperature favored
the hydrocycloaddition rate of aromatic hydrocarbons. Meanwhile, there is crucial content of
226
aliphatic alkanes, generally increasing to 25.12% alongside the increment of reaction
temperature. It was observed that this high yield of aliphatic alkanes was superior to these
yields even from the severest condition of highest initial pressure and catalyst to reactant
ratio,which indicates that the high reaction temperature (200 °C) possibly facilitated the
scission of cycloalkanes to form the aliphatic alkanes through hydrocracking reactions.27
Table 7.3 Products distribution and partial alkanes’ carbon selectivity as a function
of reaction temperature at the reaction time of 1 h.a
Temperature ( °C )
150
175
200
Aliphatic alkanes
18.88
22.91
25.12
Cycloalkanes
7.23
24.70
29.37
Hydro-aromatic hydrocarbons
14.21
11.05
11.42
Aromatic hydrocarbons
58.62
39.66
32.71
Others
1.06
1.68
1.38
1,4-dimethylcyclohexane
-
4.16
6.63
1,3-dimethylcyclohexane
-
1.63
2.05
1,2-dimethylcyclohexane
-
1.33
2.12
Ethylcyclohexane
2.78
1.83
5.78
Octane
0.91
1.22
1.58
1,2,4-trimethylcyclohexane
0.20
0.69
0.35
Propylcyclohexane
1.24
5.73
4.55
Nonane
0.89
1.41
1.98
Hexahydroindan
0.54
2.00
2.18
Decalin
0.11
0.53
0.55
Overall selectivity (% in area)
Alkanes selectivity (C mol%)
227
Decane
1.12
2.06
2.31
Undecane
1.45
2.30
2.04
Dodecane
1.72
2.34
3.12
Tridecane
1.26
1.43
1.58
Tetradecane
1.28
1.47
1.73
Pentadecane
1.17
1.40
1.64
Hexadecane
1.32
1.58
1.71
a Reaction
condition: Initial pressure, 500 psi; Raney Ni catalyst, 5 wt% with respect
to reactant mass.
On the other hand, the total amounts of aromatic hydrocarbons experienced a
gradually declined tendency as the reaction temperature went up to 200 °C. The
reduction of aromatic hydrocarbons was hydrogenated into cycloalkanes and aliphatic
alkanes via hydrogenation and hydrocracking reactions. It is noted that the share of
hydro-aromatic hydrocarbons was not significantly impacted by the reaction
temperature. Although some aromatic hydrocarbons were hydrogenated into
hydro-aromatic hydrocarbon, these hydro-aromatic hydrocarbons as the intermediates
were simultaneously converted into saturated hydrocarbons. In the gas fraction,
unreacted hydrogen was detected at the end of reaction accompanied with minor
volume of small hydrocarbons (such as methane and ethane), implying that the
experiments were not carried out under hydrogen starved conditions. These gaseous
hydrocarbons detected could be produced from the hydrocracking of liquid
hydrocarbons. These results suggest that not only hydrogenation reactions but also
hydrocracking reactions could take place at high reaction temperature.
228
Reaction temperatures also had a significant effect on the carbon selectivity of specific
alkanes including aliphatic and cyclic alkanes. It was noteworthy that the carbon
selectivity of mono-cyclic alkanes significantly increased as the reaction temperature
increased, especially from 150 to 175 °C. Meanwhile, 1, 4-dimethylcyclohexane
(C8H16) and propylcyclohexane (C9H18) gradually increased to 4.16% and 5.73%,
which indicates that the reaction temperature above 175 °C presumably accelerated the
hydrogenation rate of the corresponding aromatic hydrocarbons. For the carbon
selectivity of ethylcyclohexane (C8H16), the enhancement of reaction temperature from
175 to 200 °C led toward a remarkable increase. The carbon selectivity of polycyclic
alkanes (e.g. decalin) was slightly impacted by the increment of reaction temperature.
There were small upward tendencies towards carbon selectivity of hexahydroindan
(C9H16) and decalin (C10H18) in the range from 150 to 200 °C. This fact was possibly
ascribed to the low amounts of double-ring aromatic hydrocarbon in the liquid organic
from catalytic microwave degradation; and it is very tough for the double-ring
aromatic hydrocarbons to be hydrogenated into saturated hydrocarbon below 200 °C.
Likewise, the carbon selectivity of aliphatic alkanes was also affected by the reaction
temperature. The carbon selectivity of aliphatic alkanes, such as octane (C8H18) and
nonane (C9H20), showed a steady increment as the reaction temperature grew. These
outcomes imply that the increment of aliphatic alkanes was attributed to hydrogenation
of aliphatic olefins and hydrocracking of cycloalkanes simultaneously occurring at the
high temperature range.12
229
Nevertheless, the optimal result (over 30% of aromatic hydrocarbons plus
approximately 10% hydro-aromatic hydrocarbons remained in the hydrotreated
organics) cannot meet the specifications of conventional jet fuels, especially for
advanced jet fuels. Other variable (e.g. reaction time) should be investigated to obtain
higher valuable fuels that can directly be used as conventional jet fuels or drop-in
fuels. Apart from the aforementioned variables, reaction time was another crucial
factor that influenced product distribution and chemicals’ carbon selectivity.17 Table
7.4 summarizes products distribution and partial alkanes’ carbon selectivity with
regard to reaction temperature at the reaction time of 2 h. In particular, the production
distribution prominently shifted towards the cycloalkanes (31.23 – 53.06%) with
increasing reaction temperature to 250 °C, suggesting that hydrogenation reactions
were enhanced because of the longer reaction duration and higher reaction
temperature. The amount of cycloalkanes in the hydrotreated organics at 250 °C was
equal to that in JP-5. It was noticeable that the total amount of aliphatic alkanes
appeared to generally increase to 31.23% at 250 °C, whose amount was similar with
the content (30.8%) in conventional JP-5.23 Unlike the saturated hydrocarbons, the
total amounts of aromatic hydrocarbon and hydro-aromatic hydrocarbons distantly
declined, particularly from 200 to 250 °C. The total amounts of hydro-aromatic
hydrocarbons and aromatic hydrocarbons were below 15% at 250 °C, which satisfies
the limited ceiling of aromatic hydrocarbons in Jet-A and JP-8. It is noteworthy to
point out that this remarkable drop in the amount of aromatics is a positive effect for
the quality of JP-5 fraction provided that the legislation limit (15.9%) is not surpassed.
230
These results indicate that the obtained liquid products at 250 °C for 2 h can be
directly used as alternatives for the formulation of JP-5 navy fuel.
Table 7.4 Products distribution and partial alkanes’ carbon selectivity as a function of
reaction temperature at the reaction time of 2 h.a
Temperature ( °C )
150
200
250
Aliphatic alkanes
21.12
26.37
31.23
Cycloalkanes
31.13
40.18
53.06
Hydro-aromatic hydrocarbons
9.81
10.84
7.22
Aromatic hydrocarbons
36.19
21.12
7.78
Others
1.76
1.49
0.71
1,4-dimethylcyclohexane
8.40
8.18
9.45
1,3-dimethylcyclohexane
1.79
2.02
3.04
1,2-dimethylcyclohexane
3.77
2.98
4.21
Ethylcyclohexane
2.34
1.62
3.88
Octane
1.01
0.93
1.84
1,2,4-trimethylcyclohexane
0.25
1.05
2.92
Propylcyclohexane
5.75
9.80
10.21
Nonane
1.22
0.98
1.45
Hexahydroindan
1.81
1.34
2.02
Decalin
0.66
0.76
0.87
Decane
1.54
2.00
2.77
Undecane
2.43
3.45
4.21
Dodecane
2.21
3.80
4.56
Tridecane
1.45
2.08
2.79
Overall selectivity (% in area)
Alkanes selectivity (C mol%)
231
Tetradecane
1.47
2.08
2.92
Pentadecane
1.44
2.04
3.01
Hexadecane
1.51
2.22
3.33
a Reaction
condition: Initial pressure, 500 psi; Raney Ni catalyst, 5 wt% with respect to
reactant mass.
The carbon selectivity of mono-cyclic alkanes significantly increased at the reaction
time of 2 h, especially for propylcyclohexane, as the reaction temperature was
elevated; while the carbon selectivity for hexahydroindan and decalin slightly went up
in the range of 150 – 250 °C. Even if naphthalenes can be converted into
corresponding cycloalkanes under this condition,27 the low amounts of naphthalenes
could make the slight variation of double-cyclic alkanes. Meanwhile, the high reaction
temperature (250 °C) could allow the double-cyclic alkanes to be hydrocracked and
oligomerized, forming monocyclic alkanes or their derivatives. As expected, the
carbon selectivity of main linear alkanes also increased with the augment of reaction
temperature. From the optimal result at 250 °C, up to 26.88% of the total carbon
selectivity towards jet fuel range linear alkanes (C8 – C16) was achieved. The high
purity of the linear alkanes provides an excellent performance of the hydrotreated
organics as the replacement of jet fuels. In terms of gaseous results at 250 °C, a trace
volume of small-chain hydrocarbons were found, suggesting that the hydrocracking
and oligomerization reaction have taken place in the process. Wherefore, prolonging
the reaction time can principally enhance the hydrogenation reactions; hydrocracking
232
and oligomerization reactions might also occur jointly at the high reaction
temperature.
7.4.3 Reaction pathway for the conversion of plastics (LDPE) into JP-5 navy fuel
These observations are the key point to propose the reaction pathway for the
conversion of plastics (LDPE) into JP-5 navy fuel. Based on the quantified products
distribution in this study, and related results from catalytic microwave-induce
degradation of LDPE as a function of several variables,17 the overall reactions network
(including catalytic microwave degradation and hydrotreating process) is shown in
Fig. 7.5. Thermal degradation of polyethylene has been reported to take place through
two mechanisms: (1) random scission to yield long-chain hydrocarbons; and (2)
chain-end scission of the oligomers to yield low-molecular weight products.33, 34 The
two abovementioned mechanisms took place simultaneously, giving rise to free
radicals along with the carbon chains; thereafter the cleavage of the molecule caused
the formation of a molecule with an unsaturated end and the other with a terminal free
radical.35 Consequently the radical fragments were converted into straight chain
dienes, alkenes and alkanes through hydrogen transfer reactions.35
233
Fig. 7.5 Reaction pathways for the conversion of plastics (LDPE) into jet fuels.
The volatiles subsequently underwent catalytic cracking over ZSM-5 through two
carbocationic mechanisms.33,
36
Once these carbocations were generated, various
acid-catalyzed reaction could take place over the acid sites, including aromatization,
cyclization, cracking, isomerization, and oligemerization.12 Aromatization and
cyclization reactions proceed by the way of hydrogen transfer reactions, whilst
cracking reactions usually occurs by means of β-scission reactions. Herein, lager
intermediate olefins derived from the thermal degradation could not enter ZSM-5
micropores where the majority of the active acid sites were located. The
macromolecules were degraded on the active sites on the external surface of the zeolite
crystallites; small fractions relatively diffused into the zeolite micropores and further
reacted on the internal sites.37 Moreover, the intermediate olefins stepwise underwent
234
oligomerization, cyclization and hydrogen transfer reactions, contributing to forming
aromatic hydrocarbons.38 Eventually the liquid organic with high quality of aromatics
and aliphatic olefins were obtained when the suitable loading of well-promoted ZSM-5
catalyst was introduced.
It was observed that different reactions took place in the hydroreforming experiments
of liquid organics from the LDPE thermal degradation, such as hydrogenation,
hydrocracking, and hydroisomerization.4 Since the liquid organics resulting from the
catalytic microwave degradation step in mostly produced by a mixture of aromatic
hydrocarbons, aliphatic olefins, and alkanes, hydrogenation of the alkenes is the one of
the first step to take place.4 When the much higher-severity conditions were employed,
the aromatic hydrocarbons were partially hydrogenated into cycloalkanes or
hydro-aromatic hydrocarbons. Likewise, the total amounts of aliphatic alkanes were
enhanced because the formed cycloalkanes could further undergo the hydrocracking to
generate the aliphatic alkanes. According to the gaseous fraction, hydrocracking of
liquid hydrocarbons over the metal sites, leading toward methane and ethane, should
occur. However, the gaseous small hydrocarbons in all reactions carried out were
practically negligible.
7.5 Conclusions
This study illustrated that the microwave-induced degradation followed by
hydrotreating process is a profound approach for the production of jet fuels from
235
plastics (LDPE). Diverse variables (initial pressure, catalyst to reactant ratio, reaction
temperature, and reaction time) were taken consideration to manufacture advanced jet
fuels. In the first step of microwave-induced degradation of LDPE over well-promoted
ZSM-5 catalyst, the overall carbon yield of liquid organics was 66.18% at the catalytic
temperature of 375 °C with the catalyst to reactant ratio of 0.1; while the carbon yields
of both coke and char were approximately 1%. Such high carbon yield of liquid
organics could improve the feasibility of the process by using waste plastics as the
feedstock in an industrial scale. Furthermore, the trace yield of coke deposited on the
catalyst assisted in the promotion of the catalyst lifetime (results not shown), which
could significantly reduce the operating cost of the catalyst in a biorefinery.
After the hydrotreating process by using Raney nickel as the catalyst, the recovery of
the hydrotreated organics could reach more than 95 wt% in all experiments. It was
observed that the optimal reaction for the production of advanced jet fuels was
conducted at the reaction temperature of 250 °C for 2 h. The total amounts of aliphatic
alkanes (31.23%), cycloalkanes (53.06%), and aromatic hydrocarbons (15.00%) in the
jet fuel range were obtained under the mild reaction condition. The outcomes indicate
that the specifications of hydrotreated organics are equal to those of JP-5 navy fuel,
which can be directly used as the alternatives or additives for JP-5 navy fuel. From this
perspective, the catalytically integral processes of waste plastics by using inexpensive
catalysts deliver a novel and feasible pathway in a biorefinery, specifically targeting
JP-5 navy fuel.
236
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A. López, I. de Marco, B. M. Caballero, M. F. Laresgoiti, A. Adrados and A.
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W. Huber, Green Chemistry, 2010, 12, 1933.
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Chem., 2015, DOI: 10.1039/c5gc00516g, 4029-4036.
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General, 2013, 455, 114-121.
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Fuel, 2015, 160, 375-385.
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Engineering Chemistry Research, 2012, 51, 13915-13923.
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239
CHAPTER EIGHT
FROM PLASTICS TO DIVERSE GRADES OF JET FUELS VIA
COMBINED CATALYTIC CONVERSIONS
8.1 Abstract
High carbon yields of various grades of jet fuels were obtained from plastics through a novel
pathway.
The
consecutive
processes
principally
proceeded
via
the
catalytic
microwave-induced degradation of low-density polyethylene (a model compound of waste
plastics) followed by hydrogenation of raw organics. The catalytic microwave degradation
was conducted at the catalytic temperature of 375 °C and catalyst to feed ratio of 0.1 and 0.2
to manufacture distinct contents of aliphatic and cyclic hydrocarbons. The carbon yields of
the raw organics from the catalytic microwave degradation was 66.18 and 56.32%. Several
variables were employed to determine the suitable conditions for the production of alternative
jet fuels in the hydrogenation process. We observed that the unsaturated hydrocarbons as the
precursors of jet fuels could be converted into jet fuel range aliphatic and cyclic alkanes.
Since n-heptane medium could assist in the hydrogenation of unsaturated hydrocarbons, all
hydrogenation experiments were conducted in such medium. With the presence of 10 and 20
wt% well-promoted ZSM-5 catalyst, the overall carbon yields of hydrogenated organics with
respect to raw plastics were approximately 54 and 63%, respectively. The alterations of
reaction conditions did not have an effect on the overall carbon yield of hydrogenated
organics. The raw organics from catalytic microwave degradation by using 10 wt%
well-promoted ZSM-5 catalyst could be converted into alternatives of JP-5 navy fuel or
additives of Jet A and JP-8 in the presence of 10 wt% Raney Ni catalyst at 200 °C for 2 h.
The raw organics from catalytic microwave degradation by using 20 wt% well-promoted
240
ZSM-5 catalyst could transformed into high energy-density jet fuels (e.g. RJ-5 and JP-10)
with regard to various variables. Home-made Raney Ni catalyst showed much higher
catalytic performance than that of as-purchased Raney Ni 4200 catalyst. In this regard, the
catalytic conversions of plastics can be regarded as a clear breakthrough to produce various
grades of jet fuels.
Keywords: Jet fuels; plastics; well-promoted ZSM-5 catalyst, Home-made Raney Ni catalyst;
catalytic conversions
8.2 Introduction
The consumption of virgin plastics has increased exponentially over the past decades, since
they are acting as an indispensable ingredient of utmost importance in daily life. From 2010
to 2016, global plastic consumption is predicted to continuously rise in the similar yearly
number (4%),1 leading in parallel to growing waste plastics generated. Most waste plastics
disposed in landfills caused a serious danger towards the environment owing to plastics
degradation and subsequent contaminant generation.2 Plastic that pollutes oceans and
waterways has severe impacts on environment and economy. A new study in Science
indicated the horrifying numbers: In 2010, the study found between 4.8 and 12.7 million
metric tons (that was about 10.5 billion to 28 billion pounds) of plastic entered the oceans.
Plastic pollution has an incalculably lethal effect on everything from plankton to whales.3
The current fate to manage waste plastics are generally limited by means of landfills and
incineration.4, 5 Hence conversion of waste plastics into valuable chemicals and fuels has
attracted crucial interest worldwide. In the context of limited crude oils, jet fuels have
241
become one of the fastest growing refinery product market demands.6, 7 The demand for jet
fuels especially for high energy-density jet fuels is expected to increase by 27% in the
forthcoming years as opposed to gasoline.8 The current jet fuels originated from fossil
resources
principally consist of aliphatic alkanes (paraffins) and cycloalkanes
(naphthenes).9, 10 Nevertheless, linear-chain alkanes and branched-chain alkanes with lower
densities (~0.76-0.78 g/mL) cannot satisfy the specifications of jet fuels.11 Their low densities
result in their relatively poor volumetric heating values of conventional jet fuels.12 To
overcome the shortage of aliphatic alkanes, ring-saturated alkanes or aromatic hydrocarbons
should be synthesized and added into commercial jet fuels (e.g. Jet A and JP-8).10, 13, 14
On the other hand, the substitution of high-density jet fuels, such as JP-10 and RJ-5, has
attracted great attention.15 Significant ring strain contributes to a higher heat of combustion
(JP-10 at 39.6 MJ/L and RJ-5 at 44.9 MJ/L)
12.
As a result, most efforts to produce
high-density jet fuels have concentrated on increasing the cycloalkanes content.
Notwithstanding, there are elaborate research work in literature describing the production of
hydrocarbons from waste plastics;2, 4, 5, 16-18 research on direct route from waste plastics to
high energy-density jet fuels is still in its fancy. There is also no search reported on tailored
catalysts that converted plastics into jet fuels (kerosene). Accordingly, waste plastics can
serve as ideal feedstock for the production of high energy-density jet fuels.
In our previous work, we presented a catalytically integrated process for converting
lignocellulosic biomass into renewable cycloalkanes for jet fuels.19 These processes involved
242
three key steps: (1) catalytic microwave-induced pyrolysis of intact biomass into
aromatics-enriched bio-oils; (2) extraction of aromatics-enriched bio-oils by n-heptane; (3)
hydro-cycloaddition of the liquid organics into C8 – C16 cycloalkanes. Nevertheless, we have
also observed that the formation of renewable cycloalkanes was at relatively low carbon yield
(~ 20%) by using lignocellulosic biomass as feedstock. Instead, we intend to illustrate how
these processes can be utilized to transform plastics into various grades of jet fuels, which can
be used as the replacements of conventional jet fuels or blended into current jet fuels.
In scaling up this high energy-density jet fuel process we have adjusted the chemistry to
manufacture a mixture of different content hydrocarbons including aliphatic alkanes,
naphthenes (cycloalkanes) and minor aromatic hydrocarbons with the carbon number from 8
to 16. The overall process is outlined in Fig.8.1 and consists of two fundamental steps. It is
manifested that microwave-assisted degradation technology is one of the most promising
methods for enhancing and accelerating chemical reactions due to effective heat transfer
through microwave irradiation.20 Comparing with conventional heating, microwave-induced
degradation encloses the potentials of fast and selective heating, easy control of reaction
conditions, low reaction temperatures and energy requirements.21, 22 In this sense, the first
step is the production of aliphatic alkenes and aromatic hydrocarbons by catalytic
microwave-induced degradation. The aliphatic alkenes and aromatic hydrocarbon with
carbon number in the range of 8 to 16 is highly promising precursors for straight-chain and
cyclic alkanes in jet fuels.19,
23, 24
Comparing with various mediums, n-heptane was the
optimal solvent that can assist in the hydrogenation rate of aromatic hydrocarbons
243
19.
The
second step is thus the hydrogenation of these unsaturated hydrocarbons into jet fuel range
alkanes in n-heptane medium by using various Raney nickels as the catalysts to optimize the
results. Toward this end, this research is to demonstrate how the integrated processes can be
used to produce distinct grades of jet fuels. We show how changes in the reaction conditions
in the catalytic microwave-induce degradation and hydrogenation steps can be utilized to tune
the product selectivity.
Fig. 8.1 Reaction pathways for the conversion of plastics (LDPE) into jet fuels.
8.3 Materials and Methods
8.3.1 Materials
Low density polyethylene (LDPE) (CAS number 9002-88-4) in the form of pellets was
purchased from Sigma-Aldrich Corporation (St. Louis, MO, USA). The density and melting
point of LDPE are 0.925g/cm3 at ambient temperature and 116 °C, respectively. Parent
ZSM-5 (SiO2/Al2O3 Mole Ratio: 50) was purchased from Zeolyst International, USA. Raney
244
Ni 4200 (slurry in water) in an activated form was supplied by Sigma-Aldrich Corporation
(St. Louis, MO, USA). Nickel-Aluminum alloy powder in a non-activated type were used
from Alfa Aesar (Ward Hill, MA, USA).
8.3.2 Catalyst preparation
The activity of parent ZSM-5 was improved by suffering both hydrothermal and calcined
treatments. Under the gentle stirring, parent ZSM-5 powder was added into deionized water
(mass ratio=1) at 60 °C. After addition, the mixture was kept on stirring for 2 h under this
condition. The slurry was then dried at 105 °C till constant weight. The sequential process
was the catalyst calcination: hydrothermally treated ZSM-5 was calcined at 550°C for 5 h in a
muffle furnace. The catalysts were pelletized and sieved to 20 – 40 mesh.
Non-pyrophoric nickel catalyst (referred as NP Ni) and Home-made Raney Ni catalyst was
treated according to the modified method described in the literature.19 NP Ni was prepared
from metallic alloy powders. Under gentle stirring, Ni-Al alloy powders (2 g) were slowly
impregnated with 1.2 wt% NaOH aqueous solution (20 mL) at room temperature. After
addition, the temperature was elevated from room temperature to 80 °C and hold at the
temperature for 30 min. Additional 2 mL of 12 wt% NaOH aqueous solution was added to the
slurry and stirred gently at 80 °C for 30 min for further alkali leaching. Subsequently the
sample was washed to neutrality using distilled water and reserved in water for future
catalytic use.
245
Home-made Raney Ni catalyst was developed using 20 wt% NaOH aqueous solution to
remove Al in the following procedure. 2 g of the above Ni-Al was slowly added into 20 mL
NaOH aqueous solution under gentle stirring. After addition, the slurry was kept on stirring at
80 °C for 1 h. The excess of sodium hydroxide was finally washed with distilled water until
nearly neutral pH was reached. The obtained Raney Ni catalyst was stored in water. In order
to test the catalytic performance of Raney Ni manufactured from the process, Raney nickel
4200 (referred as Raney Ni 4200) purchased in the activated form was used as criterion for
catalyzing the hydrogenation reaction of pyrolysis oils. Raney Ni is notorious for its
pyrophoricity, and it may ignite spontaneously when dried in air. The Raney Ni 4200 slurry
and home-made Raney Ni catalyst were thus dried at 60 °C till constant weight in the
atmosphere of nitrogen to avoid contact with air, prior to the subsequent catalytic test.
8.3.3 Catalytic microwave-assisted degradation of low-density polyethylene (LDPE)
Detailed experimental setting was described in our previous studies.25, 26 Fixed loading of
low-density polyethylene pellets (50 g) for each run were placed in a 500 mL quartz flask
inside the microwave oven (Shanghai, China) by a constant microwave power setting (700
W). 0.25 g of activated carbon powder was used as the absorber for the microwave-assisted
degradation. All reactions of microwave degradation were conducted at the temperature of
350 °C for 20 min until all LDPE pellets were completely vaporized. The volatiles from the
flask passed through a packed bed catalysis reactor which was filled with the modified
ZSM-5 catalyst. As our previous work reported, a mild condition to obtain the high carbon
yield of liquid organics was set at 375 °C 26. In this regard, the catalytic temperature was held
246
constantly at 375 °C. Since the catalytic microwave degradation showed decent results even
in the presence of low catalyst loading 19, 26, low catalyst to feed ratio was used at 0.1 or 0.2
to determine and modify the product selectivity of liquid organics. The catalytic microwave
degradation was duplicated for three times in order to gain abundant liquid organics which
were thereafter combined together for the hydrogenated process.
8.3.4 Hydrogenation of raw organics derived from catalytic microwave degradation
To saturate the liquid organics evolved from catalytic microwave degradation, a closed
reaction system with a stirred stainless batch reactor of the 4592 micro stirred reactor (with a
50 mL vessel) and a 4848 reactor controller from Parr Instrument Company (Moline, IL,
USA) was used. The liquid organics (2 g), the n-heptane (6 g), alongside the catalysts with
intended ratio were loaded into the reactor. Then the reactor was sealed and vented for five
times with hydrogen to get rid of the air present in the vessel. Hydrogen was subsequently
adjusted to reach the set pressure. The automatic controller was employed to control the
temperature and the revolution of stirrer (300 rpm). The pressure inside the reactor was
recorded and the reactions proceeded at a set temperature for the intended time. After the
experiment finished, stirring was stopped and the reactor was rapidly cooled to ambient
temperature. Then, the gas was collected for analysis and the reactor was depressurized.
Consequently the liquid product was filtered to remove catalyst particles.
8.3.5 Analytical techniques
Elemental analysis of LDPE pellets, liquid samples, char, and coke deposited on spent
247
catalysts was conducted using a 2400 Series II CHN/O Elemental Analyzer (PerkinElmer,
USA).
The textural properties of the catalyst were measured in light of N2 adsorption–desorption
(Micromeritics TriStar II 3020 Automatic Physisorption Analyzer). Fresh catalysts were
initially degassed in vacuum at 300 °C for 1 h. The Brunauer–Emmett–Teller equation was
used to calculate the surface area of catalysts by using adsorption data at p/po= 0.05–0.25.
The pore volume was determined using the Barrett–Joyner–Halenda (BJH) method.
The acidity of parent ZSM-5 and well-promoted ZSM-5 catalyst was measured by
temperature-programmed desorption (TPD) of ammonia with a Micromeritics AutoChem II
2920 Chemisorption Analyzer equipped with a PFEIFFER mass spectrometer. These samples
were saturated with NH3 at the room temperature in a flow of 10% NH3 in nitrogen. After
NH3 saturation, the weakly bound NH3 was desorbed prior to the measurement at 120 °C for
1 h with a He flow rate of 25 ml/min. The desorption curve was then obtained at a heating
ramp of 10 °C/min from 120 °C to 550 °C at a He flow rate of 25 ml/min.
Powder X-ray diffraction (XRD) patterns were applied on a Rigaku Smartlab X-ray
diffractometer equipped with a Cu Kα X-ray source, which was operated at 40 kV and 40 mA.
The scattering angle 2θ was changed from 10° to 80°.
The particle size and surface morphology of the samples were evaluated with a scanning
248
electron microscope (SEM, FEI Quanta 200 F).
The chemical composition of liquid organics was characterized and qualified by Agilent
7890A GC-MS (GC–MS; GC, Agilent 7890A; MS, Agilent 5975C) with a DB-5 capillary
column. The GC was first programmed to heat to 45 °C for 3 min followed by heating to
300°C at a rate of 10°C/min. The injection sample size was 1 μL. The flow rate of the carrier
gas (helium) was 0.6mL/min. The ion source temperature was 230 °C for the mass selective
detector. Compounds were identified by comparing the spectral data with that in the NIST
Mass Spectral library. The area percent of changed concentrations of model compounds
obtained from GC/MS results was utilized to predict product concentration in liquid organics.
All the measurements were triplicated to assure reproducibility.
The gaseous product was collected in a 1L Tedlar gas bag and then offline analyzed by an
INFICON 3000 Micro-GC (INFICON Inc., Santa Clara, CA, USA) system with a thermal
conductivity detector (TCD). A standard gas mixture consisting of H2, N2, CH4, CO, CO2,
C2H4, C2H6, and C3H6 was used to calibrate the yield of non-condensable gas. Alkanes and
olefins (>C4) in gas samples were either not detected or negligible in this research. All the
measurements were triplicated to assure reproducibility.
249
8.3.6 Experimental methods and data evaluation
The coke mass was determined by the difference before and after catalytic
degradation. The weight of non-condensable gas was calculated using the following
equation.
ℎ
=
!"#
−
− ℎ
−
(1)
Overall carbon yields of the liquid, gas, and solid products and carbon selectivity of a specific
product were calculated based on the following equations.
B
C
B
C
D
=
H D=
C
E
C
C
C
× 100%
E
(2)
× 100%
E
(3)
Table 8.1 Textural properties of well-promoted ZSM-5 and Raney Ni catalysts.a
SBET
Vpore
Spore
dpore
(m2/g)
(cm3/g)
(m2/g)
nm
Parent ZSM-5
386.9
0.078
55.3
5.7
Well-promoted ZSM-5
396.2
0.097
74.1
5.2
Ni-Al alloy
0.35
0
0
0
NP Ni
157.7
0.048
38.2
5.0
Raney Ni 4200
38.1
0.112
43.0
10.4
Home-made Raney Ni
52.4
0.034
35.8
3.8
Zeolites
Nickels
a
SBET: BET surface area; Vpore: pore volume; Spore: pore surface area; dpore: average pore
size
250
8.4 Results and discussion
8.4.1 Catalyst characterization
In comparison with the characteristics of parent ZSM-5, the BET surface area, pore volume,
and pore surface area were significantly improved by the combined treatments. Furthermore,
ZSM-5 modified by means of the combined treatments resulted in the generation of
secondary porosity (mesoporosity) in the ZSM-5 zeolite matrix as shown in Table 8.1. The
average pore size of the modified ZSM-5 is 5.2 nm, which is very close to naphthalene
diameter (5.5 nm), therefore double-ring aromatics are prone to be adsorbed in the pores.27
Fig. 8.2 NH3-TPD profiles of parent ZSM-5 and well-promoted ZSM-5 catalyst.
Since the overall acid amounts can be inferred from the relative peak areas of the NH3
desorption curves; it was observed that the acidity are comparable (Fig. 8.2). For the
well-promoted ZSM-5 catalyst, the maximum of the peak shifted towards low temperatures,
which accompanied with a decrease in percentage of strong acid sites. The catalyst with the
251
decrease of strong acid sites could reduce catalytic cracking of large molecules into gaseous
molecules, thereby increasing the carbon yield of liquid organics. On the other hand, the
minimum of the peck catalysts was not impacted by the integrated treatments. Hence, the
well-promoted ZSM-5 catalyst used in the catalytic reaction was modified by the combined
treatments and then pelletized and sieved to 20 – 40 mesh.
During alkali leaching, Al content in the Ni-Al alloy powder reacted with NaOH solution. As
expected, Raney Ni primarily consisted of metallic nickel. When the Al/NaOH stoichiometry
is more than 1, insoluble Al(OH)3 was also produced through the reaction between Al and the
water.28 Wherefore NP-Ni derived from Ni-Al alloy treated by insufficient amount of NaOH
is considered as a Ni-Al(OH)3 catalyst. For Raney Ni catalyst, there is also a small amount of
residual hydrated alumina absorbed in the spongy structure of Raney Ni catalyst.28 In this
respect, the textural properties of NP-Ni, Raney-Ni 4200 and home-made Raney Ni,
compared with the parent Ni-Al alloy powder, are outlined in Table 1. BET surface area, pore
volume, and pore surface area of all catalysts were dramatically enhanced by the alkali
treatments. After dissolution of Al component, the BET surface area of home-made Raney Ni
remarkably increased from 0.35 to 52.4 m2/g, which was much higher than that of
as-purchased Raney-Ni 4200 (38.1 m2/g) and other Raney Ni catalysts reported elsewhere.29
The improved BET surface area of home-made Raney Ni could assist the adsorption of
hydrogen molecules for the hydrogenation reaction. Although the NP-Ni has the highest BET
surface area, the catalytic activity of NP-Ni was possibly hindered by hydrated alumina
adsorbed in the structure. In addition, pore volume (0.034 cm3/g) and pore surface area (35.8
m2/g) of for home-made Raney Ni catalyst were also promoted. The average pore size (3.8
nm) of home-made Raney Ni catalyst was identical to the diameter of monocyclic aromatics,
thus monocyclic aromatics are prone to diffusing into the pores.27
252
Fig. 8.3 shows the XRD patterns of the Ni-Al alloy, NP Ni, Raney-Ni 4200 and home-made
Raney Ni catalyst. It can be seen that the XRD patterns of Ni-Al alloy was comprised of the
two categories of Ni3Al2 and Ni3Al domains. As the Al component dissolved by 20 wt. %
NaOH solution, the diffractions regarding metallic Ni were obtained as amorphous nature for
both Raney-Ni 4200 and home-made Raney Ni catalyst. It was found that both Raney Ni
4200 and home-made Raney Ni catalyst mainly presented diagnostic (111), (200), and (220)
diffractions of fcc Ni at 2θ of 44.5, 51.8, and 76.3°, respectively.30 With respect to NP Ni, the
composition of Al(OH)3 was identified based on the peaks assignable to gibbsite and bayerite.
Fig. 8.3 The XRD patterns of the Ni-Al alloy powder, NP-Ni, Raney-Ni 4200, and
home-made Raney Ni catalyst.
The SEM images of Ni-Al alloy, NP-Ni, Raney-Ni 4200 and the home-made Raney Ni
catalyst were displayed in Fig. 8.4. The morphological difference between Ni-Al alloy and
other catalysts were readily visible from Fig. 8.4 (A), (B), (C), and (D). It was noticed that
253
both Raney-Ni 4200 and home-made Raney Ni catalyst were constituted by the typical
fractured and angular particles, which were comparable with other research;31 whilst Ni-Al
alloy showed the intact metallic structure. Furthermore, the small particles of home-made
Raney Ni catalyst were more dispersive than Raney-Ni 4200 particles, reaffirming a higher
BET surface. As shown in Fig. 8.4 (B), NP-Ni was composed of Ni angular particles and
irregularly oriented aluminum crystal-like particles.28
B
AA
D
C
Fig. 8.4 SEM images of the Ni-Al alloy powder (A), NP-Ni (B), Raney-Ni 4200 (C),
and home-made Raney Ni (D).
254
A
B
D
C
Fig. 8.5 Overall carbon yield distribution (C mol%) and carbon selectivity of main chemical
compounds (C mol%): catalytic microwave degradation of LDPE with catalyst to LDPE ratio
0.2 (A, C); catalytic microwave degradation of LDPE with catalyst to LDPE ratio 0.1 (B, D).
8.4.2 Catalytic transformation of LDPE into liquid organics
To produce aviation fuels for use in the aircrafts, the transformation of plastics into the
precursors of jet fuels in the carbon number range from 8 to 16 is necessitated. Based on our
previous study for catalytic microwave pyrolysis of LDPE,26 a mild catalytic temperature set
at 375 °C could obtain the high carbon yield of liquid organics. Interestingly, the low catalyst
to feed ratio can obtain a good result regarding the produce selectivity of liquid organics 26, 32.
255
Henceforth, the catalytic transformation of LDPE volatiles were subjected to the desirable
catalytic temperature by using two ratios (0.2 and 0.1) to concentrate on different grades of
jet fuel precursors. The liquid organics from catalytic microwave degradation of LDPE at the
ratio of 0.2 and 0.1 were 64.41±5.20 wt% and 50.77±3.01 wt%, respectively. There have
never been any results could achieve such high mass yield of liquid organics derived from
catalytic degradation under such low-severity conditions. The overall carbon yield
distribution from catalytic microwave degradation of LDPE is shown in Fig. 8.5 (A) at the
ratio of 0.2 and Fig. 8.5 (B) at the ratio of 0.1. It is found that the share of liquid organics was
the dominant product, occupying 56.32 and 66.18% corresponding to the ratio of 0.2 and 0.1;
whereas both the minor carbon yield of coke and char were approximately 1%, which is as
low as our aforementioned result.26 These results suggest that more catalyst loading could
facilitate catalytic cracking rate to form gaseous compounds. The gas consisted of ethylene,
ethane, hydrogen, and methane, originating from the large extent of cracking and
oligomerization reactions.33
The carbon selectivity of main chemical compounds in the liquid organics from catalytic
microwave degradation of LDPE is depicted in Fig. 8.5 (C) and Fig. 5 (D) with the ratio of
0.2 and 0.1, respectively. It is noted that waxes evolved from direct thermal degradation could
be converted into the lumps of liquid organics through catalytic cracking, cyclization, and
oligomerization reactions under the mild reaction conditions. In general, the chemical
compounds of liquid organics have been identified; and main chemical compounds were also
quantified. Over 50% carbon selectivity of chemical compounds were quantified as shown in
Fig. 8.5 (C) and Fig. 8.5 (D). The total amounts of aromatic hydrocarbons in the ratio of 0.2
were superior to those in the ratio of 0.1, indicating that more catalyst could offer abundant
active sites for the aromatization reactions to form aromatic hydrocarbons. It is observed that
256
mono-cyclic aromatic hydrocarbons in the jet fuel range were preferentially formed over the
well-promoted ZSM-5 catalyst, which could be ascribed to the fast diffusion of the
intermediates toward the active sites inside the micropores. A small amount of double-cyclic
aromatic hydrocarbons including naphthalene and its derivatives was evolved from the
oligomerization and polymerization reactions of mono-cyclic aromatic hydrocarbons.34
It is
also noted that carbon selectivity towards xylenes (C8H10) was 24.69 % and 18.34 %, which
was the dominant part in the liquid organics. Likewise, the large amount of propylbeneze
(C9H12) was also obtained, displaying the second remarkable carbon selectivity under the
both conditions. In addition, catalytic cracking of waxes from thermal degradation of LDPE
also gave rise to certain amounts of aliphatic hydrocarbons; the catalyst could not provide
enough active sites for aromatization and oligomerization reactions when using low catalyst
loading, thereby triggering high amount of aliphatic hydrocarbon under the condition of the
low catalyst loading. Consequently, the results imply that the products gained from catalytic
microwave degradation of LDPE at 375 °C is mostly made up of hydrocarbons in the jet fuel
range (C8 – C16).
8.4.3 Hydrogenation of raw organics derived from catalytic microwave degradation
8.4.3.1 The effect of reaction temperature on the chemical composition of hydrogenated
organics
Since the liquid organics produced by catalytic microwave degradation of LDPE at both the
ratio of 0.2 and 0.1 were principally comprised of C8 - C16 range unsaturated hydrocarbon. In
order to completely convert unsaturated hydrocarbons into alkanes, several variables were
investigated to determine the optimal conditions for the maximum conversion. In the
hydrotreating process, the loss of the liquid organics was ignored if considering the recovery
257
(more than 95 wt%) of the liquid organics could be achieved. Because of the significant
influence of reaction temperature in such system 19, reaction temperature was first used as the
impact factor. The liquid organics derived from catalytic microwave degradation of LDPE
with the presence of different catalyst loadings were separately taken to measure the effect of
reaction temperature.
Due to the distinct concentrations of aliphatic and cyclic hydrocarbons in the liquid organics
from catalytic microwave degradation of LDPE by using different catalyst loadings, each
sample should be investigated with respect to several variables. Herein, the raw organics
from catalytic microwave degradation with the catalyst to LDPE ratio of 0.2 were initially
subject to the hydrotreating processes. The overall carbon yields of hydrogenated organics
with respect to the raw plastics were approximately 54% without the calculation of n-heptane
carbon yield. The independent factors (such as reaction temperature and initial pressure) did
not have an impact on the carbon yield of hydrogenated organics. It is manifested that carbon
loss reactions (e.g. hydrocracking and decarbonylation) rarely took place under these mild
conditions.
258
Fig. 8.6 Chemical composition of hydrotreated organics on the basis of reaction temperature.
Reaction condition: Reactant: raw organics from catalytic microwave degradation of LDPE
over 20 wt% ZSM-5 catalyst; Initial pressure, 500 psi; Raney Ni 4200 catalyst, 10 wt% with
respect to reactant mass; Reaction time, 2 h.
It was also observed that either the Raney Ni 4200 catalyst to reactant ratio conducted at 0.05
or reaction time set for 1 h slightly converted the unsaturated hydrocarbons into alkanes.
These results indicate that the enhancement in the two variables (catalyst to reaction ratio and
reaction time) led toward a parallel augment in the yields of saturated hydrocarbons. The
insufficient saturated alkanes are due in part to the mild reaction conditions. Since these high
amounts of unsaturated hydrocarbon remained in the hydrotreated organics, which cannot be
straightforward used as jet fuels; further efforts for converting unsaturated hydrocarbons into
saturated hydrocarbons, containing both aliphatic and cyclic alkanes, should be made. Based
259
on the previous comparison, the catalyst to reactant ratio of 0.1 and reaction time for 2 h were
feasible for the testing experiments of reaction temperature. The chemical compounds of
hydrotreated organics are elucidated as a function of reaction temperature as shown in Fig.
8.6. The optimal condition for maximum yields of saturated hydrocarbons was found to be at
the reaction temperature of 250 °C. The corresponding amounts of aliphatic and cyclic
alkanes reached at the highest selectivity, while the amount of aromatic hydrocarbon dropped
to the lowest. Thus, these augments of aliphatic and cyclic alkanes took place mainly at the
expense of the aliphatic olefins and aromatic hydrocarbon because their share declined.
On the other hand, the liquid organics from catalytic microwave degradation with the catalyst
to LDPE ratio of 0.1 were thereafter used as the feedstock in the hydrotreating processes. The
overall carbon yields of hydrogenated organics regarding the raw plastics under the following
conditions were ~ 63% without the calculation of n-heptane carbon yield.
It was also found
that the reaction time had a significant impact on the hydrotreating processes. In this sense,
the reaction temperature was set at 2 h for all experiments. The total amounts of both
aliphatic alkanes and cyclic alkanes progressively increased with the increasing reaction
temperatures from 150 to 200 °C. Nevertheless, the optimal result (close to 30% selectivity of
aromatic hydrocarbons plus hydro-aromatic hydrocarbons remained in the hydrotreated
organics) cannot meet the specifications of conventional jet fuels, especially for advanced jet
fuels. Thus more catalyst loading should be employed to obtain higher valuable fuels that can
directly be used as conventional jet fuels or drop-in fuels. Table 8.2 shed light on products
260
distribution and partial alkanes’ carbon selectivity as a function of reaction temperature with
the catalytic to reactant ratio of 0.1. It was observed that the high yields of aliphatic and
cyclic alkanes were achieved even at the mild reaction temperature (150 °C), which implies
that the addition reaction rate was relatively high at the low reaction temperature. The share
of cycloalkanes grew to 57.75% as the reaction temperature increased to 200 °C, indicating
that elevated reaction temperature favored the hydrogenation rate of aromatic hydrocarbons.
The total amount of cycloalkanes in the hydrotreated organics at 200 °C was equal to the
content (52.80%) in JP-5.35 It was also noticeable that the content of aliphatic alkanes
appeared to generally increase to 40.79% at 200 °C, whose amount was a bit higher than the
content (30.8%) in conventional JP-5. Unlike the saturated hydrocarbons, the total amounts
of hydro-aromatic hydrocarbons and aromatic hydrocarbon distantly declined, particularly
from 175 to 200 °C. There were no hydro-aromatic hydrocarbons and aromatic hydrocarbons
detected at 200 °C, which meets the limited ceiling of aromatic hydrocarbons (15.9%) in the
conventional jet fuels. This remarkable drop in the amount of aromatics is a positive impact
on the quality of JP-5 fraction. In the gaseous fraction, unreacted hydrogen was detected
along with minor volume of small hydrocarbons, confirming that the experiments were not
carried out under hydrogen starved conditions. Consequently, the obtained liquid products
conducted under the mild reaction temperature can be potentially used as alternatives for the
formulation of JP-5 navy fuel.
261
Table 8.2 Products distribution and partial alkanes’ carbon selectivity as a function of
reaction temperature.a
Reaction temperature ( ºC)
150
175
200
Aliphatic alkanes
40.27
40.22
40.79
Cycloalkanes
49.21
51.29
57.75
Hydro-aromatic hydrocarbons
1.00
0.34
0
Aromatic hydrocarbons
7.73
4.83
0
Others
1.79
3.32
1.46
1,4-dimethylcyclohexane
-
3.17
-
1,3-dimethylcyclohexane
10.35
7.20
11.09
1,2-dimethylcyclohexane
1.13
2.08
4.42
Ethylcyclohexane
2.97
8.88
2.75
Octane
1.29
1.07
1.47
1,2,4-trimethylcyclohexane
1.99
1.67
4.43
Propylcyclohexane
11.93
9.94
15.44
Nonane
1.60
1.31
1.70
Hexahydroindan
3.41
3.23
3.58
Decalin
0.80
0.74
0.68
Decane
2.52
2.32
2.53
Undecane
3.67
3.63
3.42
Dodecane
3.48
3.80
3.33
Tridecane
2.42
2.39
2.82
Tetradecane
2.27
1.87
2.88
Pentadecane
2.65
2.78
2.68
Hexadecane
2.48
2.07
2.87
Overall selectivity (% in area)
Alkanes selectivity (C mol%)
a Reaction
condition: Reactant: raw organics from catalytic microwave degradation of LDPE
262
over 10 wt% ZSM-5 catalyst; Initial pressure, 500 psi; Raney Ni 4200 catalyst, 10 wt% with
respect to reactant mass; Reaction time, 2 h.
The carbon selectivity of specific alkanes including aliphatic and cyclic alkanes were also
substantially affected by the reaction temperature. The carbon selectivity of cycloalkanes
significantly
increased
as
the
reaction
temperature
increased.
Meanwhile,
1,
3-dimethylcyclohexane (C8H16) and propylcyclohexane (C9H18) gradually increased to 11.09%
and 15.44%, indicating that increasing reaction temperature presumably accelerated the
hydrogenation of the corresponding aromatic hydrocarbons. The carbon selectivity of other
cyclic alkanes (e.g. decalin) was slightly impacted by the increment of reaction temperature.
This fact was possibly ascribed to complete conversion of the aromatic hydrocarbons into
corresponding cycloalkanes. Likewise, the reaction temperature solely played a slight role in
the carbon selectivity of aliphatic alkanes. The carbon selectivity of aliphatic alkanes, such as
octane (C8H18) and nonane (C9H20), showed a steady fluctuation as the reaction temperature
grew. These outcomes imply that the stability of aliphatic alkanes was attributed to entire
addition reaction of aliphatic olefins, which could occur at the relatively low reaction
temperature.33
8.4.3.2 The effect of initial pressure on the chemical composition of hydrogenated
organics
In our aforementioned results, the raw organics from catalytic microwave degradation of
LDPE by using the 10 wt% catalyst (with respect to LDPE masss) can be readily converted
into the saturated hydrocarbons under mild reaction condition. To further transform the raw
organic from catalytic microwave degradation of LDPE with the use of 20 wt% catalyst (with
regard to LDPE mass) into saturated hydrocarbons, the effect of initial pressure on the
263
hydrotreated organics was implemented. The initial pressure within the 500 – 900 psi range
was first investigated working at 200 °C for 2 h and using the catalyst (Raney Ni 4200) to
reactant (raw organics) ratio of 0.05. The overall carbon yields of hydrogenated organics as a
function of initial pressure were around 54% after removing the carbon yield of n-heptane.
Because of the low catalyst loading, there were still a large amount (over 30 wt%) of
aromatic hydrocarbons remaining in the hydrotreated organics, even under the highest initial
pressure conducted at 900 psi. Wherefore more catalyst loading (10 wt, with respect to
reactant mass) of Raney Ni 4200 were be applied in the hydrotreating processes.
Fig. 8.7 Chemical composition of hydrotreated organics with respect to initial pressure.
Reaction condition: Reactant: raw organics from catalytic microwave degradation of LDPE
over 20 wt% ZSM-5 catalyst; Reaction temperature, 200 ºC; Raney Ni 4200 catalyst, 10 wt%
with respect to reactant mass; Reaction time, 2 h.
Fig. 8.7 shows the yields by categories obtained in these hydrotreated reactions. Seemingly,
the arguments of initial pressure enhanced the extent of hydrocycloaddition reactions toward
264
cycloalkanes, obtaining the highest amount in this fraction (84.32%) at 900 psi of hydrogen
pressure. Such high content of cycloalkanes and the aliphatic alkanes in the hydrotreated
organics could be directly used as the additives in the conventional jet fuels or the
replacements of high energy-density jet fuels (e.g. JP-10 and RJ-5). At the initial pressure of
900 psi, the most noticeable result was a drop in the total aromatic hydrocarbon contents. The
enhancement of initial pressure is supposed to conversion of aromatic hydrocarbons into
cycloalkanes and hydro-aromatic hydrocarbons. Besides, the yield of hydro-aromatic
hydrocarbons experienced a gradual decline, likely because of hydro-aromatic hydrocarbons
acting as the intermediates, which could be thereafter transformed into cycloalkanes. Hence
the increase in cycloalkanes took place at the expense of aromatic hydrocarbons and
hydro-aromatic hydrocarbons. In contrast, the yield of aliphatic alkanes slightly grew to
8.65%, even at the highest initial pressure (900 psi). In addition, no aliphatic olefins were
detected regardless of the initial pressure, which is indicative of its complete addition
reactions of aliphatic olefins. Therefore these conditions allowed for the entire removal of
aliphatic olefins. These outcomes imply that a high initial pressure contributed to more
hydrogenation reactions occurring in the reactor, because more volume of hydrogen was
provided on the surface of Raney Ni catalyst.
265
Fig. 8.8 Carbon selectivity of main alkanes with respect to initial pressure. Reaction condition:
Reactant: raw organics from catalytic microwave degradation of LDPE over 20 wt% ZSM-5
catalyst; Reaction temperature, 200 ºC; Raney Ni 4200 catalyst, 10 wt% with respect to
reactant mass; Reaction time, 2 h.
The carbon selectivity of cycloalkanes significantly increased by using the 10 wt% Raney Ni
4200 catalyst, especially for 1, 3 - dimethylcycloalkanes, as the initial pressure was elevated
(Fig. 8.8); while the carbon selectivity of ethylcyclohexane and propylcyclohexane
substantially went up in the range of 500 – 900 psi. Even if naphthalene can be converted into
decalin under these conditions,19 the carbon selectivity of decalin showed a declined tendency.
The reduced carbon selectivity of decalin was attributed to hydroisomerization of decalin to
1-methyloctahydro-1H-indene (C10H18)
36.
Likewise, the carbon selectivity of main linear
alkanes also increased with the augment of initial pressure, yet showing a slight trend. The
carbon selectivity (above 1 %) of C10 – C12 linear alkanes were the dominant aliphatic
alkanes in the hydrotreated organics. In terms of gaseous results at 900 psi, a trace volume of
small-chain
hydrocarbons
were
found,
affirming
266
that
the
hydrocracking
and
hydroisomerization reaction have taken place in the processes. Therefore, improving the
initial pressure can principally enhance the hydrogenation reaction rate; small extent of
hydrocracking and hydroisomerization reactions also occurred at the high initial pressure.
8.4.3.3 The effect of reaction time on the chemical composition of hydrogenated organics
Apart from reaction temperature and initial pressure, reaction time was another crucial factor
that affected product distribution and alkanes’ carbon selectivity. Since the raw organics from
catalytic microwave degradation of LDPE by using 10 wt% ZSM-5 catalyst (with respect to
LDPE mass) can be readily transformed, the raw organics over 20 wt% ZSM-5 catalyst were
thus implemented to determine the influence of reaction time. Interestingly, the overall
carbon yields of hydrogenated organics based on the reaction time were ~ 54%, indicating
that prolonged reaction time could not initiate any carbon loss reactions. Although the
reaction was conducted for 4 h, a large amount of aromatic hydrocarbons were remained in
the hydrotreated organics due to the presence of low catalyst loading. By contrast, the product
distribution (with 10 wt% Raney Ni 4200 catalyst) and partial alkanes’ carbon selectivity in
light of reaction time are depicted in Table 8.3. In general, the resulting organics prominently
shifted towards the cycloalkane (26.47 – 75.99%) with increase of reaction time, indicating
that addition reactions were enhanced because of longer reaction duration. The total amount
of aliphatic alkanes slightly increased to 8.04% at prolonged reaction time (4 h). The total
amount of hydro-aromatic hydrocarbons distantly declined; likewise, the yield of aromatic
hydrocarbons showed a downward tendency, ranging from 54.16 to 9.30%. As prolonged
reaction time could give rise to hydrocycloaddition reactions, resulting in the decrease of both
267
aromatic hydrocarbons and hydro-aromatic hydrocarbons. The outcome indicates that the
hydrotreated organics (at 4 h) within ~ 85% selectivity of alkanes are suitable for the
utilization as high energy-density jet fuels.
Table 8.3 Products distribution and partial cycloalkanes carbon selectivity on the basis of
reaction time.a
Reaction time (h)
1
2
3
4
Aliphatic alkanes
5.37
7.14
7.46
8.04
Cycloalkanes
26.47
56.28
71.83
75.99
Hydro-aromatic hydrocarbons
13.61
11.12
6.75
5.53
Aromatic hydrocarbons
54.16
24.07
12.62
9.3
0.4
1.39
1.36
1.14
1,4-dimethylcyclohexane
1.12
4.23
2.33
4.89
1,3-dimethylcyclohexane
5.40
9.14
20.95
16.56
1,2-dimethylcyclohexane
1.61
4.79
2.97
3.10
Ethylcyclohexane
2.76
3.21
3.60
3.43
Octane
0.45
0.71
0.51
0.52
1,2,4-trimethylcyclohexane
0.21
0.51
2.93
3.59
Propylcyclohexane
7.29
13.94
18.92
19.47
Nonane
0.41
0.42
0.47
0.49
Hexahydroindan
0.50
1.05
0.94
0.86
Decalin
1.12
5.95
5.28
5.41
Decane
0.68
0.95
1.25
1.33
Undecane
1.56
-
2.57
2.58
Dodecane
1.42
2.19
1.98
2.19
Overall selectivity (% in area)
Others
Alkanes selectivity (C mol%)
268
Tridecane
0.47
0.51
0.69
0.67
Tetradecane
0.37
0.42
0.45
0.45
Pentadecane
0.65
0.76
0.69
0.72
Hexadecane
0.53
0.66
0.57
0.60
a Reaction
condition: Reactant: raw organics from catalytic microwave degradation of LDPE
over 20 wt% ZSM-5 catalyst; Initial pressure, 500 psi; Raney Ni 4200 catalyst, 10 wt% with
respect to reactant mass; Reaction temperature, 200 ºC;
The
carbon
selectivity
toward
cycloalkanes
Reaction time, 2 h.
increased,
especially
toward
dimethylcyclohexanes, as the reaction time was elevated; whilst the carbon selectivity of
decalin and hexahydroindan decreased from the period of 2 - 4 h. That is because naphthalene
and indene could be completely converted into decalin and
hexahydroindan under the
condition of 2 h. Prolonged time could render the cycloalkanes to be hydrocracked and
hydroisomerizated.36 For the gas result from the reaction for 3 and 4 h, trace volume of
hydrocarbons was detected, which implies that the hydrocracking and hydroisomerization
reactions have taken place under these conditions.
As such, the carbon selectivity of linear
alkanes quantified was improved due to the increase of reaction time. Prolonging reaction
time to 4 h could make some cyclic alkanes be hydrocracked into aliphatic alkanes, thereby
improving the linear alkanes’ carbon selectivity. Wherefore these results demonstrate that
prolonging reaction can principally enhance the hydrogenation reactions; hydrocracking and
hydroisomerization reactions could also occur after 2 h reaction time in the process.
269
8.4.3.4 The effect of catalyst loading on the chemical composition of hydrogenated
organics
Based on the previous results, the reaction time set at 2 h are suitable for hydrogenation of
raw organics from catalytic microwave degradation of LDPE by using 20 wt% ZSM-5
catalyst.
Hence the reaction time of 2 h was used to investigate the variable of catalyst to
reactant ratio for all experiments. The overall carbon yields of hydrogenated organic with
regard to catalyst loading fluctuated at 54%. Table 8.4 shows representative results of
hydrogenated organics regarding the chemical compounds as a function of catalyst to reactant
ratio. The catalyst to reactant ratio had a vital influence in the yield of aliphatic and cyclic
alkanes. The increase in the amount of catalyst loading led toward an enhancement in the
extent of hydrogenation reactions, augmenting the share of aliphatic and cyclic alkanes, and
decreasing those of aromatic hydrocarbons especially when a ratio of 0.2 was employed. The
sufficient catalysts (catalyst to reactant ratio of 0.2) applied could give rise to a premium
result, achieving over 10% selectivity toward aliphatic alkanes and approximately 80%
toward cycloalkanes, which displayed a more advanced jet fuel.
Table 8.4 Products distribution and partial cycloalkanes carbon selectivity with respect to
catalyst to reactant ratio. a
Catalyst to reactant ratio
0.05
0.1
0.15
0.2
Aliphatic alkanes
7.41
7.14
10.31
10.75
Cycloalkanes
23.83
56.28
72.58
79.72
Hydro-aromatic hydrocarbons
14.44
11.12
4.56
1.74
Aromatic hydrocarbons
51.8
24.07
10.47
6.34
Overall selectivity (% in area)
270
Others
2.52
1.39
2.08
1.44
1,4-dimethylcyclohexane
1.12
4.23
2.22
3.87
1,3-dimethylcyclohexane
4.91
9.14
24.28
17.06
1,2-dimethylcyclohexane
0.65
4.79
2.10
5.22
Ethylcyclohexane
8.29
3.21
3.32
3.24
Octane
0.68
0.71
0.62
0.87
1,2,4-trimethylcyclohexane
0.22
0.51
2.02
2.54
Propylcyclohexane
5.10
13.94
17.10
19.20
Nonane
0.41
0.42
0.45
0.51
Hexahydroindan
0.56
1.05
0.92
1.69
Decalin
1.47
5.95
5.78
6.04
Decane
0.61
0.95
1.00
1.13
Undecane
0.11
-
0.15
0.34
Dodecane
1.31
2.19
2.06
2.38
Tridecane
0.46
0.51
0.49
0.51
Tetradecane
0.36
0.42
0.37
0.40
Pentadecane
0.61
0.76
0.67
0.68
Hexadecane
0.51
0.66
0.58
0.59
Alkanes selectivity (C mol%)
aReaction
condition:
Reactant: raw organics from catalytic microwave degradation of
LDPE over 20 wt% ZSM-5 catalyst; Initial pressure, 500 psi; Raney Ni 4200, 10 wt% with
respect to reactant mass; Reaction temperature, 200 ºC; Reaction time, 2 h.
The effect of catalyst to reactant ratio on partial alkanes’ carbon selectivity is also explained
in Table 8.4. It was observed that catalyst loading had a noticeable influence in the carbon
selectivity of mono-cyclic alkanes, including dimethylcyclohexanes and propylcyclohexane.
These results suggest that mono-cyclic aromatic hydrocarbons were prone to be converted
into corresponding alkanes if over 10 wt% catalyst (with regard to reactant ratio) was
271
introduced. Since more adequate catalysts were applied; the cycloalkanes could continue
undergoing the hydroisomerization at the spare active sites to form aliphatic alkanes. It can be
seen that the carbon selectivity of both decane and dodecane increased as more catalysts (20
wt% Raney Ni 4200) were used. Consequently, more catalyst loading provided more active
sites for the hydrocycloaddition and hydroisomerization reactions, resulting in more cyclic
and aliphatic alkanes in the hydrotreated organics.
8.4.3.5 The effect of catalysts selection on the chemical composition of hydrogenated
organics
According to investigation of aforementioned variables (reaction temperature, initial pressure,
reaction time, and catalyst to reactant ratio), the chemical composition of hydrotreated
organics derived from the reactions conducted at the mild and suitable conditions in the
presence of various catalysts (NP Ni, Raney Ni 4200, and home-made Raney Ni) is depicted
in Table 8.5. The overall carbon yields of hydrogenated organics in the presence of diverse
Raney Ni catalysts were approximately 54%, which suggests that the catalyst species did not
favor the hydrocracking of large hydrocarbons to small molecules. Comparing with the
overall product distribution by using as-purchased Raney Ni 4200 catalyst, the result over NP
Ni and home-made Raney Ni was superior to Raney Ni 4200 for the production of alkanes
under the same condition. It was observed that 76.08% selectivity toward cycloalkanes and
8.14% selectivity toward aliphatic alkanes was achieved over home-made Raney Ni, which
was even better than the result using Raney Ni 4200 when the reaction time was set for 4 h.
Total amounts of hydro-aromatic hydrocarbons and aromatic hydrocarbons were less than
272
15 %, which did not exceed the specifications of conventional jet fuels (e.g. JP-5).
Accordingly, the home-made Raney Ni is the optimal catalyst that could be used for the
production of high energy-density jet fuels under the mild conditions.
Table 8.5 Products distribution and partial alkanes’ carbon selectivity in the presence of
various catalysts.a
Catalyst categories
Raney Ni
NP Ni
Home-made Raney Ni
4200
Overall selectivity (% in area)
Aliphatic alkanes
7.14
8.04
8.14
Cycloalkanes
56.28
70.99
76.08
Hydro-aromatic hydrocarbons
11.12
5.53
4.2
Aromatic hydrocarbons
24.07
14.3
9.81
Others
1.39
1.14
1.77
1,4-dimethylcyclohexane
4.23
2.34
2.52
1,3-dimethylcyclohexane
9.14
15.36
24.25
1,2-dimethylcyclohexane
4.79
4.84
2.51
Ethylcyclohexane
3.21
3.20
3.33
Octane
0.71
0.82
1.18
1,2,4-trimethylcyclohexane
0.51
1.53
2.64
Propylcyclohexane
13.94
16.82
17.24
Nonane
0.42
0.59
0.63
Hexahydroindan
1.05
1.31
1.60
Decalin
5.95
5.50
5.69
Decane
0.95
1.03
1.09
-
0.21
0.22
Alkanes selectivity (C mol%)
Undecane
273
Dodecane
2.19
2.13
2.37
Tridecane
0.51
0.52
0.51
Tetradecane
0.42
0.40
0.44
Pentadecane
0.76
0.85
0.81
Hexadecane
0.66
0.87
0.78
a
Reaction condition:
Reactant: raw organics from catalytic microwave degradation of
LDPE over 20 wt% ZSM-5 catalyst; Initial pressure, 500 psi; Catalyst, 10 wt% with respect
to reactant mass; Reaction temperature, 200 ºC; Reaction time, 2 h .
It was also found that the carbon selectivity of both dimethylcyclohexanes and
propylcyclohexane were maximum in the presence of home-made Raney Ni. Other
cycloalkanes such as ethylcyclohexane and trimethylcyclohexane also showed the largest
carbon selectivity. Therefore home-made Raney Ni should be used to assure the production of
cycloalkanes. Of the catalysts, home-made Raney Ni displayed the highest carbon selectivity
of aliphatic alkanes. These results indicate that the addition reactions of aliphatic olefins
within hydrogen could be absolutely carried out over home-made Raney Ni, thereby forming
aliphatic alkanes. To be more specific, the carbon selectivity towards octane, decane, and
dodecane was all maximum when using the home-made Raney Ni as the catalyst. These
outcomes imply that employing home-made Raney Ni as the catalyst could obtain the highest
carbon selectivity of cyclic and aliphatic alkanes.
8.4.3.6 Process robustness: the recyclability of the Raney Ni catalyst
The conversion of plastics (LDPE) was investigated in the aforementioned reaction
conditions to assure the production of various grades of jet fuels. Raw organics deliberately
274
chosen were from catalytic microwave degradation of LDPE over 20 wt% ZSM-5 catalyst
duet to their high aromatic hydrocarbons content. Irrespective of conversion extent of
aromatic hydrocarbons to cycloalkanes, an appreciable reaction condition was employed to
determine the process robustness and reuse of the Raney Ni catalyst. That is because the
catalyst reuse is of vital importance to a heterogeneous process.37 In the reaction medium, the
catalyst was first recovered by filtration and then washed. After drying at 60 ºC for 8 h, the
used catalyst (Raney Ni 4200) was reloaded in a fresh reaction medium. The results of three
such consecutive recycling runs are demonstrated in Fig. 8.9.
The carbon yields of
hydrogenated organics regarding to raw plastics were all about 54%. Over 50% selectivity
toward cycloalkanes and approximately 7% selectivity toward aliphatic alkanes were attained,
which were equal to the results from the first run using the fresh catalyst. Hence, there was no
any loss of catalytic activity in the second recycling run, showing that the resilience and
durability of Raney Ni catalyst was decent. However, some loss of catalytic activity was
noticed when the third and fourth recycling runs were conducted, which could originate from
the deactivation of the catalyst caused by coke formation. Notwithstanding a small decrease
in the conversion of unsaturated hydrocarbons into saturated alkanes after the third recycling
run, the catalytic system is acceptably reusable because of no any loss of catalytic activity at
the second recycling run.
275
Fig. 8.9 Chemical composition of hydrotreated organics with respect to recycle times.
Reaction condition: Reactant: raw organics from catalytic microwave degradation of LDPE
over 20 wt% ZSM-5 catalyst; Reaction temperature, 200 ºC; Initial pressure, 500 psi; Raney
Ni 4200 catalyst, 10 wt% with respect to reactant mass; Reaction time, 2 h.
8.5 Conclusions
This contribution has demonstrated the feasibility of combined catalytic conversion of
plastics to various grades of jet fuels. The microwave-induced degradation of plastics
(LDPE) at moderate reaction condition principally manufactured unsaturated
hydrocarbons with aromatic hydrocarbons as major components. The variation of
well-promoted ZSM-5 catalyst loading (10 or 20 wt%) in the catalytic microwave
degradation could modify the product distribution of aliphatic and cyclic hydrocarbons,
276
which could allow the selection of straight-chain and cyclic alkanes in the following
process. The overall carbon yields of raw organics derived from catalytic microwave
degradation over 10 wt% and 20 wt% ZSM-5 catalyst were 66.18% and 56.32%,
respectively; while the carbon yields of both coke and char were approximately 1%.
In the hydrotreating process, the raw organics was hydrogenated in the n-heptane
medium; and diverse variables (reaction temperature, initial pressure, reaction time,
catalyst to reactant ratio, and catalyst selection) were taken consideration to
manufacture various grades of jet fuels by using the two classes of raw organics as
feedstock. It was found that in the presence of 10 and 20 wt% ZSM-5 catalyst the
overall carbon yields of hydrogenated organics regarding raw plastics could reach
approximately 63 and 54%, respectively. The changes of hydrogenation reaction
condition did not have an impact on the carbon yields of hydrogenated organics. For
the production of JP-5 navy fuel, the raw organic from catalytic microwave
degradation in the presence of 10 wt% ZSM-5 catalyst were converted into such grade
of jet fuel by using 10 wt% Raney Ni catalyst at the reaction temperature of 200 ºC for
2 h. The total amounts of aliphatic alkanes (40.79%) and cycloalkanes (57.75%) in the
jet fuel range were obtained under the mild reaction condition. Therefore, the
hydrotreated organics can be potentially used as the alternatives for JP-5 navy fuel or
additives in Jet A, Jet A1, and JP-8 to improve their densities. To produce other grades
of high-energy density jet fuels, the raw organics from catalytic microwave
degradation over 20 wt% ZSM-5 catalyst were transformed as a function of various
277
variables. Up to 84.32% selectivity toward cycloalkanes and 8.65% selectivity toward
aliphatic alkanes were gained at the initial pressure of 900 psi. On the other hand, the
result by using the home-made Raney Ni as the catalyst at 500 psi in the hydrotreating
process was comparable with the optimal result conducted at 900 psi using the Raney
Ni 4200 as the catalyst. The Raney Ni catalyst was thus applicable and also
appreciably reusable in the hydrotreating system. The hydrogenated organics within
large amounts of cycloalkanes could be regarded as the replacements of high
energy-density jet fuels, such as RJ-5 and JP-10. From this perspective, the catalytic
conversions of plastics by using inexpensive catalysts deliver a novel and feasible
pathway in a biorefinery, specifically targeting different grades of jet fuels.
278
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281
CHAPTER NINE
OPTIMIZING CARBON EFFICIENCY OF JET FUEL RANGE
ALKANES FROM CELLULOSE CO-FED WITH
POLYETHYLENE VIA CATALYTICALLY
TANDEM PROCESSES
9.1 Abstract
Enhanced carbon yields of renewable alkanes for jet fuels were obtained through a novel
pathway from co-feeding of cellulose and low-density polyethylene (LDPE). The consecutive
processes proceeded via the catalytic microwave-induced pyrolysis over well-promoted
ZSM-5 catalyst and hydrogenation process by using home-made Raney Ni catalyst. It was
found that parent ZSM-5 modified by hydrothermal and calcined treatments resulted in the
increase of surface area as well as the formation of more mesopores. Interestingly, the
well-promoted ZSM-5 catalyst had high selectivity toward C8 – C16 aromatic hydrocarbons in
the co-feed catalytic microwave pyrolysis. The raw organics with improved carbon yield (~
44%) were more principally lumped in the jet fuel range at the catalytic temperature of
375 °C with the LDPE to cellulose ratio of 0.75. As the Ni-Al alloy dissolved by 20 wt.%
NaOH solution, the BET surface area, pore volume, and pore surface area of home-made
Raney Ni catalyst were appreciably improved. Based on XRD analysis, the diffractions
regarding metallic Ni were achieved as amorphous nature for home-made Raney Ni catalyst.
SEM analysis confirmed that the home-made Raney Ni catalyst was constituted by the typical
282
fractured and angular particles, and the small particles of home-made Raney Ni catalyst were
dispersive. The home-made Raney Ni catalyst was assayed for hydrogenation of diverse
organics species derived from co-feed catalytic microwave pyrolysis under a low-severity
condition. It was observed that the four species of raw organics in the n-heptane medium
were almost completely converted into saturated hydrocarbons. The overall carbon yield
(with regards to co-reactants of cellulose and LDPE) of hydrogenated organics that mostly
match jet fuels was sustainably enhanced, reaching above 39%. Meanwhile, ~ 90%
selectivity toward renewable alkanes for jet fuels was attained. These enhanced hydrogenated
organics with high amounts of renewable cycloalkanes can be potentially used as
high-density jet fuels or additives for blending with commercial jet fuels.
Keywords: Bio-jet fuels; modified ZSM-5; home-made Raney Ni; co-feed catalytic
microwave pyrolysis; hydrogenation; cycloalkanes
9.2 Introduction
Municipal solid waste (MSW) has been found as one of the major challenges in the face of
the modern world;1 in particular, the sustainable disposal management of MSW was
intensively appealed in the industrialized and developing countries.2 MSW is comprised of
various categories of materials. Among it, biomass co-existing with plastics make up a
majority of the composition.3 Yet, the conventional methods for the disposal of MSW takes
away valuable land and causes numerous environmental issues.3 More attentions are paid to
energy-efficient, eco-friendly and economically sound technologies for the disposal of MSW.
Because discarded MSW represents a tremendous energy source; MSW-to-energy
283
technologies not only mitigates negative impacts on the environment, but also provides
sustainable energy. Catalytic fast pyrolysis (CFP) is recognized as the most prevailing and
promising pathway in a single process.4-8 Zeolite-based catalysts (e.g. ZSM-5) have been
proved as the most efficient catalysts to produce important petrochemicals.7, 9, 10
However, even in the presence of ZSM-5 catalyst, catalytic pyrolysis of lignocellulose alone
only generated low carbon yields (10 – 30%) of aromatic hydrocarbons; large amounts of
solid residues (carbon yields usually above 30%) were also achieved in the process.8, 11-13 It is
plagued by the high yield of coke deposited on the catalyst. That is because the coke could
rapidly deactivate the catalyst and further reduce the lifetime, triggering the catalytic process
to be impractical. These aforementioned phenomenon is principally due to the intrinsically
oxygen rich nature and hydrogen deficiency of biomass.11, 13-15 It was found that the hydrogen
to carbon effective (H/Ceff) ratio plays a vital role in the efficiency for converting biomass to
biofuel and coke formation.6, 16, 17 Thus the hydrogen-deficient (H/Ceff usually less than 0.3)
biomass produce low yields of petrochemicals, but large amounts of coke when they were
converted over zeolite-based catalysts.7, 10
To improve the carbon efficiency of aromatic production and reduce the formation of coke,
the incorporation of hydrogen-enriched co-reactant in the catalytic pyrolysis of biomass could
mitigate these problems. It is mostly likely that co-fed hydrogen-rich feedstock in the
catalytic pyrolysis of biomass could change reaction mechanism of oxygen removal from
biomass by substituting decarbonylation and decarboxylation with deoxygenation.8, 16, 18, 19 It
284
is widely known that waste plastics represent a cheaper and abundant hydrogen source,7, 10, 13,
15, 20
which were utilized
to improve carbon yield of aromatic production and lower the
coke formation in the co-feed catalytic pyrolysis. For example, polyethylene accounts for up
to 40% of gross waste plastics in MSW.21 Accordingly, the co-feeding of biomass with waste
plastics from MSW in catalytic pyrolysis is beneficial for the environment and energy
recapture.
On the other hand, the co-feeding of biomass with plastics in catalytic pyrolysis commonly
gave rise to mono-ring aromatic hydrocarbons with low carbon number (C6 – C8).15, 17, 19, 20
These hydrocarbons with low carbon numbers cannot suit current jet fuels within C8 - C16
hydrocarbons. Linear-chain and branched-chain alkanes predominantly make up C8 - C16
hydrocarbons in the current jet fuels.22, 23 By contrast, cycloalkanes are compact molecules
within robust ring strain and can be burned cleanly with high heats of combustion.24-26 As
expected, jet fuel range cycloalkanes should be synthesized and added into commercial jet
fuels (e.g. Jet A and JP-8). It was noteworthy that aromatic hydrocarbons are prone to be
transformed into cycloalkanes by hydro-cycloaddition reactions under a mild reaction
condition.27, 28 To pursue renewable cycloalkanes, highly desirable aromatic hydrocarbon (C8
– C16) as the precursors should be synthesized in co-feed catalytic pyrolysis, rather than
mono-ring aromatic hydrocarbons with low carbon numbers.
Parent ZSM-5 zeolite used for aromatic production with low carbon number are not suitable
for the production of jet fuel range aromatics in the catalytic pyrolysis. Nonetheless,
285
moderate treatments were employed to tailor zeolite properties for modifications of porosity
and acidity, which is essential for improving product selectivity of C8 – C16 aromatic
hydrocarbons in our previous study.29,
30
The mild hydrothermal and calcined conditions
favored the generation of mesopores for the diffusion of double-ring aromatics, and catalytic
sites for specific efficiency. In this regard, we have developed a tandem process to produce
renewable cycloalkane for jet fuels derived from lignocellulosic biomass individually, albeit
achieving a low carbon yield (24.68%).27 It is manifested that microwave-assisted pyrolysis
technology is one of the most promising methods for enhancing and accelerating chemical
reactions due to effective heat transfer through microwave irradiation.31 In comparison with
conventional pyrolysis, microwave-assisted pyrolysis encloses the potentials of fast and
selective heating, easy control of reaction conditions, low reaction temperatures and energy
requirements.32,
33
In the hydrotreating process, Raney-type Ni is widely employed as a
versatile catalyst for reductive conversions of organic compounds.34 It was discerned that the
catalytic activities of home-made Raney Ni catalyst used was superior to as-purchased Raney
Ni catalysts.27 We observed that aromatics from catalytic microwave pyrolysis were almost
completely converted into high-density cycloalkanes in the n-heptane medium under a
low-severity condition.
This present work aims to unravel the key interactions between cellulose and low-density
polyethylene (LDPE), which act as the model compounds of biomass and waste plastics in
MSW. To this end, cellulose and LDPE were first co-pyrolyzed using a tandem
microwave-induced system coupled with a downstream catalysis system. For converting
286
aromatics into renewable cycloalkanes for jet fuels, several species of liquid organics
extracted by the optimum solvent (n-heptane) was transformed into saturated hydrocarbons
by using home-made Raney Ni catalyst in the hydrogenation process. The influence of
organic species in the overall carbon yield and product distribution of saturated hydrocarbon
in the jet fuel range was also evaluated.
9.3 Experimental sections
9.3.1 Materials
Cellulose (CAS number 9004-34-6) was purchased from Sigma-Aldrich Corporation (St.
Louis, MO, USA). Cellulose is in the form of microcrystalline powders and particle sizes is
averaged at 50 µm. Low-density polyethylene (LDPE) (CAS number 9002-88-4) in the form
of pellets was purchased from Sigma-Aldrich Corporation (St. Louis, MO, USA). The density
and melting point of LDPE are 0.925g/cm3 at ambient temperature and 116 °C, respectively.
The elemental composition of cellulose and LDPE is described in Table S1. Parent ZSM-5
(SiO2/Al2O3 Mole Ratio: 50) was purchased from Zeolyst International, USA.
Nickel-Aluminum alloy powder in a non-activated type was used as purchased from Alfa
Aesar (Ward Hill, MA, USA).
9.3.2 Catalyst preparation
The activity of parent ZSM-5 was improved by suffering both hydrothermal and calcined
treatments. Under the gentle stirring, parent ZSM-5 powder was added into deionized water
(mass ratio=1) at 60 °C. After addition, the mixture was kept on stirring for 2 h under this
287
condition. The slurry was then dried at 105 °C till constant weight. The sequential process
was the catalyst calcination: hydrothermally treated ZSM-5 was calcined at 550°C for 5 h in a
muffle furnace. The catalysts were pelletized and sieved to 20 – 40 mesh. After reaction, the
spent catalysts were regenerated at 550°C for 2 h in a muffle as well.
Home-made Raney Ni catalyst was developed using a 20 wt% NaOH aqueous solution to
remove Al in the following procedure. 10 g of the above Ni-Al was slowly added into 100
mL NaOH aqueous solution under gentle stirring. After addition, the slurry was kept on
stirring at 80 °C for 1 h. The excess of sodium hydroxide was finally washed with distilled
water until nearly neutral pH was reached. The obtained Raney Ni catalyst was stored in
water. Raney Ni is notorious for its pyrophoricity, and it may ignite spontaneously when dried
in air. The Raney Ni catalyst was thus dried at 60 °C till constant weight in the atmosphere of
nitrogen to avoid contact with air, prior to the subsequent catalytic test.
9.3.3 Co-feed catalytic microwave-induced pyrolysis of cellulose and LDPE
Detailed experimental setting was described in our previous studies.35, 36 Cellulose powder
was first air dried at 105 °C for 24 h to remove the physically bound moisture, prior to
conducting the experiments. Co-feeding of cellulose with LDPE were placed in a 500 mL
quartz flask which was placed inside a Sineo MAS-II batch microwave oven (Shanghai,
China) by a constant microwave power setting (700 W). 0.05 g of activated carbon powder
was used as the absorber for the microwave-assisted pyrolysis. All reactions of microwave
pyrolysis were conducted at the temperature of 480 °C for 10 min to assure the sufficient
288
interactions between cellulose and LDPE.
9.3.4 Hydrogenation of liquid organics derived from co-feed catalytic microwave
pyrolysis
According to the solvents influence in the hydro-cycloaddition of a model compound
(naphthalene) in our previous study,27 the combined bio-oils evolved from co-feed catalytic
microwave pyrolysis were extracted by the optimal solvent (n-heptane). To produce
renewable cycloalkanes for jet fuels, a closed reaction system with a stirred stainless batch
reactor of the 4592 micro stirred reactor (with a 50 mL vessel) and a 4848 reactor controller
from Parr Instrument Company (Moline, IL, USA) was used. The mixture of organics and the
n-heptane medium was loaded into the reactor together with 20 wt% home-made Raney Ni
catalyst (in terms of the reactants). Then the reactor was sealed and vented for five times with
hydrogen to get rid of the air present in the vessel. Hydrogen was subsequently adjusted to
reach the set pressure (500 psi). The automatic controller was employed to control the
temperature and the revolution of stirrer (300 rpm). The pressure inside the reactor was
recorded and the reactions proceeded at 200 °C for 2 h. After the experiment finished, stirring
was stopped and the reactor was rapidly cooled to ambient temperature. Then, the gas was
collected for analysis and the reactor was depressurized. Consequently the liquid product was
filtered to remove catalyst particles.
9.3.5 Analytical techniques
Elemental analysis (C, H and N) of feedstock, liquid samples, char, and coke deposited on
289
spent catalysts was conducted using a 2400 Series II CHN/O Elemental Analyzer
(PerkinElmer, USA).
The textural properties of the catalyst were determined by means of N2 adsorption–desorption
(Micromeritics TriStar II 3020 Automatic Physisorption Analyzer). Fresh catalysts
(well-promoted ZSM-5 catalyst and home-made Raney Ni catalyst) were degassed in vacuum
at 300 °C for 1 h. The Brunauer–Emmett–Teller equation was applied to calculate the specific
surface area using adsorption data at p/po= 0.05–0.25. The pore volume was evaluated by
using the Barrett–Joyner–Halenda (BJH) method.
The acidity of parent ZSM-5 and well-promoted ZSM-5 catalyst was measured by
temperature-programmed desorption (TPD) of ammonia with a Micromeritics AutoChem II
2920 Chemisorption Analyzer equipped with a PFEIFFER mass spectrometer. These samples
were saturated with NH3 at the room temperature in a flow of 10% NH3 in nitrogen. After
NH3 saturation, the weakly bound NH3 was desorbed prior to the measurement at 120 °C for
1 h with a He flow rate of 25 ml/min. The desorption curve was then obtained at a heating
ramp of 10 °C/min from 120 °C to 550 °C at a He flow rate of 25 ml/min.
Powder X-ray diffraction (XRD) patterns were applied on a Rigaku Smartlab X-ray
diffractometer equipped with a Cu Kα X-ray source, which was operated at 40 kV and 40 mA.
The scattering angle 2θ was changed from 10° to 80°.
290
The particle size and surface morphology of the samples were evaluated with a scanning
electron microscope (SEM, FEI Quanta 200 F).
The chemical composition of the bio-oils was characterized and qualified by Agilent 7890A
GC-MS (GC–MS; GC, Agilent 7890A; MS, Agilent 5975C) with a DB-5 capillary column.
The GC was first programmed to heat to 45°C for 3 min followed by heating to 300°C at a
rate of 10°C/min. The injection sample size was 1 μL. The flow rate of the carrier gas
(helium) was 0.6mL/min. The ion source temperature was 230 °C for the mass selective
detector. Compounds were identified by comparing the spectral data with that in the NIST
Mass Spectral library. The area percent of changed concentrations of model compounds
obtained from GC/MS results was utilized to predict product concentration in bio-oils. All the
measurements were triplicated to assure reproducibility.
The moisture content in the bio-oils was determined by a Karl Fischer (KF) compact titrator
(V20 Compact Volumetric KF Titrator, Mettler-Toledo).
The gaseous product was collected in a 1L Tedlar gas bag and then offline analyzed by an
INFICON 3000 Micro-GC (INFICON Inc., Santa Clara, CA, USA) system with a thermal
conductivity detector (TCD). A standard gas mixture consisting of H2, N2, CH4, CO, CO2,
C2H4, C2H6, and C3H6 was used to calibrate the yield of non-condensable gas. Alkanes and
olefins (>C4) in gas samples were either not detected or negligible in this research. All the
measurements were triplicated to assure reproducibility.
291
9.3.6 Experimental methods and data evaluation
A central composite experimental design (CCD) was employed to optimize the process
conditions and product yields from co-feeding of cellulose and LDPE in catalytic microwave
pyrolysis (Table 1). The catalytic temperature (X1, °C) and LDPE to cellulose ratio (X2) were
chosen as independent variables. The loading of cellulose was 20 g for each run, while the
catalyst to co-reactants ratio was kept constantly at 0.2 in the co-feed catalytic microwave
pyrolysis. In these experiments based on CCD, the mass of LDPE pellets varied from 8 to 22
g, while packed bed catalytic temperature ranged from 269 to 481 °C. Four additional
experiments were conducted as the controls (Entry 14 – 17, Table 9.1).
The coke mass was determined by the difference before and after catalytic
co-pyrolysis. Because of ex-situ catalysis, char was left in the quartz flask inside
microwave reactor; while the coke was formed on the well-promoted ZSM-5 catalyst
in the packed-bed catalysis reactor. The weight of non-condensable gas was calculated
using the following equation.
ℎ
=
−
− ℎ
−
(1)
Overall carbon yields of the liquid, gas, and solid products and carbon selectivity of a specific
product were calculated based on the following equations.
B
C
B
C
D
=
H D=
C
C
C
E
C
× 100%
E
E
292
(2)
× 100%
(3)
Table 9.1 Experimental design and product yield distribution.a
Yield (wt%)
Catalytic
LDPE
Temperature (ºC)
cellulose ratio
Bio-oil
Gas
Char
Coke
1
300
0.5
45.88
38.10
14.45
1.57
2
300
1
50.10
39.09
9.51
1.30
3
450
0.5
38.31
45.90
14.56
1.23
4
450
1
46.34
43.57
9.34
0.75
5
375
0.75
45.22
42.51
11.00
1.27
6
375
0.75
45.16
42.84
10.78
1.22
7
375
0.75
46.02
42.13
10.59
1.26
8
375
0.75
45.35
42.34
11.12
1.19
9
375
0.75
45.49
42.41
10.87
1.23
10
375
0.4
39.12
43.75
15.54
1.59
11
375
1.1
50.45
40.03
8.63
0.89
12
269
0.75
48.75
38.71
11.04
1.50
13
481
0.75
42.18
46.23
10.95
0.64
14
375
1.45
51.45
40.71
7.03
0.81
15
250
0.75
50.27
37.13
10.87
1.73
16
500
0.75
40.97
47.58
10.88
0.57
17
375
0
33.54
48.55
16.09
1.82
Entryb
to
a
reaction condition: reaction temperature 480 ºC; reaction time, 10 mins.
b
Entry-1 to Entry-13 were conducted based on central composite design; Entry-14 and
Entry-16 were added as the controls; Entry-17 is the control in the absence of LDPE.
293
Table 9.2 Textural properties of well-promoted ZSM-5 and home-made Raney Ni catalysts.a
SBET
Vpore
Spore
dpore
(m2/g)
(cm3/g)
(m2/g)
nm
Parent ZSM-5
386.9
0.078
55.3
5.7
Well-promoted ZSM-5
396.2
0.097
74.1
5.2
Ni-Al alloy
0.35
0
0
0
Home-made Raney Ni
52.4
0.034
35.8
3.8
a
SBET: BET surface area; Vpore: pore volume; Spore: pore surface area; dpore: average pore
size
9.4 Results and discussion
9.4.1 Catalyst characterization
Comparing with the characteristics of parent ZSM-5, the BET surface area, pore volume, and
pore surface area were significantly promoted by the combined treatments. Furthermore, the
well-promoted ZSM-5 by means of the combined treatments gave rise to the generation of
secondary porosity (mesoporosity) in the ZSM-5 zeolite matrix as listed in Table 9.2. As
expected, the average pore size of the modified ZSM-5 is 5.2 nm, which is consistent with
naphthalene diameter (5.5 nm), thus double-ring aromatics can be readily adsorbed in the
pores.37 Since the overall acid amounts can be inferred from the relative peak areas of the
NH3 desorption curves; it was observed that the acidity are comparable (Fig. 9.1). For the
well-promoted ZSM-5 catalyst, the maximum of the peak shifted towards low temperatures,
which accompanied with a decrease in percentage of strong acid sites. The catalyst with the
decrease of strong acid sites could reduce catalytic cracking of large molecules into gaseous
molecules, thereby increasing the carbon yield of liquid organics. On the other hand, the
minimum of the peck catalysts was not impacted by the integrated treatments. Hence, the
294
well-promoted ZSM-5 catalyst used in the catalytic reaction was modified by the combined
treatments and then pelletized and sieved to 20 – 40 mesh.
Fig. 9.1 NH3-TPD profiles of parent ZSM-5 and well-promoted ZSM-5 catalyst.
For the production of home-made Raney Ni catalyst, Al in the Ni-Al alloy powder was
dissolved by NaOH solution during alkali leaching. As such, the textural properties of
home-made Raney Ni catalyst were also demonstrated in Table 9.2. After the alkali treatment,
BET surface area, pore volume, and pore surface area were substantially improved. The BET
surface area of the Raney Ni catalyst remarkably increased by dissolution of Al composition
from 0.35 to 52.4 m2/g. In addition, pore volume and pore surface area progressively went up
to 0.034 cm3/g and of 35.8 m2/g, respectively. It is well known that the high BET surface area
could assist the adsorption of hydrogen for the hydrogenation reaction. The average pore size
of home-made Raney Ni catalyst is 3.8 nm, which is suited for the diffusion of mono-cyclic
aromatics in the pores.37
295
Fig. 9.2 shows the XRD patterns of the Ni-Al alloy powder and home-made Raney Ni
catalyst. It can be seen that the XRD patterns of Ni-Al alloy was comprised of the two
categories of Ni3Al2 and Ni3Al domains. As the Al component dissolved by 20 wt. % NaOH
solution, the diffractions regarding metallic Ni were obtained as amorphous nature for
home-made Raney Ni catalyst. It was also found that the home-made Raney Ni catalyst
mainly presented diagnostic (111), (200), and (220) diffractions of fcc Ni at 2θ of 44.5, 51.8,
and 76.3°, respectively.38
Fig. 9.2 The XRD patterns of the Ni-Al alloy powder and home-made Raney Ni
catalyst.
The SEM images of Ni-Al alloy and home-made Raney Ni catalyst are displayed in Fig. 9.3.
The morphological difference between Ni-Al alloy and home-made Raney Ni catalyst are
readily visible from Fig. 9.3 (A) and (B). It can be seen that the home-made Raney Ni
catalyst is constituted by the typical fractured and angular particles, which is identical to
other research;39 whereas Ni-Al alloy showed the intact metallic structure. It was also found
that the small particles of home-made Raney Ni catalyst were dispersive, reaffirming the high
296
BET surface.
A
B
Fig. 9.3 SEM images of Ni-Al alloy powder (A) and home-made Raney Ni catalyst (B).
9.4.2 Product yield distributions from co-feed catalytic microwave pyrolysis
The product yield distribution from co-feeding of cellulose with LDPE in the catalytic
microwave pyrolysis are listed on the basis of catalytic temperature and LDPE to cellulose
mass ratio in Table 9.1.
The mass yields of bio-oil and non-condensable gas were strikingly
affected by the two variables in the range from 33.54 to 51.45 wt% versus 37.13 to 48.55
wt%, respectively. The optimal condition for maximizing bio-oil yield was presented to be at
a catalytic temperature of 375 ºC with the LDPE to cellulose ratio of 1.45; while the
maximum gas yield was predicted at 375 °C with the absence of LDPE. In general, the water
content in the bio-oil varied from 9.67 to 29.74 wt% depending on the reaction conditions.
Without the application of LDPE in the catalytic microwave pyrolysis, the water content
could attain ~ 30 wt%. However, the co-feeding of LDPE in the catalytic microwave
297
pyrolysis of cellulose reduced the water content down to 10 wt% in the bio-oil. As discerned,
olefins from thermal decomposition of LDPE could react with cellulose-derived furans to
form aromatic hydrocarbons through Diels-Alder reactions followed by dehydration at
Brönsted acid sites.18, 40 As a result, if the extent of dehydration reaction was enhanced in the
co-feed catalytic microwave pyrolysis; the increasing trend of water content in the bio-oil
was observed. Furthermore, the production of aromatic hydrocarbon from catalytic cracking
of olefins and subsequent cyclization was also influenced by the two independent variables.36
Solid carbonaceous residue (char and coke) from thermal degradation and catalytic reforming
can be distinguished because of the ex-situ catalysis. It can be seen that catalytic temperature
and LDPE to cellulose ratio affected mass yields of both char and coke. The char yield was in
the range from 7.03 wt% at the highest LDPE to cellulose ratio of 1.45 to 16.09 wt% without
the usage of LDPE. The char residue from the co-feed catalytic microwave pyrolysis was
lower than catalytic microwave pyrolysis of cellulose alone. These results indicate that
co-feeding of cellulose with LDPE in the catalytic microwave pyrolysis was thought to have
a pronounced impact on the char formation. It is mainly due to the fact that with the presence
of hydrogen originating from the thermal decomposition of LDPE, the recondensation
reactions of oxygenates evolved from cellulose were partially inhibited.20, 41 Therefore the
char formation was strongly reduced when cellulose was co-pyrolyzed with LDPE.
Given the coke deposited on the catalyst resulting in the loss of the active sites and
micropores blockage,30 the co-feed catalytic microwave pyrolysis is necessitated to aid in
298
coke reduction and the extension of
catalyst lifetime. The coke deposition yield with regard
to catalytic temperature and LDPE to cellulose ratio is also depicted in Table 9.1, ranging
from 0.57 to 1.82 wt%. As expected, the coke deposition was impacted by the co-feeding of
LDPE, which abated the formation of coke precursors. It was found that the highest yield of
coke was obtained with the absence of LDPE. Formation and deposition of coke was
considerably attenuated as LDPE was co-fed. It has been suggested that in the presence of
ZSM-5 catalyst, cellulose-derived furans could react with LDPE-derived olefins via
Diels-Alder reactions followed by dehydration reactions, thereby reducing the coke formation
from polymerization of furans.6, 7 Furthermore, plastics-derived hydrocarbons could provide
hydrogen
for cellulose-derived
oxygenates,
mitigating the
coke formation from
hydrogen-deficient oxygenates in the zeolite-catalyzed conversions.20 For ANOVA analysis
of the experiment based on CCD, the P-value of the four models (the bio-oil, gas, char, and
coke yield) are all less than 0.05, implying the models are significant to present the
relationships between product yields and the two independent variables. The coefficient of
determination (R2) for the four models are all larger than 0.96, evidencing that these models
fairly represent the relationships with respect to the independent variables. Consequently,
these four models can be used to predict the maximum product yields in accordance to the
two variables.
9.4.3 The effect of catalytic temperature on the co-feed catalytic microwave pyrolysis
As reported in our previous studies, catalyst temperature had a pronounced effect on product
distribution by using ZSM-5 as the catalyst.30, 35 In this respect, catalytic temperature was
299
chosen as the variable to investigate the product distribution and even evaluate the quality of
liquid organics. Representative results of carbon yields regarding product distribution from
co-feed catalytic microwave pyrolysis were at an array of catalytic temperature between 250
and 500 °C at the same plastics to biomass ratio (0.75) as shown in Fig. 9.4. The carbon yield
of char did not change alongside the variation of catalytic temperature, since the packed-bed
catalysis reactor was located downstream from the microwave pyrolysis reactor. Notably,
increasing the catalyst temperature during co-feed catalytic microwave pyrolysis decreased
the carbon yield of coke. Since direct decomposition of plastics produced a large amount of
long-chained waxes,15 these waxes were readily adsorbed on the catalytic surface at the low
catalytic temperature, generating the waxy coke that enveloped the whole catalyst. Likewise,
the polymerization of aromatics with oxygenates to form coke also took place at the low
catalytic temperature. When the catalytic temperature was elevated, catalytic cracking of
waxes absorbed on the catalyst was facilitated, contributing to lighter compounds that were
more easily accessible to the pores. Thus the carbon yield of coke decreased at the expense of
large molecules cracking as the catalytic temperature was enhanced.
300
Fig. 9.4 The overall carbon yields regarding product distribution from co-feed catalytic
microwave pyrolysis in light of catalytic temperature at the same LDPE to cellulose ratio
(0.75).
Because the thermal decomposition of plastics is an endothermic reaction, the yield of
plastics-derived olefins should go up as the catalytic temperature increases at the tested
region.10 The increasing trend from degradation of plastics could strengthen the Diels-Alder
reaction, contributing to the increase of liquid aromatics.17 Nonetheless, an excessively high
catalytic temperature resulted in reverse Diels-Alder reactions to produce light olefins instead
of liquid aromatics.6, 20 It was also manifested that increasing catalytic temperature favored
the catalytic cracking of oxygenates into non-condensable gas, preventing the formation of
liquid aromatics.4, 6 Consequently, the carbon yield of liquid organics gradually decreased
owing to the elevated catalytic temperature. By contrast, the carbon yield of gas
301
monotonically enhanced at the expense of liquid organics formation as the catalytic
temperature went up to 500 °C.
The detailed carbon yields and selectivity corresponding to catalytic temperature can be more
clearly discerned in Table 9.3. As for the liquid organics from co-feed catalytic microwave
pyrolysis at 250 °C, large amounts of aliphatic olefins and cellulose-derived oxygenates were
produced in the liquid organics. This result indicates that such low catalytic temperature
could not favor the zeolite-catalyzed reactions to form aromatics.
As the catalytic
temperature went up to 375 °C, there was a minor amount of aliphatic hydrocarbons detected
in the liquid organics. Besides, aromatic oxygenates (such as phenol) existing in the liquid
organics was much lower than those from catalytic microwave pyrolysis of biomass
individually under the same conditions. It is likely that LDPE-derived olefins participated in
the Diels-Alder reaction with cellulose-derived furans followed by dehydration reaction or
directly aromatize to achieve enhanced aromatic hydrocarbons. As listed in Table 9.3,
aromatic hydrocarbons were predominantly composed of toluene, xylenes, trimethylbenzenes,
indane, indene, naphthalene, and their derivatives. The carbon selectivity toward
ethylbenzene and xylenes ranged from 6.76 to 8.02% and 18.95 to 26.92%, respectively;
whilst the carbon selectivity toward trimethylbenzene declined from 15.21 to 13.01% as
catalytic temperature increased from 250 to 500 °C. These outcomes suggest that the
dealkylation reaction of aromatics was also promoted at the elevated catalytic temperature,
leading to
increase of simple aromatic hydrocarbons.10, 15 Comparing with the amount of
oxygenates (e.g. phenol) achieved at low catalytic temperature, only a trace amount of
302
oxygenates was obtained at the catalytic temperature of 500 °C. It could be inferred from
these results that under such condition, hydrogen transfer reactions could promote the
cracking of the side chains substituted on phenyl on the catalyst surface to form aromatic
hydrocarbons.42 To make the precursors lumped in the jet fuel range, we observed that the
aromatic hydrocarbons obtained at catalytic temperature of 375 °C were more desirable.
Table 9.3 Detailed carbon yield distribution and product carbon selectivity as a function of
catalytic temperature.a
Catalytic temperature ( ºC)
250
269
375
481
500
46.35
45.02
43.87
40.37
38.87
15.41
15.18
15.28
15.19
Overall carbon yield (C mol%)
Liquid organics
Char
15.17
Coke
2.53
2.19
1.80
0.93
0.83
Gasb
35.95
37.38
39.15
43.42
45.11
Toluene
5.32
6.23
5.23
4.89
4.56
Ethylbenzene
6.76
6.34
7.03
7.66
8.02
p-xylene/m-xylene
18.95
19.35
23.33
24.55
26.92
Trimethylbenzene
15.21
16.02
15.32
13.23
13.01
Indane
1.32
1.43
1.29
2.03
2.32
Indene
1.54
1.62
1.78
1.97
1.89
Phenol
0.95
0.75
0.65
0.43
0.19
p-cresol/m-cresol
1.98
1.74
1.02
0.47
0.18
Naphthalene
5.70
6.01
6.67
6.99
6.74
1-methylnaphthalene
5.32
5.78
6.65
7.43
7.78
2-methylnaphthalene
1.25
1.94
2.02
1.99
2.34
Liquid carbon selectivity (C mol%)
303
Anthracene
0.67
1.03
1.23
1.45
1.53
Pyrene
0.23
0.12
0.32
0.34
0.45
Methane
5.25
6.23
6.43
5.67
6.02
Carbon monoxide
28.32
24.17
14.26
16.65
18.23
Carbon dioxide
20.35
18.11
16.74
18.27
18.45
Ethylene
22.74
26.45
32.77
38.78
40.11
Ethane
10.18
12.23
13.50
11.57
9.26
Propane
13.16
12.81
16.30
9.06
7.93
Gaseous carbon selectivity (C mol%)
a Reaction
condition: Catalyst, 20 wt% with respect to feed; LDPE, 75 wt% with respect to
cellulose; Reaction temperature, 480 ºC; Reaction time, 10 min.
b
Determined by difference
Non-condensable gas was another major co-product from co-feeding of cellulose with LDPE
in the catalytic microwave pyrolysis. The composition of gaseous fraction with respect to
catalytic temperature is also depicted in Table 9.3. Because of the co-feed catalytic pyrolysis,
oxygen content was rejected from furans by reacting with olefins through Diels-Alder
reactions followed by dehydration reaction, rather than decarbonylation and decarboxylation
reactions which produce CO and CO2.5, 6, 40 Thus the carbon yields of both CO and CO2 were
very low. However, it was observed that the carbon yields of both CO and CO2 increased as
the catalytic temperature go up from 375 to 500 °C. It is reaffirmed that the elevated catalytic
temperature had a negative impact on the Diels-Alder reactions,15 rejecting oxygen content
from oxygenates by decarbonylation and decarboxylation reactions. It is worth noting that the
carbon yield of ethylene gradually increased as the catalytic temperature went up to 500 °C. It
is inferred that increasing catalytic temperature accelerated catalytic cracking of waxes
304
(long-chain hydrocarbons) into light olefins. In addition, elevated catalytic temperature also
prevented light olefins from undergoing the aromatization reaction to form aromatic
hydrocarbons.8, 10
9.4.4 The effect of LDPE to cellulose ratio on the co-feed catalytic microwave pyrolysis
It was ensured that the optimal condition to obtain the lumps of aromatics hydrocarbons for
jet fuels was set at 375 °C. Since cellulosic biomass is more abundant and cheaper than waste
plastics, it is desired to enhance the mass ratio of cellulosic biomass in the co-reactants if this
cannot decrease the carbon yield of liquids substantially. To evaluate how the effect of LDPE
proportion affected the product distribution in the co-feed catalytic microwave pyrolysis, the
overall carbon yield of product distribution at 375 °C as a function of LDPE to cellulose ratio
is shown in Fig. 9.5. The carbon yield of liquid organics increased nonlinearly with the
increase of LDPE to cellulose ratio. In comparison with the run conducted in the absence of
LDPE, the carbon yield of liquid organics rapidly increased as the ratio went up to 0.75. This
increasing trend indicates that there was a positive synergy between the co-reactants. This
trend can be easily rationalized because the degradation of LDPE produced large amounts of
olefins with high H/Ceff ratio, reacting with cellulose-derived oxygenates by the Diels-Alder
reaction to augment liquid organics yield.43,
44
It is manifested that the yields of liquid
organics from co-feed catalytic microwave pyrolysis were much higher than the optimal
result (above 30%) from catalytic microwave pyrolysis of LDPE alone.36 Nevertheless, as the
LDPE to cellulose ratio increased from 0.75 to 1.45, the carbon yield of liquid organics
slightly went up. It has been suggested that when LDPE was fed in excess relative to
305
cellulose, the conversion of excessive olefins into aromatic hydrocarbons required multiple
reaction steps (e.g. oligomerization, cyclization, and aromatization reactions ) and was
therefore less efficient.20 Consequently, the carbon yield of liquid organics was not
significantly improved as the LDPE to cellulose ratio increased from 0.75 to 1.45.
Fig. 9.5 The overall carbon yields regarding product distribution from co-feed catalytic
microwave pyrolysis as a function of LDPE to cellulose ratio at the same catalytic
temperature (375 °C).
The carbon yields of char and coke gradually decreased with the increase of LDPE to
cellulose ratio as shown in Fig. 9.5. It was found that both the carbon yields of char and coke
dramatically declined as the LDPE to cellulose ratio went up to 0.75; while they show slight
increasing tendencies as the ratio started 0.75 to 1.45. It is reaffirmed that there was a
306
synergistic effect between the co-reactants in the ratio range from 0 to 0.75; whereas
excessive LDPE could not further enhance the synergy. Hence, there exists an optimal LDPE
to cellulose ratio (0.75) for the Diels-Alder reaction in co-feed catalytic microwave pyrolysis.
The increasing carbon yield of gas with the increasing LDPE to cellulose ratio was attained.
As outlined above, the Diels-Alder reaction could not be further enhanced at the high ratios,
thus the light olefins from the catalytic cracking of LDPE assisted the carbon yield of
non-condensable gas.
Table 9.4 lists how the carbon selectivity for the aromatic species changes with the addition
of LDPE. In general, the carbon selectivity of main monocyclic aromatics first went up and
then decreased as the LDPE to cellulose ratio increased; whist the carbon selectivity of main
polycyclic aromatics gradually declined. For instance, the maximum carbon selectivity of
aromatic hydrocarbons was xylenes, which increased from 20.24 to 23.33% and then
declined to 21.47. Unlike the xylenes, the second abundant carbon selectivity toward
aromatics was trimethylbenzene, showing a decreasing tendency from 17.45 to 13.12%. On
the other hand, the carbon selectivity toward polycyclic aromatics, such as naphthalene,
experienced an appreciable drop as the LDPE to cellulose ratio increased. It is attributed to
the fact LDPE-derived olefins competed with monocyclic aromatics for reacting with
cellulose-derived oxygenates, which suppressed the polymerization reaction of monocyclic
aromatics to form polycyclic aromatics.20 From the perspective of aromatic production with
the carbon number in the jet fuel range, the run conducted at the LDPE to cellulose ratio of
0.75 was preferred in the co-feed catalytic microwave pyrolysis. Adding more LDPE would
307
not considerably improve the carbon yield of liquid yield; yet, it would be a waste of LDPE.
Table 9.4 Detailed carbon yield distribution and product carbon selectivity on the basis of
LDPE to cellulose ratio. a
LDPE to cellulose ratio
0
0.4
0.75
1.1
1.45
Liquid organics
24.53
36.05
43.87
44.23
44.88
Char
31.78
24.04
15.18
11.31
8.83
Coke
3.76
2.57
1.80
1.22
1.06
Gasb
39.93
37.34
39.15
43.24
45.23
Toluene
4.34
4.67
5.23
5.76
6.61
Ethylbenzene
5.54
6.32
7.03
8.05
9.27
p-xylene/m-xylene
20.24
21.54
23.33
21.58
21.47
Trimethylbenzene
17.45
16.34
15.32
14.36
13.12
Indane
1.05
1.33
1.29
1.37
1.01
Indene
2.43
2.15
1.78
1.24
0.97
Phenol
1.25
0.97
0.65
0.70
0.43
p-cresol/m-cresol
2.01
1.45
1.02
0.32
0.11
Naphthalene
9.05
7.97
6.67
5.23
4.17
1-methylnaphthalene
8.43
7.69
6.65
4.54
3.65
2-methylnaphthalene
4.78
3.69
2.02
1.54
1.01
Anthracene
2.05
1.07
1.23
1.04
0.44
Pyrene
1.24
0.89
0.32
0.29
0.17
Methane
8.97
7.23
6.43
5.67
5.06
Carbon monoxide
40.78
26.98
14.26
12.19
10.21
Carbon dioxide
30.87
24.33
16.74
13.34
10.76
Overall carbon yield (C mol%)
Liquid carbon selectivity (C mol%)
Gaseous carbon selectivity (C mol%)
308
Ethylene
6.67
20.67
32.77
40.48
48.47
Ethane
8.50
11.76
13.50
12.98
12.21
Propane
4.21
9.03
16.30
15.34
13.29
a Reaction
condition: Catalyst, 20 wt% with respect to feed; Reaction temperature, 480 ºC;
Catalytic temperature, 375 ºC; Reaction time, 10 min.
b
Determined by difference
The gaseous composition with respect to LDPE to cellulose ratio is also demonstrated in
Table 9.4. The carbon yield of methane gradually decreased with the increase of LDPE to
cellulose ratio; whereas, the carbon yield of both CO and CO2 decreased. It is mainly
attributed to the Diels-Alder reaction that was facilitated between cellulose and LDPE,
instead of the decarbonylation and decarboxylation reactions. Unlike the CO and CO2, the
carbon yield of ethylene was found to increase in the most pronounced way from 6.67 to
48.47%, which in turn made the carbon yields of CO and CO2 decreased. These outcomes
verifies that the increasing yield of ethylene could originate from catalytic cracking of waxes
that was from the thermal degradation of LDEP. Besides, cellulose-derived oxygenates could
promote the degradation of large molecules to light olefins in the co-feed catalytic
pyrolysis.10
9.4.5 Hydrogenation of liquid organics derived from co-feed catalytic microwave
pyrolysis for jet fuels
Considering the negative effect of water in the hydrogenation process,28 the small amount of
water in the raw bio-oil from the co-feed catalytic microwave pyrolysis should be removed.
Since n-heptane played an important role in the hydrogenation process,27 the raw bio-oil was
309
thus separated and extracted by the optimum solvent (n-heptane). The liquid organics and
n-heptane in every samples were separated and weighed to measure the loss of organics in the
extracting step. As the liquid organics produced by the feeding of cellulose with LDPE in the
catalytic microwave pyrolysis (at the catalytic temperature of 375 °C with the LDPE to
cellulose ratio of 0.75) principally consisted of C8 - C16 aromatics, the controllable adjustment
of aromatic hydrocarbons with 8 - 16 carbon numbers are considered as precursors that
should be converted into jet fuels. In addition, the run that obtained the maximum carbon
yield of liquid organics (at the catalytic temperature of 375 °C with the LDPE to cellulose
ratio of 1.45) and the run that obtained the liquid organics within the lowest amount of
oxygenates (at the catalytic temperature of 500 °C with the LDPE to cellulose ratio of 0.75)
were hydrogenated for the optimization of jet fuels. Liquid organics acting as the control
from the experiment that cellulose was subjected to catalytic microwave pyrolysis alone was
hydrogenated as well. As reported previously,
naphthalene was completely transformed into
saturated decalin in the n-heptane medium;27 thus the mass ratio of reactant to solvent was set
at 1:7 for the following hydrogenation process.
The product distribution and the carbon selectivity toward main alkanes from the
hydrogenation of diverse raw organics by using the home-made Raney Ni as the catalyst are
elucidated in Table 9.5. Unreacted hydrogen (over 99 vol%) was detected at the end of all
reactions, suggesting that the reactions were not conducted under hydrogen starved
conditions. The overall carbon yields of these hydrogenated organics (L-1, L-2, L-3, and L-4)
with respect to the co-reactants of cellulose and LDPE were 21.91, 39.18, 34.71 and 40.08%,
310
respectively. These overall carbon yields of hydrogenated organics for jet fuels were much
higher than that in our previous research.27 Among the hydrogenated organics, the result for
the production of cycloalkanes from hydrogenated L-3 was superior to the others under the
same condition. It was observed that more than 90% selectivity towards cycloalkanes was
achieved from hydrogenated L-3. The high amounts of high-density cycloalkanes could be
potentially used as high-density jet fuel (e.g., JP-10 and RJ-5). Yet it was found that the
overall carbon yield of the hydrogenated organics was a little lower than those of both
hydrogenated L-2 and L-4. For the production of aliphatic alkanes, hydrogenated L-4
obtained the maximum selectivity (31.45%). From this perspective, the contents of
hydrogenated L-4 were consistent with that in JP-5, containing 31.23% of aliphatic alkanes,
53.06% of cycloalkanes, and 15% of remaining aromatic hydrocarbons.45 The total amounts
of hydro-aromatic hydrocarbons and aromatic hydrocarbons were less than 15%, which meet
the specifications of commercial jet fuels. A minor mount of other compounds were found in
the all hydrogenated organics except hydrogenated L-1. As mentioned above, these were
some oxygenates in the L-1 from catalytic microwave pyrolysis of cellulose alone. These
oxygen content could not be removed by hydrogenation reaction under the low-severity
condition; for example, phenol in L-1 was just convert into cyclohexanol via
hydro-cycloaddition reaction. Less than 1 vol% of small hydrocarbons (such as methane,
ethane, and propane) were obtained, which suggests that there was almost no carbon loss by
hydrocracking reaction in the hydrogenation system.
It was found that the maximum carbon selectivity toward main monocyclic alkanes were all
311
from hydrogenated L-3, except trimethylcyclohexanes. Trimethylcyclohexanes derived from
the hydro-cycloaddition of trimethylbenzenes had the highest carbon selectivity from
hydrogenated L-1. That is possibly because the dealkylation reaction was favored at the high
catalytic temperature (500 °C). It is noticeable that the maximum carbon selectivity of main
aliphatic alkanes was all achieved from hydrogenated L-4, which is reaffirmed that the high
LDPE loading contributed to the high carbon yield of aliphatic hydrocarbons. Of the four raw
organics for hydrogenation, the hydrogenated organics derived from L-2 produced the highest
amounts of saturated alkanes matching jet fuels. With the consideration of both overall
carbon yield and product distribution, the raw organics (L-2) is thus the optimal source to
manufacture renewable jet fuels.
Table 9.5 Products distribution
and partial alkanes’ carbon selectivity for hydrogenation of
diverse liquid organicsa
Organics speciesb
L-1
L-2
L-3
L-4
Aliphatic alkanes
-
10.71
1.34
31.45
Cycloalkanes
83.69
78.15
90.87
57.79
Hydro-aromatic hydrocarbons
8.42
5.28
5.01
4.01
Aromatic hydrocarbons
2.74
4.63
2.03
5.89
Others
5.15
1.23
0.75
0.86
1,4-dimethylcyclohexane
4.35
4.25
5.37
3.68
1,3-dimethylcyclohexane
11.54
12.78
16.76
12.11
1,2-dimethylcyclohexane
3.56
6.09
8.25
6.54
Ethylcyclohexane
2.67
3.13
3.67
3.25
Overall selectivity (% in area)
Alkanes selectivity (C mol%)
312
Octane
-
0.67
0.15
1.23
1,2,4-trimethylcyclohexane
18.54
16.89
14.23
11.98
Propylcyclohexane
4.45
3.21
3.64
2.13
Nonane
-
0.76
0.09
1.92
Hexahydroindan
4.56
3.32
4.78
2.21
Decalin
13.65
6.23
8.23
4.15
Decane
-
0.64
-
1.32
Undecane
-
2.33
0.21
4.41
Dodecane
-
0.76
-
1.47
Tridecane
-
0.28
0.12
1.89
Tetradecane
-
0.65
-
2.08
Pentadecane
-
0.41
-
2.15
Hexadecane
-
0.38
0.11
1.76
a
Reaction condition: Initial pressure, 500 psi; Raney Ni catalyst, 20 wt% with respect to
reactant mass;
b
Reaction temperature, 200 ºC; Reaction time, 2 h.
L-1: reactant from the experiment conducted at catalytic temperature of 375 ºC in the
absence of LDPE; L-2: reactant from the experiment conducted at catalytic temperature of
375 ºC with LDPE to cellulose ratio of 0.75; L-3: reactant from the experiment conducted at
catalytic temperature of 500 ºC with LDPE to cellulose ratio of 0.75; L-4: reactant from the
experiment conducted at catalytic temperature of 375 with LDPE to cellulose ratio of 1.45.
313
H
H
H
H
H
H
Fig. 9.6 Proposed reaction pathways for the conversion of cellulose and LDPE into jet fuel
range alkanes.
9.4.6 Reaction pathway for the conversion regrading co-feeding of cellulose with LDPE
into jet fuels
These observations are the key point to propose the reaction pathway for the conversion
regarding co-feeding of cellulose with LDPE into jet fuels. The results of the hydrogenated
314
organics for jet fuels were evidenced that the reaction pathway was more complicated than
those mentioned previously.7, 9, 20 Based on the quantified products distribution in this study,
related results from catalytic microwave-induce degradation of LDPE,36 and lignocellulosic
biomass to jet fuel through combined catalytic conversions;30, 46 the overall reactions network
(including co-feed catalytic microwave pyrolysis and hydrogenation process) is proposed in
Fig. 9.6. In the primary route regrading cellulose, cellulose underwent a series of dehydration,
decarboxylation, and decarbonylation to form furan compounds (e.g. furfural) during thermal
degradation.12, 30 For the decomposition of LDPE by another route, thermal degradation of
LDPE took place through two mechanisms: random scission and chain-end scission.47, 48 The
two aforementioned mechanisms occurred simultaneously, giving rise to free radicals
together with the long carbon chains.49 Meanwhile, the radical fragments could also be
transformed into straight chain dienes, alkenes and alkanes through hydrogen transfer
reactions.49 The hydrogen from the thermal degradation of LDPE could be provided for
cellulose-derived oxygenates that acted as the strong acceptor of hydrogen, suppressing the
char formation.
The long-chain hydrocarbons from the thermal degradation of LDPE could subsequently go
through catalytic cracking in the presence of well-promoted ZSM-5 catalyst by two
carbocationic mechanisms to form light olefins.47, 50 These olefins could react with furans
through the Diels–Alder reaction followed by the dehydration reaction to generate aromatic
hydrocarbons. Meanwhile, these LDPE-derived olefins could individually subject to
oligomerization, cyclization and aromatization reactions to form aromatic hydrocarbons.15
315
Likewise, these furans derived from cellulose could go through decarbonylation,
aromatization, and oligomerization reactions inside the pores of well-promoted ZSM-5
catalyst to form aromatic hydrocarbons alone. As reported previously, the interactions
through the hydrocarbon pool mechanism could exist in addition to Diels–Alder reaction.7
Since the carbon yield of liquid organics significantly increased and the coke yield
dramatically decreased as the LDPE was introduced, the Diels–Alder reaction was thus the
dominant reaction pathway during the co-feed catalytic microwave pyrolysis in comparison
to the hydrocarbon pool mechanism.
Furthermore, it was observed that the liquid organics resulting from the co-feed catalytic
microwave pyrolysis mostly produced by a mixture of aromatic hydrocarbons, aliphatic
olefins, and alkanes. In the hydrogenation process by using the home-made Raney Ni as the
catalyst, the hydrogenation of the aliphatic olefins was the one of the first step to take place.51
The
aromatic
hydrocarbons
were
thereafter
hydrogenated
into
cycloalkanes
or
hydro-aromatic hydrocarbons via hydro-cycloaddition reactions under the very mild reaction
condition. In addition, the hydroisomerization reaction could took place between the
dimethylcyclohexanes; and trace volume of small hydrocarbons was also produced by
hydrocracking reactions. From the detailed analysis of hydrogenated organics, it was found
that these hydrogenated organics can be potentially used as additives to make the desired jet
fuels by blending with other hydrocarbons.
316
9.5 Conclusions
In summary, the co-feed microwave-induced pyrolysis followed by hydrogenation process is
a profound approach for the improved carbon yield of renewable alkanes for jet fuels from
cellulose and LDPE. There was a positive synergy between cellulose and LDPE in the
co-feed catalytic microwave pyrolysis, which could significantly improve the carbon yield of
raw organics and suppress the formations of char and coke. The excellent catalytic properties
of the well-promoted ZSM-5 catalyst were strongly related to the mesopores formed through
integrated treatments. The raw organics with improved carbon yield (43.87%), which was
derived from the run conducted at the catalytic temperature of 375 °C with the LDPE to
cellulose ratio of 0.75, were lumped in the jet fuel range.
The home-made Raney Ni catalyst manufactured showed good textural properties such as
high BET surface area. In the hydrogenation process, four species of raw organics were
hydrogenated under a low-severity condition. Given the overall carbon yield and product
distribution, the raw organics conducted at the catalytic temperature of 375 °C with the LDPE
to cellulose ratio of 0.75 in the co-feed catalytic microwave pyrolysis, is more suitable for the
production of renewable alkanes for advanced jet fuels (especially for high-density jet fuels).
The overall carbon yield regarding cellulose and LDPE was 39.18%. Meanwhile, ~ 90%
selectivity toward jet fuel range alkanes was obtained in such hydrogenated organics, and up
to 75% selectivity belong to high-density cycloalkanes.
It is more likely that in the future
these renewable alkanes would be potentially used as additives in jet fuels.
317
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321
CHAPTER TEN
ENHANCEMENT OF JET FUEL RANGE ALKANES FROM
CO-FEEDING OF LIGNOCELLULOSIC BIOMASS WITH
PLASTICS VIA TANDEM CATALYTIC CONVERSIONS
10.1 Abstract
Enhanced carbon yields of jet fuel range alkanes were manufactured from co-feeding of
lignocellulosic biomass with plastics. The consecutive processes proceeded via the co-feed
catalytic microwave-induced pyrolysis and hydrogenation process. In the co-feed catalytic
microwave pyrolysis by using ZSM-5 as the catalyst, parent ZSM-5 fabricated by
hydrothermal and calcined treatments contributed to the increase of surface area as well as
the formation of more mesopores. Liquid organics with enhanced carbon yield (40.54%) were
more principally lumped in the jet fuel range from the co-feed catalytic microwave pyrolysis
performed at the catalytic temperature of 375 °C with the plastics to biomass ratio of 0.75. To
manufacture home-made Raney Ni catalyst, the BET surface area, pore surface area, and pore
volume of the home-made Raney Ni catalyst were considerably improved when the Ni-Al
alloy was dissolved by the NaOH solution. In the hydrogenation process, we observed the
three species of raw organic derived from the co-feed catalytic microwave pyrolysis were
almost completely converted into saturated hydrocarbons under a low-severity condition. The
improved carbon yield (38.51%) of hydrogenated organics regarding co-reactants of biomass
and plastics predominantly match jet fuels. In the hydrogenated organics, over 90%
322
selectivity toward jet fuel range alkanes was attained. In this respect, these hydrogenated
organics with high amounts of renewable cycloalkanes can be potentially served as
high-density jet fuels or additives for blending with civilian jet fuels.
Keywords: Bio-jet fuels; well-promoted ZSM-5; home-made Raney Ni; co-feed catalytic
microwave pyrolysis; hydrogenation; cycloalkanes
10.2 Introduction
The limitation of fossil resources accompanied with huge environmental issues accelerate the
exploitation of renewable resources to substitute petroleum hydrocarbon fuels, with a
particular concentration on the development of new generation biofuels.1-3 Lignocellulosic
biomass has attracted essential attention as a carbon-neutral resource for promising potentials
to produce transportation biofuels and versatile chemicals.4, 5 Catalytic fast pyrolysis (CFP) of
lignocellulose to valuable biofuels and chemicals is recognized as the most prevailing and
promising route in a single process
6-11.
Dozens of lignocellulose and biomass-derived
feedstock have been subjected to catalytic pyrolysis for the production of advanced biofuels
with more favorable properties.6-8,
12, 13
Zeolite-based catalysts (e.g., ZSM-5) have been
recognized as the highest-efficiency catalysts to produce considerable petrochemicals
(aromatics and olefins).9, 14, 15
However, even in the presence of highly efficient catalyst, catalytic pyrolysis of
lignocellulose can solely manufacture low carbon yield (10 – 30%) of aromatic hydrocarbons;
323
large amounts of solid residues, including both biochar and coke (carbon yields usually above
30%), are achieved in the process.10,
16-19
As aromatics are considered as the high
energy-density hydrocarbons production 11, such low carbon yield of aromatic hydrocarbons
from catalytic pyrolysis is not cost-effective to scale up the process in a biorefinery. Further,
the process is commonly plagued by the high production of coke deposited on the catalyst
because the coke could rapidly deactivate the catalyst and reduce its lifetime, resulting in the
catalytic process to be impractical.20 Therefore, these are the huge challenges in the face of
commercializing the catalytic pyrolysis for the production of renewable petrochemicals.
It is discerned that the petrochemicals with the low carbon yield of aromatics and high coke
formation are mainly associated with the oxygen-enriched intrinsic nature and hydrogen
deficiency of lignocellulose.16, 18, 21 In addition, the hydrogen to carbon effective (H/Ceff) ratio
plays a significant role in coke formation and converting efficiency of biomass into advanced
biofuels.8, 22, 23 Thus the hydrogen-deficient (H/Ceff usually less than 0.3) biomass produces
low carbon yields of petrochemicals and large formation of coke when the lignocellulose
were transformed over zeolite catalysts.9, 15 In order to improve the carbon efficiency of
aromatics and minimize the coke formation, it is reasonable that the incorporation of high
H/Ceff ratio co-reactants with biomass in the catalytic pyrolysis could help mitigate these
issues. It is observed that co-feeding of lignocellulose with hydrogen-rich feedstock in the
catalytic pyrolysis can modify the reaction mechanism of oxygen removal by substituting
decarbonylation and decarboxylation with dehydration.10, 22, 24, 25 Waste plastics represent a
cheaper and abundant hydrogen sources, which can be used to improve carbon efficiency of
324
aromatics and lower the coke formation in the catalytic co-pyrolysis.9, 15, 18, 21, 26 Tremendous
quantities of waste plastics are generated each year worldwide. Among the waste plastics,
polyethylene formed by the polymerization of olefins with a H/Ceff ratio of 2 accounts for up
to 40% of gross waste plastics.27 Hence co-feeding of lignocellulose with waste plastics in
catalytic pyrolysis is remarkably beneficial for the environment and energy recapture.
Although the co-feeding of biomass with plastics in catalytic pyrolysis has a significant
synergy for aromatic production and coke reduction, these aromatics are commonly
mono-ring aromatic hydrocarbons with low carbon numbers (C6 – C8).21, 23, 25, 26 Yet, these
aromatics with low carbon numbers cannot satisfy the requirement of jet fuels 19; that is due
to the fact that the current jet fuels are principally comprised of linear-chain and
branched-chain alkanes in the range of C8 - C16.
28, 29
It is well known that cycloalkanes in
jet fuels can be burned cleanly with high heats of combustion; that is because cycloalkanes
are compact molecules within robust ring strain.30, 31 To promote the qualities of current jet
fuels, cycloalkanes with the carbon numbers in the jet fuel range can be synthesized and
added into civilian jet fuels.29, 32 The representative techniques on bio-jet fuel production
include catalytic hydrodeoxygenation of vegetable oils33, 34 and Fischer-Tropsch synthesis of
biomass-derived syngas.5, 11 Nonetheless, aromatics and cycloalkanes cannot be generated by
the two technologies. On the other hand, the promising pathway for the production of
cycloalkanes are the hydro-cycloaddition of aromatic hydrocarbons.35-38 To pursue the
precursors of renewable cycloalkanes for jet fuels, highly desirable aromatic hydrocarbon in
the jet fuel range (C8 – C16) should be generated.
325
Due to conventional ZSM-5 zeolite with microporous structure, it contributes to mono-ring
aromatics with low carbon numbers. In our previous study, moderate treatments were used to
fabricate ZSM-5 properties including its porosity and acidity, which is vital for improving
product selectivity toward aromatic hydrocarbons.39,
40
It was shown that the mild
hydrothermal and calcined modifications favored the generation of mesopores and more
catalytic sites. Accordingly our well-modified catalyst is suitable for the production of
aromatic hydrocarbons in the C8 – C16 range. Besides we have developed a combined process
to manufacture renewable cycloalkane from lignocellulosic biomass, albeit obtaining a low
carbon yield.35 In addition, it is manifested that microwave-assisted pyrolysis technology is
one of the most promising techniques to enhance the degradation reactions because of rapid
heat rate by microwave irradiation
41.
In comparison with conventional pyrolysis,
microwave-assisted pyrolysis encloses the advantages of rapid heating, easy control, and low
energy input.42,
43
In the hydro-cycloaddition process, the catalytic performance of
home-made Raney Ni catalyst was superior to purchased Raney Ni catalysts.35 During the
process, the aromatics in the n-heptane medium were almost completely converted into
saturated hydrocarbons under a low-severity condition.
Initially, the production of C8 – C16 aromatics was the target compounds from the co-feeding
of lignocellulosic biomass and plastics in the catalytic pyrolysis. To best of our knowledge,
the co-feed catalytic microwave-induced pyrolysis has not yet been explored. The main
objective of this research presented in this paper was the production of jet fuel range alkanes
326
with enhanced carbon yield. In this respect, there is no literature reported on the production
of jet fuel range alkanes from the co-feeding of lignocellulosic biomass with plastics.
Furthermore, the reaction mechanisms regarding the co-feeding of lignocellulosic biomass
with plastics through catalytic combined conversions have never been established. Therefore,
this study can fill this research gap in the production of jet fuels from co-feeding of
lignocellulosic biomass and plastics via tandem catalytic processes.
The innovations aspects of this study were the novel pathways from co-feeding of
lignocellulosic biomass with plastics by using modified catalysts, namely well-promoted
ZSM-5 catalyst and home-made Raney Ni catalyst with the purpose of enhancing the carbon
yield of aromatics and lowering the coke formation and improving the carbon selectivity of
jet fuel range alkanes. Herein, biomass and plastics were co-pyrolyzed using a tandem
microwave-induced pyrolysis system and a catalytic system, aiming to determine whether
there is an apparent synergy between co-feeding of biomass and plastics in the catalytic
microwave-induced pyrolysis. For maximizing the carbon efficiency of jet fuel range alkanes,
the various species of bio-oils derived from co-feed catalytic microwave-assisted pyrolysis
was first extracted by the optimum solvent (n-heptane). The organic mixture was
subsequently hydrogenated into saturated hydrocarbons by using home-made Raney Ni
catalyst. The influence of organic species on the carbon yield and product distribution of
saturated hydrocarbon in the jet fuel range was also evaluated.
327
10.3 Experimental sections
10.3.1 Materials
Douglas fir sawdust pellets (7 mm in diameter and 15 mm in length) were leveraged as
lignocellulosic biomass, which were supplied by Bear Mountain Forest Products Inc., USA.
Low-density polyethylene pellets (LDPE) was purchased from Sigma-Aldrich Corporation
(St. Louis, MO, USA). Its density and melting point are 0.925g/cm3 and 116 °C, respectively.
Parent ZSM-5 (SiO2/Al2O3 Mole Ratio: 50) was supplied by Zeolyst International, USA.
Nickel-Aluminum alloy powder in a non-activated form was purchased from Alfa Aesar
(Ward Hill, MA, USA).
10.3.2 Catalyst preparation
The parent ZSM-5 was fabricated by hydrothermal and calcined treatments. Parent ZSM-5
powder was added into deionized water (mass ratio=1) at 60 °C under the gentle stirring. The
slurry was kept at this condition for 2 h. The slurry was subsequently dried at 105 °C until the
weight was constant. The sequential process was calcination: parent ZSM-5 tailored by
hydrothermal treatment was calcined at 550°C for 5 h in a muffle furnace. The obtained solid
was pelletized and sieved to 20 – 40 mesh. For the regeneration of promoted ZSM-5, the
spent catalysts were burned at 550°C for 2 h in a muffle as well.
Home-made Raney Ni catalyst was fabricated by using a 20 wt% NaOH aqueous solution to
dissolve Al compostion in the Ni-Al alloy powder. In this case, 10 g of Ni-Al alloy powder
was slowly added into 20 wt% NaOH aqueous solution (100 mL) under gentle stirring. After
328
addition, the slurry was held on stirring at 80 °C for 1 h. The excess of NaOH solution was
consequently washed by distilled water until neutral pH was reached. Since Raney Ni is
notorious for its pyrophoricity, it can ignite spontaneously when it is exposed in air. The
obtained home-made Raney Ni catalyst was therefore stored in water. Prior to the subsequent
catalytic test, the Raney Ni catalyst was dried at 60 °C in the atmosphere of nitrogen to avoid
contact with air.
10.3.3 Co-feed catalytic microwave-induced pyrolysis of lignocellulosic biomass and
plastics
Detailed experimental description was shown in our previous research.44,
45
Prior to the
experiments, Douglas fir sawdust pellets were air dried at 105 °C for 24 h to remove the
physically bound moisture. The co-feeding of Douglas fir sawdust pellets with LDPE pellets
were placed in a 500 mL quartz flask, which was put inside a Sineo MAS-II batch microwave
oven (Shanghai, China) by a constant microwave power setting (700 W). 0.05 g of activated
carbon powder was leveraged as microwave absorbers for the microwave-assisted pyrolysis.
All runs for microwave-assisted pyrolysis in the microwave oven chamber were implemented
at 480 °C for 10 min.
10.3.4 Hydrogenation of liquid organics derived from co-feed catalytic microwave
pyrolysis
Due to the solvents influence in the hydro-cycloaddition of a model compound (naphthalene)
in our previous study,35 the mixed bio-oils from co-feed catalytic microwave pyrolysis were
329
extracted by the optimal solvent (n-heptane). For the hydrogenation process of extracted
organics, a closed reaction system with a stirred stainless batch reactor of the 4592 micro
stirred reactor (with a 50 mL vessel) and a 4848 reactor controller from Parr Instrument
Company (Moline, IL, USA) was employed. Herein, the mixture of organics and the
n-heptane medium was placed into the vessel with 20 wt% home-made Raney Ni catalyst
(with respect to the reactants). Then the reactor was sealed and vented for three times by
hydrogen to expel the air in the vessel. Hydrogen was subsequently adjusted to the set
pressure (500 psi). The automatic controller was used to control the temperature and the
revolution of stirrer (350 rpm). The pressure inside the reactor was recorded and the reactions
were conducted at 200 °C for 2 h. After the experiment was quenched, stirring was stopped
and the vessel was rapidly cooled to ambient temperature. Then, the gas was collected by a
1L Tedlar gas bag and the reactor was depressurized. Consequently the liquid products were
filtered to remove catalyst particles.
10.3.5 Analytical techniques
Elemental analysis (C, H and N) of reactants, liquid samples, char, and coke deposited on
spent catalysts was measured by using a 2400 Series II CHN/O Elemental Analyzer
(PerkinElmer, USA).
The textural properties of the catalysts were evaluated in light of N2 adsorption–desorption
(Micromeritics TriStar II 3020 Automatic Physisorption Analyzer). Fresh catalysts
(well-promoted ZSM-5 catalyst and home-made Raney Ni catalyst) were degassed in vacuum
330
at 300 °C for 1 h. The Brunauer–Emmett–Teller equation was utilized to figure out the
surface area by using adsorption data at p/po= 0.05 – 0.25. The pore volume was measured
through the Barrett–Joyner–Halenda (BJH) method.
The acidity was determined by temperature-programmed desorption (TPD) of ammonia with
a Micromeritics AutoChem II 2920 Chemisorption Analyzer equipped with a PFEIFFER
mass spectrometer. The samples were initially saturated with NH3 at the room temperature in
a flow of 10% NH3 in nitrogen. After NH3 saturation, the weakly bound NH3 was desorbed
prior to the measurement at 120 °C for 1 h at a He flow rate of 25 mL/min. The desorption
curve was finally calculated at a heating ramp of 10 °C/min from 120 °C to 550 °C with a He
flow rate of 25 ml/min.
Powder X-ray diffraction (XRD) patterns were applied on a Rigaku Smartlab X-ray
diffractometer equipped with a Cu Kα X-ray source, which was operated at 40 kV and 40 mA.
The scattering angle 2θ was changed from 10° to 80°.
The particle size and surface morphology of the samples were evaluated with a scanning
electron microscope (SEM, FEI Quanta 200 F).
The chemical composition of all liquid samples was characterized and qualified by Agilent
7890A GC-MS (GC–MS; GC, Agilent 7890A; MS, Agilent 5975C) with a DB-5 capillary
column. The GC was first heated to 45°C for 3 min, followed by heating to 300°C at a rate of
331
10°C/min. The injection of the sample volume was 1 μL. The flow rate of the carrier gas
(helium) was 0.6mL/min. The ion source temperature was 230 °C for the mass selective
detector. Compounds were identified by comparing the spectral data with that in the NIST
Mass Spectral library. All the measurements were triplicated to ensure reproducibility.
The moisture content in the bio-oils was measured by a Karl Fischer (KF) compact titrator
(V20 Compact Volumetric KF Titrator, Mettler-Toledo).
The gas was collected in a 1L Tedlar gas bag and then offline analyzed by an INFICON 3000
Micro-GC (INFICON Inc., Santa Clara, CA, USA) system with a thermal conductivity
detector (TCD). A standard gas mixture including H2, N2, CH4, CO, CO2, C2H4, C2H6, and
C3H6 was quantify to calibrate the yield of non-condensable gas. The gas composition (>C4)
were either not detected or negligible in this research. All the measurements were triplicated
to assure reproducibility.
10.3.6 Experimental methods and data evaluation
A central composite experimental design (CCD) was employed to optimize the process
conditions and product yields from co-feeding of biomass with plastics in catalytic
microwave pyrolysis (Table 10.1). The catalytic temperature (X1, °C) and plastics to biomass
ratio (X2) were used as independent variables. The loading of Douglas fir sawdust pellets was
20 g for each run, while the catalyst to reactants ratio was constantly held at 0.2 in co-feed
catalytic microwave pyrolysis. In these experiments based on CCD, the mass of LDPE pellets
332
varied from 8 to 22 g, while the catalytic temperature ranged from 269 to 481 °C. Four
additional experiments were conducted as the controls (Entry 14 – 17, Table 10.1).
Table 10.1 Experimental design and product yield distribution.a
Yield (wt%)
Catalytic
Plastics to biomass
Temperature (ºC)
ratio
Bio-oil
Gas
Char
Coke
1
300
0.5
39.97
39.27
17.21
3.55
2
300
1
45.32
39.93
11.86
2.89
3
450
0.5
34.55
45.73
16.74
2.98
4
450
1
40.24
46.07
12.02
1.67
5
375
0.75
42.54
41.33
13.11
3.02
6
375
0.75
41.38
42.44
13.20
2.98
7
375
0.75
42.09
41.93
12.87
3.11
8
375
0.75
41.68
42.00
13.44
2.88
9
375
0.75
41.77
42.24
13.02
2.97
10
375
0.4
35.22
43.61
17.57
3.60
11
375
1.1
43.85
41.72
11.56
2.87
12
269
0.75
42.85
40.57
13.24
3.34
13
481
0.75
38.23
47.56
12.78
1.43
14
375
1.45
45.57
42.76
9.09
2.58
15
200
0.75
47.00
36.18
13.28
3.54
16
500
0.75
36.47
49.56
12.95
1.02
17
375
0
30.57
43.89
21.69
3.85
Entryb
a
reaction condition: reaction temperature 480 ºC; reaction time, 10 mins.
b
Entry-1 to Entry-13 were conducted based on central composite design; Entry-14 and
Entry-16 were added as the controls; Entry-17 is the control in the absence of plastics.
333
The coke mass was calculated by the difference before and after catalytic process.
Because of ex-situ catalysis, char was left in the quartz flask inside microwave over
chamber; while the coke was formed on the ZSM-5 catalyst in the packed-bed
catalysis reactor. The weight of gas product was determined using the following
equation.
ℎ
=
−
− ℎ
−
(1)
Overall carbon yields of the liquid, gas, and solid products and carbon selectivity of a specific
product were calculated based on the following equations.
B
C
B
C
D
=
C
H D=
E
C
C
C
× 100%
E
(2)
× 100%
E
(3)
Table 10.2 Textural properties of well-promoted ZSM-5 and home-made Raney Ni catalysts.a
SBET
Vpore
Spore
dpore
(m2/g)
(cm3/g)
(m2/g)
nm
Parent ZSM-5
386.9
0.078
55.3
5.7
Well-promoted ZSM-5
396.2
0.097
74.1
5.2
Ni-Al alloy
0.35
0
0
0
Home-made Raney Ni
52.4
0.034
35.8
3.8
a
SBET: BET surface area; Vpore: pore volume; Spore: pore surface area; dpore: average pore
size
10.4 Results and discussion
10.4.1 Catalyst characterization
After the integral treatments, the BET surface area, pore surface area, and pore volume were
334
considerably improved. Besides, it was found that parent ZSM-5 fabricated by means of the
integrated treatments leaded to the generation of mesopores in the ZSM-5 zeolite matrix as
listed in Table 10.2. The average pore size of the promoted ZSM-5 is 5.2 nm, which is
idenical to naphthalene diameter (5.5 nm), therefore double-ring aromatics are readily
adsorbed in the mesopores.46 The overall acid amounts can be determined by the relative peak
areas of the NH3 desorption curves; it can be seen that the acidity are comparable (Fig. 10.1).
For the promoted ZSM-5 catalyst, the maximum peak shifted to low temperatures,
accompanying with a decrease of strong acid sites. The decrease of strong acid sites could
reduce the extent of catalytic cracking of large molecules into gaseous molecules, thereby
increasing the carbon yield of liquid organics. However, the minimum peck was not impacted
by the integrated treatments.
Fig. 10.1 NH3-TPD profiles of parent ZSM-5 and well-promoted ZSM-5 catalyst.
For the manufacture of home-made Raney Ni catalyst, Al in the Ni-Al alloy powder reacted
335
with NaOH solution during alkali leaching. The textural properties of Ni-Al alloy and
home-made Raney Ni catalyst were also demonstrated in Table 10.2. According to the alkali
treatment, BET surface area, pore surface area, and pore volume were substantially enhanced.
The BET surface area of home-made Raney Ni catalyst dramatically increased from 0.35 to
52.4 m2/g. Moreover, pore surface area and pore volume progressively went up to 35.8 m2/g
and 0.034 cm3/g, respectively. Accordingly, the decent BET surface area of the Raney Ni
catalyst contributed to the adsorption of hydrogen for the hydrogenation reaction. The
average pore size of the Raney Ni catalyst is 3.8 nm, which is suitable for the diffusion of
mono-cyclic aromatics in the pores. 46
Fig. 10.2 The XRD patterns of the Ni-Al alloy powder and home-made Raney Ni
catalyst.
Fig. 10.2 shows the XRD patterns of the Ni-Al alloy powder and home-made Raney Ni
catalyst. It can be seen that the XRD patterns of Ni-Al alloy is comprised of the two
336
categories of Ni3Al2 and Ni3Al domains. As the Al component dissolved by 20 wt. % NaOH
solution, the diffractions regarding metallic Ni were shown as amorphous nature. It was also
found that the home-made Raney Ni catalyst mainly presented diagnostic (111), (200), and
(220) diffractions of fcc Ni at 2θ of 44.5, 51.8, and 76.3°, respectively.47
The SEM images of Ni-Al alloy powder and home-made Raney Ni catalyst are displayed in
Fig. 10.3. The morphological difference between Ni-Al alloy powder and home-made Raney
Ni catalyst are readily visible from Fig. 10.3 (A) and (B). The home-made Raney Ni catalyst
is constituted by the typical fractured and angular particles, which is identical to other
research;48 whereas Ni-Al alloy powder only had the intact metallic structure. In addition, it is
observed that the small particles of home-made Raney Ni catalyst are dispersive, reaffirming
the high BET surface.
A
B
Fig. 10.3 SEM images of Ni-Al alloy powder (A) and home-made Raney Ni catalyst (B).
10.4.2 Product yield distributions from co-feed catalytic microwave pyrolysis
The product yield distribution from co-feeding of lignocellulosic biomass with plastics in the
337
catalytic microwave pyrolysis are summarized on the basis of catalytic temperature and
plastics to biomass mass ratio in Table 10.1. It was found that the mass yields of bio-oil and
gas were considerably impacted by the two variables in the range from 30.57 to 47.00 wt%
versus 36.18 to 49.56 wt%, respectively. The water content in the bio-oil varied from 10 to 30
wt% depending upon the reaction conditions. Without the addition of plastics in the catalytic
microwave pyrolysis, the water content could reach 30 wt%, while the water content from the
co-feed catalytic microwave pyrolysis was 10 – 20 wt%. It is evidenced that plastics-derived
olefins could react with furan compounds from the decomposition of cellulose and
hemicellulose to generate aromatics through Diels-Alder reactions followed by dehydration.24,
49
As a result, if the extent of dehydration reaction was enhanced in the co-feed catalytic
microwave pyrolysis; the increasing trend of water content in the bio-oil was observed. In
addition, the yield of olefins produced via catalytic cracking and subsequent cyclization for
the production of aromatics could also be influenced by the two independent variables.45
Accordingly, the catalytic temperature and plastics to biomass ratio play an essential role
bio-oil and gas yields. The maximum bio-oil yield was observed at the catalytic temperature
of 200 ºC with the plastics to biomass ratio of 0.75; while the maximum gas yield was found
at 500 °C with the ratio of 0.75.
Owing to the ex-situ catalysis, char derived from microwave-assist co-pyrolysis and coke
from catalytic step can be distinctly separated. It was also found that catalytic temperature
and plastics to biomass ratio affected both mass yields of char and coke as well. The char
yield was in the range from 9.09 wt% at the highest plastics to biomass ratio of 1.45 to 21.69
338
wt% without the introduction of plastics. Thus the char residue from the catalytic microwave
co-pyrolysis of lignocellulosic biomass with plastics was lower than that from biomass alone.
These results indicate that co-feeding of biomass with plastics in the catalytic microwave
pyrolysis was discerned to have a pronounced impact at the char formation. It is mainly
attributed to the fact that with in the surrounding of hydrogen from the direct decomposition
of plastics, the polymerization reactions of oxygenates were partially inhibited for the char
formation.26,
50
In addition, lignin as the main component of lignocellulosic biomass
generated reactive free radicals upon pyrolysis.51 These free radicals were relatively unstable
and could readily polymerize during pyrolysis; therefore, the free radicals became strong
hydrogen acceptors.52 Hydrogen evolved from plastics could be likely absorbed by free
radicals, and the polymerization and cross-linking reactions to form char were thus
suppressed.21 As a result, the formation of char was strongly reduced when biomass was
co-pyrolyzed with plastics. From other studies reported elsewhere, the HHV of the resulting
char could be obviously improved in the co-pyrolysis of biomass and plastics as well.50, 53
Since the coke deposited on the catalyst could cause the deactivation of zeolite-based
catalysts,40 the co-feed catalytic microwave pyrolysis should aid in reducing coke formation
and extending catalyst lifetime. The coke yield with regard to catalytic temperature and
plastics to biomass ratio is also outlined in Table 10.1, ranging from 1.02 to 3.85 wt%. As
expected, the coke formation was notably affected by the co-feeding of plastics. It can be
seen that the highest coke yield was obtained without the addition of plastics. The catalyst
coke was attenuated as the plastics was co-fed, declining to ~ 1 wt%. It has been reaffirmed
339
that in the presence of the well-promoted ZSM-5 catalyst, the resultant furan compounds
could react with plastics-derived olefins via Diels-Alder reactions rather than polymerization
reactions, thereby reducing the coke formation.8,
9
Furthermore, the hydrogen-deficient
oxygenates from the thermal decomposition of cellulose and hemicellulose could accept the
hydrogen from plastics degradation or olefins oligomerization, resulting in the decrease of
coke formation.26 For most phenolic compounds from the thermal degradation of lignin were
too large to enter zeolite pores;54 these phenolic compounds could easily deposit on the
catalyst surface and further polymerize to form coke.21 Conversely, these phenolic
compounds could abstract hydrogen atoms from evolutions of plastics during the co-feed
catalytic microwave pyrolysis. Hence these phenolic compounds were stabilized and the
polymerization reaction for the generation of coke could be inhibited. In the ANOVA analysis
based on CCD, the P-value from the four models (bio-oil, gas, char, and coke yield) are all
smaller than 0.05, implying that these models are appropriate to present the relationships
between all product yields and the two independent variables. The coefficient of
determination (R2) of these models are all more than 0.95, which evidences that these models
are favorable to mimic the relationships. Accordingly, these models can be employed to
predict the maximum product yields regarding the two independent variables.
340
Fig. 10.4 The overall carbon yields regarding product distribution from co-feed catalytic
microwave pyrolysis in light of catalytic temperature at the same plastics to biomass ratio
(0.75).
10.4.3 The effect of catalytic temperature on the co-feed catalytic microwave pyrolysis
As reported in our previous work, catalyst temperature played a vital in the product
distribution.40, 44 In this case, catalytic temperature was also used to investigate the product
distribution and evaluate the qualities of liquid organics. Representative results of product
distribution during co-feed catalytic microwave pyrolysis in light of catalytic temperature are
shown in Fig. 10.4. As expected, the carbon yield of char did not significantly vary alongside
the changes of catalytic temperature because of the ex-situ catalysis. Increasing the catalyst
temperature during co-feed catalytic microwave pyrolysis lowered the carbon yield of
341
catalyst coke. It was noted that direct degradation of plastics could produce a large amount of
long-chained waxes.21 However, these waxes with long carbon chins could be readily
adsorbed on the catalytic surface at the low catalytic temperatures, forming the waxy coke
that finally enclosed the whole catalyst. Meanwhile, phenolic compounds could also be
abstracted on the catalyst and stepwise polymerize to yield coke at the low catalytic
temperatures.23 When the elevated catalytic temperatures implemented, catalytic cracking of
waxes and phenolic compounds was favored, generating lighter compounds which were more
easily accessible to enter the mesopores. Consequently, the carbon yield of coke was reduced
at the expense of large molecules as the catalytic temperature was enhanced.
Since the degradation of plastics to generate olefins is an endothermic reaction, the yield of
olefins should go up as the catalytic temperature increased at the tested region.15 The
increased yield of olefins could strengthen the Diels-Alder reaction, contributing to the
increase of aromatics.23 Nonetheless, an excessively high catalytic temperature can favor
reverse Diels-Alder reactions to form light olefins instead of aromatics.8,
26
It was also
manifested that increasing catalytic temperature improved the decomposition of
biomass-derived oxygenates into small molecular, preventing the formation of liquid
aromatics.6, 8 As a result, the carbon yield of liquid organics gradually decreased due to the
enhanced catalytic cracking reactions in the catalytic temperature region.
On the contrary,
the carbon yield of gas monotonically promoted at the expense of liquid organics degradation
as the catalytic temperature went up, reaching 46.22% at 500 °C.
342
Table 10.3 gives a detailed carbon yield and selectivity on the basis of catalytic temperature.
In comparison with the liquid organics obtained by co-feed catalytic microwave pyrolysis at
200 °C, the aliphatic olefins and oxygenates were the primary species in the liquid organics.
This result indicates that such low catalytic temperature could not favor the zeolite-catalyzed
reactions to form aromatics. As the catalytic temperature was elevated to 375 °C, the carbon
selectivity toward aromatic hydrocarbons remarkably increased. There was still a low amount
of aliphatic hydrocarbons in the liquid organics (at 375 °C). Besides, there was a trace
amount of aromatic oxygenates (such as phenol) existing in the liquid organics, whose
content was much lower than those from catalytic microwave pyrolysis of biomass alone
under the same conditions.44 It is likely that plastics-derived olefins participated in the
Diels-Alder reaction with furan compounds followed by dehydration reaction to expel these
oxygenates. Aromatic hydrocarbons obtained were predominantly made up of toluene,
xylenes, trimethylbenzenes, naphthalene, and their derivatives. It was obtained that the
carbon selectivity toward toluene and xylenes ranged from 10.72 to 18.42% and 18.78 to
27.41%, respectively; whilst the carbon selectivity toward trimethylbenzene decreased from
17.13 to 9.01% as catalytic temperature went up from 375 to 500 °C. These outcomes
evidence that the dealkylation reaction of aromatics was also promoted at the elevated
catalytic temperature, leading to increased amounts of simple aromatics.15, 21 It is noted that a
trace amount of oxygenates was obtained at the catalytic temperature of 500 °C. It could be
inferred from these results that under such condition, hydrogen transfer reactions could
enhance the cracking of the side chains substituted on phenyl to yield aromatic
hydrocarbons.54 It is also possible that the stabilized phenolic compounds could be directly
343
deoxygenated via demethoxylation followed by dehydroxylation to obtain simple aromatic
hydrocarbons.21 To make the precursors match the jet fuels, these results suggest that the
aromatic hydrocarbons obtained at catalytic temperature of 375 °C are more desirable.
Table 10.3 Detailed carbon yield distribution and product carbon selectivity as a function
of catalytic temperature. a
Catalytic temperature ( ºC)
200
269
375
481
500
Liquid organics
42.23
40.84
40.54
36.22
34.65
Char
18.02
17.56
17.87
18.34
17.67
Coke
5.98
5.42
4.23
2.78
1.46
Gasb
33.77
36.18
37.36
42.66
46.22
Toluene
4.97
6.48
10.72
14.62
18.42
Ethylbenzene
1.18
1.63
2.29
2.66
1.99
p-xylene/m-xylene
2.50
7.67
18.78
25.27
27.41
Trimethylbenzene
4.35
16.74
17.13
10.41
9.01
Indane
0.39
0.72
-
2.24
2.00
Indene
1.60
1.30
0.79
1.91
1.59
Phenol
1.92
1.78
1.51
1.46
1.19
p-cresol/m-cresol
2.96
2.19
1.74
1.52
1.09
Naphthalene
5.51
3.90
5.33
6.96
6.43
1-methylnaphthalene
-
1.12
5.29
7.08
7.50
2-methylnaphthalene
-
0.97
2.13
0.90
0.95
Anthracene
-
-
0.85
1.29
1.28
Pyrene
-
0.12
-
0.09
0.24
Overall carbon yield (C mol%)
Liquid carbon selectivity (C mol%)
Gaseous carbon selectivity (C mol%)
344
Methane
12.45
13.11
13.65
14.02
14.25
Carbon monoxide
30.32
26.54
16.43
18.61
20.11
Carbon dioxide
22.45
20.45
18.98
20.10
20.32
Ethylene
18.45
24.20
31.45
36.54
40.22
Ethane
7.18
8.43
11.23
5.09
2.91
Propane
9.15
7.27
8.26
5.64
2.19
a Reaction
condition: Catalyst, 20 wt% with respect to feed; Plastics, 75 wt% with respect
to biomass; Reaction temperature, 480 ºC; Reaction time, 10 min.
b
Determined by difference
Non-condensable gas was another major co-product from co-feeding of biomass with plastics
in the catalytic microwave pyrolysis. The composition of gaseous fraction regarding catalytic
temperature is also listed in Table 10.3. Because of the co-feeding of biomass with plastics in
the catalytic microwave pyrolysis, oxygen content was expelled from furan compounds by
reacting with olefins through Diels-Alder reaction followed by dehydration reaction with the
water as product, instead of decarbonylation and decarboxylation reactions.7, 8, 49 Thus the
carbon yields of CO and CO2 are very low. However, it was observed that the carbon yields
of CO increased as the catalytic temperature went up from 375 to 500 °C. It is discerned that
higher catalytic temperature had a negative impact at the Diels-Alder reaction,21 resulting in
rejecting oxygen content from furans by decarbonylation reaction. It is noteworthy that the
carbon yield of ethylene dramatically increased as the catalytic temperature went up to
500 °C. These results indicate that increasing catalytic temperature favored the catalytic
cracking of long-chain waxes into light olefins. In addition, elevated catalytic temperature
also prevented the aromatization reaction of light olefins to form aromatic hydrocarbons. 10, 15
345
10.4.4 The effect of plastics to biomass ratio on the co-feed catalytic microwave pyrolysis
It was manifested that the optimal condition to obtain aromatic hydrocarbons lumped in the
jet fuels range was set at 375 °C. Since lignocellulosic biomass is more abundant and cheaper
than waste plastics, it is desired to enhance the mass ratio of cellulosic biomass in the
co-reactants if this cannot decrease the carbon yield of liquid organics substantially. In order
to investigate the effect of plastics proportion in the co-feed catalytic microwave pyrolysis,
the overall carbon yield of product distribution at 375 °C with respect to plastics to biomass
ratio is depicted in Fig. 10.5.
It can be seen that the carbon yield of liquid organics
increased nonlinearly with the augment of plastics to biomass ratio. Comparing with liquid
organics without using plastics as the co-reactant, the liquid carbon yield rapidly increased as
the ratio went up to 0.75. This increasing tendency indicates that there was a positive synergy
between the two reactants. Furthermore, the substantial trend was attributed to the fact that
large amounts of olefins with high H/Ceff ratio reacted with furan compounds by the
Diels-Alder reaction to enhance liquid organics yield.55, 56 According to our previous study,45
the highest yield (above 30%) of liquid organics from catalytic microwave pyrolysis of LDPE
alone is much less than the yields from co-feed catalytic microwave pyrolysis. Yet as the
plastics to biomass ratio went up from 0.75 to 1.45, the carbon yield of liquid organics
slightly augmented. Wherefore, this slight tendency suggests that the co-feeding of biomass
and plastics in catalytic microwave pyrolysis through the Diels-Alder reaction for the
enhancement of liquid organics yield could not take place to a larger extent or could even
reach the ceiling.
346
Fig. 10.5 The overall carbon yields regarding product distribution from co-feed catalytic
microwave pyrolysis as a function of plastics to biomass ratio at the same catalytic
temperature (375 °C).
Both the carbon yields of char and coke gradually declined with an increasing plastics to
biomass ratio as shown in Fig. 10.5. The carbon yields dramatically first decreased as the
plastics to biomass ratio went up to 0.75; whereas they represent slight augmented trends as
the ratio started 0.75 to 1.45. These could be explained by the fact that there was a
significantly synergistic effect between the two feedstock in the ratio range from 0 to 0.75;
whereas the synergy could not be further appreciably enhanced as more plastics were
employed. Hence, there exists an optimal plastics to biomass ratio (0.75) for the Diels-Alder
reaction in co-feed catalytic microwave pyrolysis if considering high value of waste plastics.
347
As stated above, the Diels-Alder reaction could not be further enhanced at the high ratios
(0.75 – 1.45), thus the light olefins derived from the catalytic cracking improved the carbon
yield of non-condensable gas.
Table 10.4 lists how the carbon selectivity toward aromatics changes with the addition of
plastics. In general, the carbon selectivity toward monocyclic aromatics went up as the
plastics to biomass ratio increased; whist the carbon selectivity toward polycyclic aromatics
gradually declined. For instance, the maximum carbon selectivity of aromatic hydrocarbons
was xylenes, which increased appreciably from 4.28 to 21.33%. As such, the second highest
carbon selectivity of aromatics was trimethylbenzene, displaying the same tendency as
xylenes from 8.32 to 16.25%. By contrast, the carbon selectivity toward polycyclic aromatics
experienced a remarkable drop with the increasing plastics to biomass ratio. It is due to the
reason that Diels-Alder reaction between aromatics and light olefins, or hydrogen could
suppress the polymerization reaction of monocyclic aromatics to generate polycyclic
aromatics.21, 22, 26 Meanwhile, light olefins alone were prone to the formation of monocyclic
aromatic hydrocarbons by using ZSM-5 catalyst in the catalytic microwave pyrolysis.
45
In
this regard, the aromatics with the carbon number in the jet fuel range were preferred from
the experiment conducted at the low plastics to biomass ratio.
The gaseous composition with regard to plastics to biomass ratio is also illustrated in Table
10.4. The carbon yield of methane retains at around 12% with the maximum yield at the ratio
of 0.75, whereas, the carbon yield of both CO and CO2 decreased with the increase of plastics
348
to biomass ratio. It is mainly attributed to the Diels-Alder reaction that was accelerated
between biomass and plastics, instead of the decarbonylation and decarboxylation reaction.
Unlike the CO and CO2, the carbon yield of ethylene was found to augment in the most
pronounced way from 5.32 to 56.34%. These outcomes verifies that the introduction of
plastics in the catalytic microwave pyrolysis enhanced the H/Ceff ratio; and biomass-derived
oxygenates
was
futher
transformed
into
more
aromatics
and
olefins.
Besides,
biomass-derived oxygenates could promote the degradation of large molecules to light olefins
in the co-feed catalytic pyrolysis.15
Table 10.4 Detailed carbon yield distribution and product carbon selectivity on the basis of
plastics to biomass ratio. a
Plastics to biomass ratio
0
0.4
0.75
1.1
1.45
Liquid organics
25.64
34.18
40.54
41.34
42.66
Char
40.55
27.27
17.87
14.88
11.25
Coke
7.53
5.65
4.23
3.87
3.34
Gasb
26.28
32.90
37.36
39.91
42.75
Toluene
2.85
4.04
10.72
8.09
10.82
Ethylbenzene
1.43
1.54
2.29
2.86
3.20
p-xylene/m-xylene
4.28
7.53
18.78
18.24
21.33
Trimethylbenzene
8.32
13.23
17.13
17.73
16.25
Indane
-
1.31
-
1.58
1.31
Indene
2.65
2.11
0.79
0.96
0.60
Phenol
4.53
1.92
1.51
1.11
0.80
Overall carbon yield (C mol%)
Liquid carbon selectivity (C mol%)
349
p-cresol/m-cresol
6.31
3.51
1.74
0.83
0.58
Naphthalene
8.02
7.78
5.33
4.16
3.23
1-methylnaphthalene
6.07
5.64
5.29
4.60
3.13
2-methylnaphthalene
3.13
2.63
2.13
1.86
1.54
Anthracene
1.82
0.45
0.85
0.91
0.24
Pyrene
0.35
0.34
-
0.21
0.12
Methane
12.94
13.22
13.65
12.44
11.68
Carbon monoxide
40.40
25.67
16.43
13.12
10.75
Carbon dioxide
28.12
22.36
18.98
14.39
11.25
Ethylene
5.32
23.19
31.45
46.84
56.34
Ethane
6.54
8.74
11.23
7.32
5.75
Propane
6.68
6.82
8.26
5.89
4.23
Gaseous carbon selectivity (C mol%)
a Reaction
condition: Catalyst, 20 wt% with respect to feed; Reaction temperature, 480 ºC;
Catalytic temperature, 375 ºC; Reaction time, 10 min.
b
Determined by difference
10.4.5 Process robustness: the recyclability of the well-promoted ZSM-5 catalyst
Irrespective of conversion extent of co-feed catalytic microwave pyrolysis, an appreciable
reaction condition was employed to determine the process robustness and reuse of the
well-promoted ZSM-5 catalyst. That is because the catalyst recyclability is of essential
importance for a heterogeneous process.57 Thus, the spent catalyst was regenerated to
evaluate the recyclability in the co-feed catalytic microwave pyrolysis. The conversions were
all conducted at the catalytic temperature of 375 °C with the plastics to biomass ratio of 0.75
as presented in Fig. 10.6. Since the co-feeding of biomass and plastics in the microwave
reactor was pyrolyzed under the same condition, the carbon yield of char did not significantly
350
change along with the recycle times of the catalyst. However, the carbon yield of liquid
organics was dramatically affected by the catalyst recycle times, showing a slight decrease
with the increase of recycle times. This result verifies that although coke deposited on the
catalyst surface could be removed, the mesopores were permanently blocked by the coke,
triggering the loss of active sites. Thus the Diels-Alder reaction between the biomass and
plastics for the improvement of liquid organics yield was inhibited as well. Interestingly, the
carbon yield of coke first decreased then increased. The decreased carbon yield of coke was
possibly due to the fact that the Diels-Alder cycloaddition was inhibited in the pores; yet the
catalytic cracking still took place on the surface of the catalyst active sites, which could
mitigate the polymerization to form coke. For the increased carbon yields, the active sites on
the catalyst surface were also be deactivated, leading to the waxes absorbed on the catalyst
that could not be cracked. According to the trend of gas yield, it is reaffirmed that the active
sites inside the catalyst pores was first deactivated by the pore blockage, the active sites on
the catalyst surface were subsequently inhibited with the increasing recycle times of the
catalyst.
Table 10.5 describes the detailed carbon yields and product carbon selectivity with respect to
catalyst reused times. It was found that carbon selectivity toward all aromatics species
decreased with the increasing reused times. Nonetheless, a large amount of aliphatic
hydrocarbons were detected in the liquid organics. The results indicate that the aromatization
and oligomerization of olefins to form aromatics were suppressed due to catalyst pores
blockages; the Diels-Alder cycloaddition between the co-reactants for improving aromatics
351
carbon selectivity was thus inhibited. It is noticed that both the carbon selectivity of CO and
CO2 first decreased and then went up; whilst the carbon selectivity of C2H4 showed an inverse
tendency, first increasing and then decreasing from the fresh to the fifth usage. It is
manifested that the catalytic cracking of both waxes and oxygenates occurred individually on
the catalyst surface, forming these relevant gas in the first phase (from the fresh use to the
third use). However, as the active sites was deactivated on the surface from the third use to
the fifth use, the catalytic cracking of waxes and oxygenates was suppressed as well.
Fig. 10.6 The overall carbon yields regarding product distribution from co-feed catalytic
microwave pyrolysis on the basis of catalyst recycle times at the catalytic temperature of
375 °C with the plastics to biomass ratio of 0.75.
352
Table 10.5 Detailed carbon yield distribution and product carbon selectivity with respect
to catalyst reused times. a
Recycle times
Fresh
2nd
3rd
4th
5th
Liquid organics
40.54
39.23
38.01
37.98
36.77
Char
17.87
18.02
17.75
18.22
17.68
Coke
4.23
3.88
3.35
4.34
4.89
Gasb
37.36
38.87
40.89
39.46
40.66
Toluene
10.72
8.18
7.84
6.23
5.96
Ethylbenzene
2.29
3.02
3.39
2.55
3.96
p-xylene/m-xylene
18.78
18.05
18.61
15.20
14.25
Trimethylbenzene
17.13
17.03
16.68
17.20
16.60
Indane
-
-
1.78
-
1.37
Indene
0.79
1.46
1.41
1.19
0.95
Phenol
1.51
1.58
1.60
1.69
1.96
p-cresol/m-cresol
1.74
2.07
1.98
1.99
2.27
Naphthalene
5.33
5.57
5.86
4.72
3.87
1-methylnaphthalene
5.29
5.38
4.66
3.23
3.15
2-methylnaphthalene
2.13
1.01
0.96
1.02
0.72
Anthracene
0.85
0.32
1.27
0.34
0.21
-
0.53
0.34
0.27
0.22
Methane
13.65
14.54
13.22
12.55
13.76
Carbon monoxide
16.43
16.33
15.65
20.11
22.44
Carbon dioxide
18.98
18.78
17.44
23.44
24.59
Ethylene
31.45
36.76
40.15
33.01
27.44
Ethane
11.23
7.45
6.24
4.56
6.35
Overall carbon yield (C mol%)
Liquid carbon selectivity (C mol%)
Pyrene
Gaseous carbon selectivity (C mol%)
353
Propane
a Reaction
8.26
6.14
7.30
6.33
5.42
condition: Catalyst, 20 wt% with respect to feed; Plastics, 75 wt% with respect
to biomass; Reaction temperature, 480 ºC; Catalytic temperature, 375 ºC; Reaction time,
10 min.
b
Determined by difference
10.4.6 Hydrogenation of liquid organics derived from co-feed catalytic microwave
pyrolysis for jet fuels
Given the Raney Ni catalyst could be poisoned by water in the hydrogenation process,36 the
water in the raw bio-oil samples should be removed. Since n-heptane severed as the vital
medium in the hydrogenation process,35 two phase of mixed bio-oil samples was separated by
the optimum solvent (n-heptane). The liquid organics mixed with n-heptane from every
samples were separated and weighed to measure the loss of organics. ~94 wt% of liquid
organics were existing in the organic phase by extracting all samples. Since liquid organics
produced from catalytic microwave co-pyrolysis (at 375 °C with the plastics to biomass ratio
of 0.75) were principally comprised of aromatics from C8 – C16, these aromatics are viewed
as precursors for civilian jet fuels. In addition, the run that obtained the maximum liquid
carbon yield (at 375 °C with the ratio of 1.45) and the run that gained the liquid organics
within lowest oxygenated content (at 500 °C with the ratio of 0.75) were also hydrogenated
for the optimization of jet fuels. As reported, naphthalene in n-heptane was completely
transformed into saturated decalin,35 thus the mass ratio of reactant to n-heptane was set at
1:7 in terms of adding certain mass of n-heptane.
The product distribution and carbon selectivity toward main alkanes from the hydrogenation
354
of diverse raw organics are elucidated in Table 10.6. After all the reactions, more than 99 vol%
of unreacted hydrogen was detected, indicating that the hydrogenation reactions were not
implemented under hydrogen-starved conditions. The overall carbon yields of these
hydrogenated organics (R-1, R-2, and R-3) with respect to the co-reactants of biomass and
plastics were 38.51, 32.92, and 40.53%, respectively. These carbon yields of hydrogenated
organics were much higher than that in our previous research 35. Compared with the overall
product distribution of hydrogenated R-1 and R-3, the result regarding the production of
cycloalkanes from hydrogenated R-2 was more superior to the others under the same
condition. It was observed that ~ 90% selectivity toward cycloalkanes was achieved from
hydrogenated R-2. The high amounts of high-density cycloalkanes can be potentially used as
the replacement of high-density jet fuels (e.g., JP-10 and RJ-5). For the production of
aliphatic alkanes, hydrogenated R-3 obtained the maximum amount, which is close to 30%.
In this sense, the contents of hydrogenated R-3 were identical to that in JP-5 navy fuel (31.23%
of aliphatic alkanes, 53.06% of cycloalkanes, and 15% of aromatic hydrocarbons).58 The total
amounts of hydro-aromatic hydrocarbons and aromatic hydrocarbons were comparable in all
hydrogenated organics, whose amounts satisfy the regulations (less than 15%) of current jet
fuels. A trace mount of other compounds were detected in all hydrogenated organics.
Likewise, lower than 1 vol% of small hydrocarbons (e.g., methane, ethane, and propane)
were achieved under the low-severity condition, reaffirming that there was almost no carbon
loss in the hydrogenation system.
355
Table 10.6 Products distribution
and partial alkanes’ carbon selectivity for hydrogenation of
diverse liquid organics.a
Organics speciesb
R-1
R-2
R-3
Aliphatic alkanes
14.04
2.15
29.24
Cycloalkanes
74.28
89.27
59.21
Hydro-aromatic hydrocarbons
3.03
3.24
1.12
Aromatic hydrocarbons
6.13
4.17
8.56
Others
2.52
1.17
1.87
1,4-dimethylcyclohexane
3.64
5.27
3.43
1,3-dimethylcyclohexane
12.38
17.26
12.67
1,2-dimethylcyclohexane
5.23
7.98
7.23
Ethylcyclohexane
3.02
3.87
4.05
Octane
0.75
0.17
0.98
1,2,4-trimethylcyclohexane
16.67
10.76
15.89
Propylcyclohexane
2.29
3.54
1.27
Nonane
0.88
0.14
1.72
Hexahydroindan
2.89
4.56
2.02
Decalin
5.76
7.87
3.71
Decane
0.58
-
1.02
Undecane
2.12
0.77
3.15
Dodecane
0.85
-
1.14
Tridecane
0.34
0.21
1.21
Tetradecane
0.63
0.15
1.32
Pentadecane
0.45
-
1.23
Overall selectivity (% in area)
Alkanes selectivity (C mol%)
356
Hexadecane
a
0.49
0.14
0.98
Reaction condition: Initial pressure, 500 psi; Raney Ni catalyst, 20 wt% with respect to
reactant mass; Reaction temperature, 200 ºC; Reaction time, 2 h.
b R-1:
reactant from the experiment conducted at catalytic temperature of 375 ºC with plastics
to biomass ratio of 0.75; R-2: reactant from the experiment conducted at catalytic temperature
of 500 ºC with plastics to biomass ratio of 0.75; R-3: reactant from the experiment conducted
at catalytic temperature of 375 with plastics to biomass ratio of 1.45.
It was discerned that hydrogenated R-3 attained the maximum carbon selectivity toward
monocyclic alkanes, except dimethylcyclohexane. Dimethylcyclohexane derived from the
hydro-cycloaddition of xylenes presented the highest carbon selectivity in hydrogenated R-2.
That is because the dealkylation reaction was facilitated at the high catalytic temperature
(500 °C), giving rise to high yield of xylenes. It is also noticeable that the maximum carbon
selectivity toward aliphatic alkanes was attained in hydrogenated R-3 and none of unsaturated
aliphatic hydrocarbons was found, suggesting that these high amounts of unsaturated
aliphatic hydrocarbons were completely transformed into saturated aliphatic hydrocarbons
under the mild reaction condition. Of the three raw organics for hydrogenation, the
hydrogenated organics derived from R-1 obtained the highest amounts of saturated alkanes
lumped in the jet fuel range. With the consideration of both reaction condition and product
distribution, the raw organics (R-1) is the optimal source to manufacture renewable jet fuels
with high carbon yield.
357
10.4.7 Reaction pathway for the conversion regrading co-feeding of biomass with
plastics into jet fuels
These observations are the key support to show the reaction route for the conversion
regarding co-feeding of lignocellulosic biomass with plastics into jet fuels. The resultant
observation of the hydrogenated organics is manifested that the reaction pathway was much
more complicated than those proposed previously.9, 14, 26 Based on the quantified products in
this research, related results from catalytic microwave-induce degradation of LDPE alone,45
and lignocellulosic biomass to jet fuel through combined catalytic conversions;40,
59
the
overall reactions network (co-feed catalytic pyrolysis and hydrogenation process) is outlined
in Fig. 10.7. For the dominant route in lignocellulosic biomass, cellulose went through a
sequence of dehydration, decarbonylation, and decarboxylation to generate furan compounds
in the thermal degradation.17, 40 As such, it was evidenced that hemicellulose was likely to
depolymerize into furan compounds,59 which was comparable with the result of cellulose
degradation
40.
Unlike cellulose and hemicellulose, lignin in lignocellulosic biomass was
primarily decomposed into phenolic compounds. As for the degradation of plastics in another
route, thermal degradation of plastics usually occurred through two mechanisms (random
scission and chain-end scission).60,
61
The two abovementioned mechanisms took place
simultaneously, generating free radicals together with the long carbon chains.62 At the same
time, the radical fragments could be transformed into straight chain hydrocarbons via
hydrogen transfer reactions.62 The hydrogen from the thermal degradation of plastics was
provided for biomass-derived oxygenates that acted as the strong acceptor, suppressing the
char formation.
358
The waxes with large molecule weigh from the thermal degradation of plastics subsequently
underwent catalytic cracking over ZSM-5 through two carbocationic mechanisms, giving rise
to light olefins.60,
63
These olefins thereafter reacted with furan compounds by the
Diels–Alder reaction followed by the dehydration reaction to form aromatic hydrocarbons.
Meanwhile, these plastics-derived olefins could individually subject to cyclization,
aromatization, and oligomerization reactions to obtain aromatic hydrocarbons.21 Likewise,
these furan compounds could individually go through decarbonylation, aromatization, and
oligomerization reactions inside the mesopores of well-promoted ZSM-5 to form aromatic
hydrocarbons as well. As previously reported, it is confirmed that the interaction via the
hydrocarbon pool mechanism could be found in addition to Diels–Alder reaction.9 Since the
carbon yield of liquid organics significantly increased and the coke yield dramatically
decreased, the Diels–Alder reaction was the dominant reaction pathway during the co-feed
catalytic pyrolysis in comparison with the hydrocarbon pool mechanism. In addition,
phenolic compounds were individually converted into aromatic hydrocarbons over ZSM-5
catalyst by the dehydration, cracking, and oligomerization reactions.17, 64
359
H
H
H
H
H
H
Fig. 10. 7 Proposed reaction pathways for the conversion of lignocellulosic biomass and
plastics into jet fuel range alkanes.
The liquid organics resulting from the co-feed catalytic microwave pyrolysis mostly gave rise
to a mixture of aromatic and aliphatic hydrocarbons. It was shown that the hydrogenation of
aliphatic olefins was the first step to take place in the hydrogenation system.65 The aromatic
360
hydrocarbons were stepwise hydrogenated into cyclic alkanes or hydro-aromatic
hydrocarbons through hydro-cycloaddition reactions in the presence of home-made Raney Ni
catalyst. Furthermore, the hydroisomerization reaction could took place among the
dimethylcyclohexanes; and a trace volume of small hydrocarbons was also generated by
hydrocracking reactions. As a result, these hydrogenated organics can be potentially used as
alternatives of civilian jet fuels or additives to manufacture desired jet fuels by mixing with
other hydrocarbons.
10.5 Conclusions
In summary, the microwave-induced co-pyrolysis followed by hydrogenation process is a
profound route for the improved carbon yield of jet fuel range alkanes from co-feeding of
lignocellulosic biomass with plastics. Two variables (catalytic temperature and plastics to
biomass ratio) were evaluated in the co-feed catalytic microwave pyrolysis to improve the
carbon yield of aromatics with the carbon number in the jet fuel range. The results showed
that there was a positive synergy between lignocellulosic biomass and plastics in the co-feed
catalytic microwave pyrolysis, which could significantly improve the carbon yield of liquid
organics and suppress the formations of char and coke. Although the optimal condition to
maximize the carbon yield (42.66%) of liquid organics in the step of co-feed catalytic
microwave pyrolysis was at 375 °C with the plastics to biomass ratio of 1.45, the compounds
with low carbon numbers (e.g. toluene) was also obtained, which cannot meet the
specifications of jet fuels. Yet the chemical compounds of liquid organics with high carbon
yield (40.54%), which was derived from the reaction conducted at 375 °C with the ratio of
361
0.75 primarily satisfy the specifications in the jet fuel range.
After the hydrogenation process by using home-made Raney Ni as the catalyst under a
low-severity condition, the overall carbon yield of three hydrogenated organics (R-1, R-2,
and R-3) regarding raw lignocellulosic biomass and plastics were 38.51, 32.92, and 40.53%,
respectively. Given the overall carbon yield and product distribution, the raw organics
conducted at the catalytic temperature of 375 °C with the biomass to plastics ratio of 0.75 in
the co-feed catalytic microwave pyrolysis is more suitable for the production of renewable
alkanes for advanced jet fuels (especially for high-density jet fuels). Herein, ~ 90%
selectivity toward jet fuel range alkanes was obtained from such raw organics, and up to 75%
selectivity belongs to high-density cycloalkanes.
It is more likely that in the future these
alkanes can be potentially used as additives or replacements for jet fuels.
362
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366
CHAPTER ELEVEN
TECHNO-ECONOMIC EVALUATION OF RENEWABLE
CYCLOALKANES FOR JET FUELS FROM
LIGNOCELLULOSIC BIOMASS
11.1 Abstract
This study reported a novel pathway to obtain renewable cycloalkanes for jet fuels from
diverse lignocellulosic biomasses and presents an economic analysis of the combined
processes in a small scale biorefinery based on the experimental results. The consecutive
processes principally proceeded via the catalytic microwave-induced pyrolysis of
lignocellulosic biomasses followed by hydrogenation process of raw bio-oil; the
non-condensable gas and char were also attained as the by-products in the catalytic
microwave-induced pyrolysis. In the bench-scale experiments, we observed that the
maximum carbon yield (24.76%) of desired aromatics in the carbon number range from 8 to
16 was achieved from catalytic microwave pyrolysis of hybrid poplar at 500 °C, comparing
with loblolly pine and Douglas fir used as feedstock. After the hydrogenation process, the
highest selectivity (95.20%) towards jet fuel range cycloalkanes was obtained. As expected, a
scenario using 15 dry tonne woody biomass per day is developed in a small scale biorefinery;
and the hybrid poplar is used as feedstock to manufacture 0.25 million gallons renewable
cycloalkanes for one year. The assessments showed that the plant including microwave
assisted ex-situ catalytic pyrolysis (MAECP) and hydrogenation (HG) systems is profitable.
367
The equipment costs have the largest contribution to the total capital investment, whereas the
feedstock and chemicals costs have the largest contribution to the total annual production cost.
It is estimated that the Return of Investment (ROI) is 30.33% for each year. Sensitivity
analysis indicates that liquid organics yield and cycloalkane selling price significantly impact
the ROI.
Keywords: Jet fuels; techno-economic analysis; microwave assisted ex-situ catalytic
pyrolysis; hydrogenation; cycloalkanes
11.2 Introduction
In the face of diminishing fossil resources, establishing energy security is essentially required
for sustainable development of society.1 With this context, the transportation engines
consumed the largest amounts of crude oil derived from fossil resources; the continuous
supplies of transportation hydrocarbon fuels thus render the global energy crisis.2 In light of
the report from EIA, the global petroleum consumption surpassed more than 40% growths
from 23.04 to 32.56 billion barrels per year during the period from 1980 to 2012. In the
United States, a total of 7 billion barrels of petroleum oils was consumed in 2010; meanwhile,
71% of crude oil was toward the integrated demand of gasoline, jet, and diesels fuels, whose
satisfactions were 5.2 billion barrels.1 Therefore, with careful consideration of substantial
increase in the consumption of transportation hydrocarbon fuels, the future availability of
petroleum resources is unpredictable.
368
On the other hand, a more sustainable replacement of conventional fuels evolved from
petroleum resources in the transportation engines with renewable alternatives was stipulated
in the Energy Independence and Security Art.3 It was reported that 36 billion gallons of
renewable biofuels are annually blended into the petroleum-derived hydrocarbon fuels by
2022. Nonetheless, only 14 billion gallons of renewable biofuels or drop-in fuels can be
manufactured in U.S., most of which are produced by means of corn-based ethanol and
triglyceride-derived biodiesel.1 Consequently corn and triglyceride derived biofuels cannot
meet the annual demand toward 36 billion gallons of renewable transportation fuels. Herein it
is urgent that new technologies leveraging abundant carbon-based sources to generate
renewable biofuels have to come to the market.
Instead, lignocellulosic biomass is considered as the most appropriate and long-term
substitution of petroleum resources due to its broad availability.4 The agricultural and forest
resources reserving in U.S. can potentially satisfy one-third of the country’s petroleum
demands.5 It was reported that there is an indicator that the demand of gasoline in the U.S. is
expected to decline by 22% till the end of 2040 because of the improved efficiency in
gasoline and electric vehicles 3. On the contrary, the demands for diesel and jet fuel
continuously increase by 27% in the forthcoming decades. From this perspective, it is
essential to shift renewable biofuels production toward diesel and jet fuels in the long term.
It is widely known that the conventional jet fuels derived from fossil resources are primarily
consist of linear-chain and branched-chain alkanes.6, 7 Yet, these linear-chain alkanes and
369
branched-chain alkanes are readily sensitive to pyrolytic cracking in the engines of aircrafts.8,
9
Conversely, cycloalkanes within robust ring strain are regarded as compact molecules,
which can contribute to a more dense jet fuel. In addition, the content of cycloalkanes can be
burned cleanly with high heats of combustion and low freezing points, comparing with the
low densities (~ 0.76-0.78 g/cm3) of straight-chain alkanes.8, 10 To overcome the shortage of
aliphatic alkanes, jet fuel range cycloalkanes or aromatic hydrocarbons should be synthesized
and added into commercial jet fuels (e.g. Jet A and JP-8).7, 11 Most efforts to produce military
jet fuels have concentrated on increasing the cycloalkanes content; for example, JP-5 navy
fuel contains 52.8% of cycloalkanes, 30.8% of aliphatic alkanes, and 15.9% of aromatic
hydrocarbons.12
To manufacture renewable jet fuel range compounds, Lei et al have accomplished a series of
investigations from lignocellulosic biomasses to the production of jet fuel range (C8 – C16)
aromatics through catalytic microwave-induced pyrolysis.13-15 It was observed that
aromatics-enriched bio-oils can be transformed into renewable cycloalkanes for high
energy-density jet fuels via hydrogenation reaction.16,
17
In our previous studies,
lignocellulosic biomasses were comparably converted into aromatics by catalytic microwave
pyrolysis over well-promoted ZSM-5 catalyst. The organic phase of the raw bio-oils derived
from catalytic microwave pyrolysis was thereafter extracted by the optimum solvent
(n-heptane). Huber et al. claimed that if noble metal catalysts can be substituted by normal
metal catalysts for hydrotreating processes, the minimum selling price of renewable jet fuels
could be significantly reduced.1
Therefore, Ni-based catalysts is more reasonable to be used
370
for hydrotreatment of raw bio-oils due to their high catalytic activities of hydrogenation and
low cost. As a result, the organic mixture was converted into desired hydrocarbons for jet
fuels in the final step, including cycloalkanes and minor aromatics, which satisfies basic
specifications of conventional jet fuels by using Raney Ni as the catalysts.
To date, research in microwave-induced pyrolysis has been concentrated on its applications to
handle wastes such as plastics, tires, agricultural wastes, wood wastes, and various organic
wastes. Despite the variety of research that has been conducted on microwave-induced
pyrolysis, the applications of microwave heating in the industrial scale is hampered by an
apparent lack of understanding in microwave systems and the technical basis on designing
commercial equipment for the pyrolytic system. In order to scale up the pyrolysis processes, a
number of studies analyzed the techno-economics of biomass fast pyrolysis to produce
bio-oils and upgrade to hydrocarbons fuels.18-22 However, there has been no available data
based on the energy usage or efficiency of microwave-assisted reactors to make a compelling
comparison in terms of energy balance, efficiency and economic analysis. In addition, limited
information is available on the economic evaluation of the microwave pyrolysis processes in
the literature in order to determine economic viability of the process.
To this end, the main objective is to investigate the technical and economic aspects of a
scale-up biorefinery developed for converting lignocellulosic biomass into renewable
cycloalkanes for jet fuels via integral processes. This study performed techno-economic
analysis of small scale mobile Microwave Assisted Ex-situ Catalytic Pyrolysis (MAECP)
371
system coupled with hydrogenation (HG) system based on current existing data and our
previous report.17 The inputs for the model include investment and financial assumptions,
processing capacity, biomass feedstock options, product options, operation costs, land costs,
and revenues.
Fig. 11.1 Simplified flow diagram of a biorefinery process for the production of cycloalkanes
for jet fuels.
Table 11.1
Proximate and elemental analyses of diverse lignocellulosic biomasses
Characteristics
Hybrid poplar
Loblolly pine
Douglas fir
sawdust
sawdust
sawdust pellets
Moisture
5.34
5.50
4.82
Volatile matter
75.87
78.23
76.08
Fixed carbon
17.36
15.74
18.89
Ash
1.43
0.53
0.21
Carbon
50.20
49.30
47.90
Hydrogen
6.02
6.02
6.55
Proximate analysis (wt%)
Elemental analysis (wt%)
372
Nitrogen
0.59
0.06
0.08
Oxygena
43.19
44.62
45.57
aby
difference
11.3 Process model description
The overall process for the production of renewable cycloalkanes for jet fuels from
lignocellulosic biomass is primarily divided into two systems, including Microwave Assisted
Ex-situ Catalytic Pyrolysis (MAECP) systems and hydrogenation (HG) system. Fig. 11.1
outlines simplified flow diagram of a biorefinery process for the production of jet fuel range
cycloalkanes. First, several experiments were conducted on catalytic microwave-induced
pyrolysis of various biomass sources (hybrid poplar, loblolly pine, Douglas fir) in the bench
scale. Table 11.1 details the proximate and elemental analysis of the three biomass sources.
Hybrid poplar have a higher carbon content and lower oxygen content than that of loblolly
pine and Douglas fir. Furthermore, the ash content in hybrid poplar is much higher than that
in loblolly pine and Douglas fir. These results indicate that the compositions in the feedstock
could have a significant influence in the production of raw bio-oils in the microwave-induced
catalytic system. Of the three biomasses, the optimal biomass species for maximizing the
organic yields was hybrid poplar as shown in Table 11.2, 11.3 and 11.4.
373
Table 11.2 Experimental design and product yield distribution of hybrid poplar.
Catalytic
Catalyst to
Temperature (ºC)
biomass ratio
H-1
300
H-2
Runa
Yield (wt.%)
Organics Gas
Char
Coke
0.15
18.65
40.54
22.59
2.68
300
0.35
17.38
42.87
22.22
3.23
H-3
450
0.15
15.65
48.84
21.52
1.27
H-4
450
0.35
14.44
50.34
22.88
2.25
H-5
269
0.25
19.21
42.22
22.59
3.57
H-6
375
0.25
17.20
44.82
22.54
2.00
H-7
375
0.25
17.43
45.01
23.17
2.03
H-8
375
0.25
17.53
45.21
22.34
2.12
H-9
375
0.25
17.11
44.76
22.56
1.98
H-10
375
0.25
17.33
44.88
21.89
2.23
H-11
481
0.25
15.46
48.74
21.98
1.54
H-12
375
0.11
18.23
43.21
22.20
1.89
H-13
375
0.39
16.21
46.02
22.91
2.43
H-14
500
0.25
14.87
50.05
23.01
1.27
H-15
375
0.50
15.89
47.23
22.63
2.98
H-16
-
-
22.07
39.20
22.52
-
a
H-1 to H-13 was conducted based on central composite design; H-14 and H-15 were
added as the controls; H-16 is the control in the absence of catalyst.
374
Table 11.3 Experimental design and product yield distribution of loblolly pine.
Catalytic
Catalyst to
Temperature (ºC)
biomass ratio
L-1
300
L-2
Runa
Yield (wt.%)
Organics Gas
Char
Coke
0.15
17.78
41.38
20.20
2.25
300
0.35
16.99
44.02
20.11
3.99
L-3
450
0.15
15.01
47.54
20.10
1.18
L-4
450
0.35
13.56
52.02
19.98
1.93
L-5
269
0.25
18.23
42.22
19.65
3.23
L-6
375
0.25
16.67
46.23
19.88
2.44
L-7
375
0.25
16.74
47.02
20.22
2.45
L-8
375
0.25
16.23
46.37
20.34
2.66
L-9
375
0.25
17.05
46.43
19.55
2.22
L-10
375
0.25
16.22
46.82
19.43
2.54
L-11
481
0.25
14.23
50.12
20.22
1.54
L-12
375
0.11
17.43
44.45
20.45
1.74
L-13
375
0.39
16.00
47.82
20.24
2.72
L-14
500
0.25
13.89
52.13
20.53
1.25
L-15
375
0.50
16.01
49.45
19.79
3.04
L-16
-
-
21.02
40.98
19.34
-
a
L-1 to L-13 was conducted based on central composite design; L-14 and L-15 were
added as the controls; L-16 is the control in the absence of catalyst.
375
Table 11.4 Experimental design and product yield distribution of Douglas fir.
Catalytic
Catalyst to
Temperature (ºC)
biomass ratio
D-1
300
D-2
Runa
Yield (wt.%)
Organics Gas
Char
Coke
0.15
17.88
39.45
22.25
3.03
300
0.35
17.11
40.25
21.50
3.98
D-3
450
0.15
15.48
46.20
21.85
2.02
D-4
450
0.35
13.88
49.66
21.63
2.65
D-5
269
0.25
18.67
40.20
21.78
4.02
D-6
375
0.25
16.88
44.02
22.54
2.82
D-7
375
0.25
16.77
44.34
22.78
2.94
D-8
375
0.25
16.96
44.21
22.07
2.78
D-9
375
0.25
16.54
44.02
21.74
3.02
D-10
375
0.25
16.82
44.01
21.56
2.98
D-11
481
0.25
14.65
47.89
22.21
2.44
D-12
375
0.11
18.00
42.88
21.85
2.71
D-13
375
0.39
15.99
44.23
21.90
3.55
D-14
500
0.25
14.02
49.11
22.05
1.95
D-15
375
0.50
16.32
45.56
22.12
3.90
D-16
-
-
20.88
38.18
21.25
-
a
D-1 to D-13 was conducted based on central composite design; D-14 and D-15 were
added as the controls; D-16 is the control in the absence of catalyst.
When pyrolysis without catalyst; the liquid organic yield of hybrid poplar went up to 22.07
wt%, exceeding those from loblolly pine and Douglas fir. It was also found that the
selectivity of aromatic hydrocarbons as the precursors of jet fuels ranged from 17.95 to
94.33%, 23.75 to 93.16%, 2.72 to 92.62% as catalytic temperature increased to 500 °C as
illustrated in Fig. 11.2, 11.3, and 11.4, implying that the elevated catalytic temperature
376
favored high selectivity towards aromatic hydrocarbons. It was reported that cellulose as the
feedstock for catalytic pyrolysis produced higher amount of aromatic hydrocarbons than
hemicellulose and lignin 23. It could also be inferred from these results that higher amount of
aromatic hydrocarbons from catalytic microwave pyrolysis of hybrid poplar was due to
higher cellulose content.
In addition, It was also notewothy that 24.76% carbon yield of
aromatics was obtained from catalytic microwave pyrolysis of hybrid poplar, which was
higher than those of loblolly pine (22.91%) and Douglas fir (21.92%) as described in Table
11.5. In general, hybrid poplar could be considered as the better resource for the production
of aromatics.
The organic composition from hybrid
popolar (%)
100
4.41
90
11.68
Sugars
80
21.44
Esters
70
Furans
60
29.16
50
94.33
83.71
40
30
Guaiacols
Phenols
66.18
19.62
Hydrocarbons
32.85
20
10
Other
Alcohols
17.95
ketones/aldehydes
0
Acids
Raw organics
269
375
481
Catalytic temperature(°C )
500
Fig. 11.2 The organic composition from hybrid poplar as a function of catalytic temperature
at the same biomass to catalyst ratio (0.25).
377
100
The organic composition from loblolly pine (%)
5.13
90
10.86
Other
Sugars
80
70
Esters
19
Furans
60
Guaiacols
50
40
26.09
93.16
82.7
27.36
Phenols
Hydrocarbons
57.17
30
Alcohols
20
10
34.37
ketones/aldehydes
23.75
Acids
0
Raw
organics
269
375
481
Catalytic temperature(°C )
500
Fig. 11.3 The organic composition from loblolly pine as a function of catalytic temperature at
the same biomass to catalyst ratio (0.25).
378
100
The organic compostion from Doulas fir (%)
5.91
12.84
90
Sugars
80
Esters
70
39.27
60
Furans
Guaiacols
50
83.62
25.54
40
92.62
Phenols
Hydrocarbons
30
Alcohols
40.61
20
10
Other
24.14
ketones/aldehydes
5.46
2.72
Acids
0
Raw
organics
269
375
481
500
Catalytic temperature(°C )
Fig. 11.4 The organic composition from Douglas fir as a function of catalytic temperature at
the same biomass to catalyst ratio (0.25).
Table 11.5 Products distribution and main aromatics carbon selectivity from diverse
lignocellulosic biomasses at 500 ºC. a
Biomass species
Hybrid poplar
Loblolly pine
Douglas fir
Gasa
42.34
43.08
40.53
Char
30.63
31.73
34.29
Coke
2.27
2.28
3.66
Overall carbon selectivity (C mol%)
379
Aromatics
24.76
22.91
21.52
Toluene
4.77
5.18
4.32
Ethylbenzene
2.60
2.39
1.79
p-xylene/m-xylene
13.34
12.03
12.54
Trimethylbenzene
5.22
4.05
4.37
Indane
2.49
2.08
2.99
Indane
5.86
4.05
5.06
Phenol
1.34
1.86
1.52
p-cresol/m-cresol
1.59
1.55
2.03
Naphthalene
17.83
13.59
10.77
1-methylnaphthalene
2.87
3.72
1.68
2-methylnaphthalene
14.20
16.70
12.30
Anthracene
0.78
1.59
1.25
Pyrene
1.20
0.26
2.40
Aromatics carbon selectivity (C mol%)
a by
difference
aReaction
condition: Catalyst, 25 wt% with respect to biomass; Reaction temperature, 480
ºC; Reaction time, 10 min.
For the utilization of biomass feedstock in the biorefinery, hybrid poplar enters the
microwave pyrolyzer operating at 480°C and atmospheric pressure (see Fig. 11.1). Inside the
microwave heating chambers, feedstock is moved forward by means of an auger. Standard
cyclones remove solids mostly consisting of char particles, which is entrained by the
pyrolytic vapors in the pyrolyzer. The pyrolytic volatiles from the microwave pyrolyzer pass
through a fixed bed reactor which is loaded with ZSM-5 zeolite catalysts. Then the vapors
can be condensed by a cooling system, obtaining aromatics-enriched bio-oils. Herein, the
condensation and distillation systems are carried out with recycling water or coolant to
380
minimize water usage in the whole process. The non-condensable gas formed is exerted to
preheat the feedstock and a small part of non-condensable gas is recycled as carrier gas for
purging the microwave pyrolyzer. Then the non-condensable gas was collected and sold as
combusting gas.
Table 11.6 Products distribution and partial cycloalkanes carbon selectivity as a function of
diverse biomass sources.
Biomass sources
Hybrid poplar
Loblolly pine
Douglas fir
Cycloalkanes
95.20
94.06
92.70
Cycloolefins
1.01
0.89
1.20
Hydro-aromatic hydrocarbons
0.23
0.22
0.34
Hydro-cyclic alcohols
2.67
3.12
3.56
Aromatic hydrocarbons
0.67
1.20
1.05
Other oxygenated aromatics
0.22
0.51
1.15
1,4-dimethylcyclohexane
4.23
5.48
1.52
1,3-dimethylcyclohexane
5.20
3.27
5.23
1,2-dimethylcyclohexane
2.34
1.53
2.11
Ethylcyclohexane
5.87
4.59
5.03
1,2,4-trimethylcyclohexane
3.06
2.29
1.36
Propylcyclohexane
1.32
0.88
1.79
Hexahydroindan
10.34
11.23
10.84
Decalin
16.76
16.34
17.05
Bicyclohexyl
1.23
1.65
1.32
Perhydrofluorene
1.89
2.05
1.34
Tetradecahydroanthracene
3.02
3.76
3.49
Overall selectivity (% in area)
Cycloalkanes selectivity (C mol%)
381
aReaction
condition: Initial pressure, 500 psi; Raney nickel catalyst, 20 wt% with respect to
reactant mass; Reaction temperature, 250 ºC; Reaction time, 2 h.
As parent oils derived from catalytic pyrolysis contain some water and the water content
could poison the catalysts in the hydrogenation process;14, 16 thus the water share must be
removed before entering the hydrotreater. The coupling of liquid-liquid extraction and
separation will be employed prior to the hydrogenation process. Since n-heptane is the
optimal medium that could assist in the hydrogenation of aromatics,17 n-heptane should also
be leveraged to extract the organics for the pilot-scale plant. The solvent can be reused after
the distillation and sent back to the extracting process. For the experimental result, the loss of
organics in water phase can be neglected if considering that all recoveries of organics from
different biomass sources could reach ~94 wt%. Accordingly, the organic mixture will
undergo the hydrogenation reaction in the hydrotreater. In the experimental results, it was
observed that 95.20% selectivity towards cycloalkanes was achieved from hybrid poplar as
depicted in Table 11. 6 at 250 °C with the initial hydrogen pressure of 500 psi. Compared
with the overall product distribution from loblolly pine and Douglas fir in the experimental
results, the result for producing jet fuel range cycloalkanes from hybrid poplar was more
superior to the others under the same condition. For the hydrotreated used in the pilot-scale
plant, a fixed bed reactor identical to the conventional one commonly used in
hydroprocessing of petroleum to finished fuels is employed. As such, the projected
temperature and pressure are held at 250 °C and 500 psi. The unreacted hydrogen is recycled
and compressed for the usage of hydrogenation process. Finally the effluent from the
hydrogenation process is sent to distillation decanter where the remaining n-heptane is
382
distilled and recycled back to extracting decanter. The renewable cycloalkanes for jet fuels
are finally achieved and safely stored prior to being delivered to the market.
11.4 Results and Discussion
11.4.1 Overall mass balances
This scenario is based on a small scale pyrolytic system with a 0.75 t/h capacity. The 0.75
t/h system can be readily mounted on a trailer and transited to other locations. The entire
system is expected to have at least a 10-year life time. The facility will be run 20 hours per
day and proceed 15 tonne woody biomass per day. It is assumed that operation time of the
facility is 329 days per year (equivalent capacity factor of 90%).19 The reactor mass balances
are calculated in terms of our previous experimental results.
According to the results,
assume that 15% liquid organics, 50% non-condensable gas and 23% char can be generated
by the MAECP system at 480°C for 10 min with the power of 700 W. It is also assumed that
the share of aromatics-enriched organics can be extracted from the bio-oils. Fig. 11.5
demonstrates the overall input-output analysis for the coupled processing units in the
biorefinery. The MAECP system delivers the bio-oils (3.85 tonne per day) to the extracting
process. The weight of solvent used for the liquid-liquid split is the same as the weight of
aromatics for each run. The hydrogen requirement for the hydrogenation process is 147.3 kg
for each day. After the hydrogenation and distillation (in the HG system) of mixed organics,
jet fuel range cycloalkanes (2.4 tonne per day or 245427 gallons per year) are obtained as the
final product for sale.
383
Fig. 11.5 Overall input-output analysis for MAECP and HG systems in this biorefinery.
11.4.2 Capital cost
Capital cost of MAECP and HG systems mainly include installed equipment cost. Table 11.7
shows a summary of estimated capital costs for required equipment in the integral processes
according to the current yields. Ruan et al.24 have proposed that the equipment cost of a direct
microwave pyrolysis system is $200,000 at 2008. For this MAECP system leveraged in our
proposed process, a fixed-bed reactor for the catalytic reforming costing $8,000 is introduced
to Ruan’s microwave pyrolysis system.24 The conservative contingency factor was for any
miscellaneous equipment, which may not be considered in the analysis due to its infancy of
development. Herein, a 20% contingency factor is applied for the base case, which is the
same as other research
19.
Insurance and taxes are assumed as 1.5% of the total installed
equipment cost.
It is well known that capital costs do not have a linear relationship with the plant size. In fact,
there is a strong declined economy of the scale as plant sizes are reduced. The cost for
384
installed equipment are on the basis of the cost correlations explained by Xing et al.6 and
Jones et al.25 The scale is adjusted to the appropriate scaling term by using the following
equation:
B
E
=B
× d
B
e
E
E
D g
f
D
Where “n” is the scale factor, typically, 0.6 to 0.7.
Table 11.7 Summary of estimated capital costs for required equipment in the integrated processes.
Item
Amount ($)
Total installed cost
Cost of installed equipment
208,000
Project contingency
20% Contingency fund
41,600
Other costs
Permits, documents, etc.
1,500
Insurance and taxes
1.5% of the total installed equipment cost
3,120
Liquid-liquid extraction
Extracting decanter
10,000
Hydrogenation
Hydrotreater
87,000
Distillation
Distillation decanter
5,000
Project contingency
20% Contingency fund
20,400
Insurance and taxes
1.5% of the total installed equipment cost
1,530
MAECP system
HT system
Total capital invest
378,150
In the HG system, there is 0.25 million gallons cycloalkanes for jet fuels can be produced in
one year; however the capacity of the system should be doubled due to the solvent introduced
in the process. As we assume, the same volume of solvent is employed to extract liquid
organics from the parent bio-oils. Therefore the capacity of installed equipment is to produce
0.5 million gallon product. Xing et al.6 have proposed specific information with regard to the
385
installed capital cost of the biorefinery scale for producing 0.5 million gallons alkanes per
year. Wherefore the installed capital cost for our biorefinery is comparable with the
hydrogenation and distillation systems used by Xing et al. In addition, some equipment, such
as pumps and heat exchangers, are not currently considered and will be added in our future
analysis when more detailed information is available.
Table 11.8 Summary of operating costs involved in the processes from hybrid poplar to
cycloalkanes.
Inputs
Value
Amount
($/unit)
($/year)
Quantity
MAECP system
Feedstock
4935
83
409,605
Chemicals (ZSM-5)
123.5
1,600
197,400
Electricity purchased
—
—
27,648
Ni-Al alloy powder
14.8
1500
22,207
Hydrogen
48.5
2000
96,923
NaOH
5.48
300
1,644
n-heptane
10
2500
25,000
Transportation of cycloalkanes to
790
4
3,160
Electricity purchased
—
—
17,000
Process labor
6580
3×12/h
236,880
Machinery depreciation
—
—
37,000
Maintenance
—
—
13,000
HT system
Market
Others
Total operating costs
1,087,467
386
11.4.3 Operating cost
Operating costs are comprised of all variable and fixed costs. Table 11.8 indicates the detailed
variable and fixed operating costs. Variable costs are expenses that usually change with the
varied amounts of the products. Variable costs included inputs costs, electrical and
maintenance, and other miscellaneous expenses. The price of feedstock within the feedstock
procurement area can remarkably impact cash flow of the system. The assumed feedstock
cost of woody biomass is 83 $•t-1.20 Chemicals (commercial ZSM-5 zeolite catalyst) used for
catalytic reforming could be regenerated by means of burning off the coke deposited on the
catalyst using air combustion.26 The price of commercial ZSM-5 zeolite is 1.6 $ • kg-1
according to the price in December 2010.20 The price of parent Ni-Al alloy powder is 1500 $•
t-127. In the manufacturing process, the yield of Raney Ni catalyst is 50 wt% in the presence
of NaOH, whose price is 300 $•t-1. Since the solvent can recycled and reused, loss of the
solvent is negligible; thus assuming only 10 ton solvent will be applied in the biorefinery.
Hydrogen is supplied from an external source, and its price varies depends upon the location
of a new plant to be built;6 the price of hydrogen considered in our project is 2000 $•t-1
according to the calculation reported elsewhere 1. Electricity is purchased for both MAECP
and HG systems.24 The transportation cost of renewable cycloalkanes for jet fuels was
estimated on a cost-per-ton basis, which is 4.00 $•t-1 according to Petrolia (2008).28 Fixed
operating costs, consisting of processing labor, maintenance, and machinery depreciation are
also determined. The number, expertise, and salaries of personnel required to operate the
facility is estimated, which is equal to that in Ruan’s 2008 report.24 Given all of the
aforementioned expenses, total operating cost is calculated to be 1,087,467 $/year.
387
11.4.4 Economic assessment of the biorefinery process
In this study, these renewable cycloalkanes can be directly used as additives to make the
desired jet fuels by blending with other hydrocarbons. According to the price (4.6 $/gallon) of
regular Jet A,29 we assume that these jet fuel range cycloalkanes are sold as regular jet fuels.
The char produced by the facility is sold for 20 $•t-1, a value with respect to the 2010 price of
coal and char’s lower relative heating value.20 After preheating the biomass, the cool
non-condensable gas are collected and sold as natural gas. The price for natural gas is 5
$/MMBTU in terms of the average from the U.S. Energy Information Administration’s (EIA)
Annual Energy Outlook’s 20-year price forecasts.21 The income tax rate is assumed as 35 %,
which is comparable with other reports.1, 21 Table 11.9 summarizes the detailed incomes from
this designed biorefinery.
Table 11.9 Annual income from the sales of final products.
Income
Value
Amount
($/unit)
($/year)
Quantity
Sale of cycloalkanes (gallon)
245,427
4.6
1,128,964
Sale of Char (dry tonne)
1,135
20
22,701
Sale of gas (MMBTU)
22445
5
112,226
Total income
1,263,891
The profit of the MAECP system and HG system is figured out in light of the difference
between the products income and operating cost. The systems could annually make a profit of
$176,424 and net revenue of $114,676 when feedstock price is 83 $•t-1 and the selling price
for cycloalkanes, char, and non-condensable gas are 4.6 $/gallon, 20 $•t-1, and 5 $/MMBTU,
388
respectively. Therefore the Return of Investment (ROI) is 30.33% per year or the investment
can be recovered in 3.3 years. The integral systems have generated profitable projections with
respect to the economic feasibility; however, the processes still need further investigation to
assure whether the profit could be improved.
Feedstock price (75%, 100%, 125%)
Cycloalkanes selling price (4.2, 4.6, 5.0 $/gallon)
Organics yield (14wt%, 15wt%, 16wt%)
ZSM-5 price (75%, 100%, 125%)
Capital cost (75%, 100%, 125%)
Hydrogen price (75%, 100%, 125%)
Gas selling price (4, 5, 6 $/MMBTU)
Char selling price(10, 20, 30 $/t)
Income tax rate (30%, 35%, 40%)
0.00
10.00
20.00
30.00
ROI (%)
40.00
50.00
60.00
Fig. 11.6 Sensitivity analysis on return of investment (ROI).
11.4.5 Sensitivity Analysis
Key variables are selected according to their potential effect on the ROI. Although the
microwave-assisted pyrolysis process could reduce the cost of feedstock preparations, our
greatest sensitivity still depends on the price of feedstock. The coupling of catalytic
microwave pyrolysis and hydrotreating process to generate jet fuel range cycloalkanes is a
developing technology, and the yields and selling prices of jet fuel range cycloalkanes are the
389
particularly significant variables in this project. Capital cost is another important sensitivity
variable due to the uncertainties associated with the estimate in this study. A detailed analysis
of the sensitivity is shown in Fig. 11.6, with the feedstock price ranging from 75 to 125% of
the current price, cycloalkanes selling prices ranging from 4.0 to 5.2 $/gallon, organics yield
from 14 to 16%, ZSM-5 zeolite price varying from 75 to 125%, capital cost changing from 75
to 125% etc. Sensitivity analysis results described indicate an enormous impact from liquid
organics yield and cycloalkane selling price. Slight improvement in the liquid organics yield
and selling price of cycloalkanes could substantially increase the ROI whereas lower organics
yield and selling price of cycloalkanes result in a rapid decrease of ROI.
11.4.6 Limitations of the Analysis
The current analysis has several limitations, many due to the lack of current large scale
operation data as well as the conceptual nature of the analysis. The following is a summary of
the most significant limitations of this analysis.
(1) Proposed design for bio-jet fuel production from MAECP system and HG system has not
been tested in a pilot scale.
(2) The process of catalysts preparation and regeneration are not been considered.
(3) How to deal with the aqueous phase from the bio-oils has not been provided because of
the lacking indication.
(4) Lack of criteria for cycloalkanes price, since biomass-derived renewable cycloalkanes for
jet fuels have not gained acceptance of existing markets
(5) The liquid organic yield from the pilot scale microwave pyrolyzer of hybrid poplar has
390
not been ascertained.
11.5 Conclusions
Preliminary economic analysis for the integrated processes was conducted according to the
simplified process flow diagram and material balances with the consideration of microwave
assisted ex-situ catalytic pyrolysis (MAECP) system and (hydrogenation) HG system. From
the experimental result, the optimal condition to maximize the C8 - C16 aromatics in the step
of catalytic microwave pyrolysis was at 500 °C with the catalyst to biomass ratio of 0.25.
These aromatics in n-heptane medium were thoroughly hydrogenated into jet fuel range
cycloalkanes at 250 °C. Furthermore, the integral processes detailed here were illustrated to
deliver up to 95.20% selectivity towards jet fuel range cycloalkanes from hybrid poplar under
very mild conditions, which were superior to those of loblolly pine and Douglas fir. The
cycloalkanes can be directly used as additives to make the desired jet fuels by blending with
other hydrocarbons.
Based on the experimental results, the techno-economic analysis shows that the coupled
system can generate a profit of 176,424 $/year. The installed equipment cost and operating
cost were calculated based on the target product capacity (0.25 million gallons per year). We
estimate that the Return of Investment (ROI) is 30.33% per year or the investment can be
recovered in 3.3 years. Sensitivity analysis shows that liquid organics yield and cycloalkane
selling price significantly affect the ROI. Slight improvement in the liquid organics yield and
selling price of cycloalkanes could substantially increase the ROI. To summarize, these
391
integrated processes within inexpensive catalysts under mild conditions delivers a novel and
feasible route that specifically targets jet fuel range cycloalkanes from lignocellulosic
biomass in a biorefinery.
392
11. 6 References
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394
CHAPTER TWELVE
THERMAL BEHAVIOR AND KINETIC STUDY FOR CATALYTIC
CO-PYROLYSIS OF BIOMASS WITH PLASTICS
12.1 Abstract
The present study aims to investigate the thermal decomposition behaviors and kinetics of
biomass (cellulose/ Douglas fir sawdust) and plastics (LDPE) in a non-catalytic and catalytic
co-pyrolysis over ZSM-5 catalyst by using a thermogravimetric analyzer (TGA). It was found
that there was a positive synergistic interaction between biomass and plastics according to the
difference of weight loss (∆ ), which could decrease the formation of solid residue at the
end of the experiment. The DTG curves demonstrated that the addition of catalyst tended to
reduce the degradation temperature. The first order reaction model well fitted for both
non-catalytic and catalytic co-pyrolysis of biomass with plastics.
As for the activation
energy (E) and pre-exponential factor (A), the addition of plastics in the non-catalytic
pyrolysis of biomass could appreciably decrease the activation energy (E). Moreover, the
presence of ZSM-5 catalyst could further decrease the activation energy (E) based on the
result of the blends without catalyst. The activation energy (E) of Cellulose-LDPE-Catalyst
and DF-LDPE-Catalyst are only 89.51 and 54.51 KJ/mol, respectively.
The kinetics
analysis showed that adding catalyst doesn’t change the decomposition mechanism. As a
result, the kinetic study on catalytic co-pyrolysis of biomass with plastics was suggested that
the catalytic co-pyrolysis is a promising technique that can significantly reduce the energy
395
input.
Keywords: Catalytic co-pyrolysis; kinetics; TGA; biomass; plastics
12.2 Introduction
Petrochemicals (aromatics and olefins) produced from petroleum sources are crucial
feedstock for the manufacture of valuable products, such as solvents, synthetic fibers, plastics,
and pharmaceuticals.1 Meanwhile, benzene, toluene, and xylenes (BTX) with the light olefins
(ethylene, propylene, and butadiene) are the most important preliminary petrochemicals.2 As
stated above, CFP of biomass over zeolite-based catalysts is a promising technique that can
convert biomass directly into high-value added aromatics and olefins. Yet many issues
accompanied with valuable products have also been encountered. For example, coke is
rapidly and dramatically formed in CFP and deposits on the catalyst, block the micropores
and consequently leads to the deactivation of the catalysts.3 It is worth noting that a relatively
low carbon yield (usually 10 – 30%) of aromatics were observed in CFP. In the liquid
products, a large amount of aromatics belong to polycyclic aromatic hydrocarbons (PAH).
The high concentration of PAH limit the utilization of liquid product since PAH are well
known to be carcinogenic, mutagenic and environmentally unfriendly.4
To increase the petrochemicals yield, improve the aromatic selectivity, and minimize the coke
formation, co-feed catalytic pyrolysis of biomass with hydrogen-enriched feedstock (such as
396
plastics and alcohols) is a novel and feasible route to solve these problems. It was reported
that the co-feeding of alcohols5 and light olefins (ethylene and propylene)6 in the catalytic
pyrolysis of biomass could enhance the yield of petrochemicals over ZSM-5 catalyst. Plastics
as the ideal co-reactant are usually added in the catalytic pyrolysis of biomass. Since waste
plastics can cause serious environmental issues and threaten human health, the utilization of
waste plastics in catalytic co-pyrolysis process is not only beneficial for the reduction of
pollutants, but also in favor of energy recovery. Numerous studies reveal that the co-feed
catalytic pyrolysis of biomass with plastics could enhance petrochemicals yield, improve the
selectivity of desired aromatic compounds, and reduce the coke formation.7-13 These results
evidence that there is a positive synergetic effect between biomass and plastics in the co-feed
catalytic pyrolysis process. According to state of the art research, most studies reported the
results of hydrocarbon plastics (e.g., PE, PP, and PS) mixed with biomass in catalytic
co-pyrolysis using HZSM-5 as the catalyst, and observed a positive synergy in the
improvement of aromatic yield and selectivity and the reduction of coke generation.
Therefore, catalytic co-pyrolysis of lignocellulosic biomass with polymers should be
advocated as the economically feasible route for the utilization of both lignocellulosic
biomass and polymers in the future.
Understanding the kinetics of non-catalytic co-pyrolysis and catalytic co-pyrolysis is vital to
the design, optimization, and scaling up of industrial biomass conversion applications. In
general, biomass is comprised primarily of cellulose, hemicellulose, and lignin. Each of the
structural constituents of biomass pyrolyzes with different rates, mechanisms and pathways.14
397
Therefore the kinetics of thermal decomposition of biomass materials involves a large
number of reactions in parallel and series, making it complicated. Pyrolysis models, which
can be classified into three principal categories: single-step global models, semi-global
models, and multiple-step models.15 Single-step global reaction models usually describe the
overall rate of devolatilization from the biomass substrate. Semi-global models assume that
biomass pyrolysis products are separated into three distinct fractions: volatiles, tars, and char.
Multiple-step models are more complex than single-step global and semi-global models. For
example, the primary pyrolysis of cellulose was a simple first-order endothermic reaction
with activation energy of 238 kJ/mol was reported.16 "Broido-Shafizadeh models" has been
proposed to predict the evolution of each main product fraction.17 It was observed that the
pyrolysis of cellulose consisted of three consecutive first-order reactions: the first reaction
releases approximately 30% volatiles; the second reaction released practically no volatiles;
while the third reaction releases the remaining 70% of the volatile matter.
Thermogravimetric analysis (TGA) is the most commonly applied thermo-analytical
technique for thermal study of biomass pyrolysis.18 TGA measures the decrease in substrate
mass caused by the release of volatiles during thermal decomposition as a function of time.19
The pyrolysis kinetics of many kinds of biomass have been widely studied using
thermogravimetric analysis method. The kinetics study of catalysis pyrolysis has been studied
for several kind of biomass and catalyst, such as tobacco rob mixed with catalyst (dolomite
and NiO),20 corn stalk with sodium carbonate or potassium carbonate as catalyst,21 wheat
straws with three kind of catalyst (solid acid catalysts, bifunctional catalysts, and industrial
398
catalysts).22
In the present work, the non-catalytic co-pyrolysis and catalytic co-pyrolysis behaviors of
biomass with plastics were investigated under inert atmosphere using a thermogravimetric
analysis to gain overall understanding of the interactions between biomass and plastics. Even
though the non-catalytic kinetics have been reported, the catalytic co-pyrolysis of biomass
with plastics has never be proposed. Hence this study would fill the knowledge gap of
catalytic co-pyrolysis. In this regard, the thermal events occurring during co-pyrolysis of
biomass with plastics were identified and the kinetic data were obtained to fit
thermogravimetric data, the global process as being considering one to four consecutive first
order reactions.
12.3 Experimental
12.3.1 Materials
Douglas fir sawdust pellets (7 mm in diameter and 15 mm in length) were leveraged as
lignocellulosic biomass, which were supplied by Bear Mountain Forest Products Inc., USA.
The cellulose (CAS number 9004-34-6) is purchased from Sigma-Aldrich Corporation (St.
Louis, MO, USA). Cellulose is in the form of microcrystalline powders and particle sizes of
cellulose are averaged at 50 µm. Low-density polyethylene pellets (LDPE) was purchased
from Sigma-Aldrich Corporation (St. Louis, MO, USA). Its density and melting point are
0.925g/cm3 and 116 °C, respectively. Parent ZSM-5 (SiO2/Al2O3 Mole Ratio: 50) was
supplied by Zeolyst International, USA. The woody biomass and LDPE pellets were all
399
grinded into small particle size (1-2 mm).
12.3.2 Catalyst preparation
The parent ZSM-5 was fabricated by hydrothermal and calcined treatments. Parent ZSM-5
powder was added into deionized water (mass ratio=1) at 60 °C under the gentle stirring. The
slurry was kept at this condition for 2 h. The slurry was subsequently dried at 105 °C until the
weight was constant. The sequential process was calcination: parent ZSM-5 tailored by
hydrothermal treatment was calcined at 550°C for 5 h in a muffle furnace. The obtained solid
was pelletized and sieved to 100 mesh.
12.3.3 Thermogravimetric analysis
The thermal degradation behavior of biomass (cellulose and Douglas fir sawdust) with LDPE
was evaluated by a thermogravimetric analyzer (Mettler Toledo 188 TGA/SDTA 851,
Switzerland). For each test, approximately 5 mg sample was placed into an alumina crucible
and heated from 25 to 600 °C by a heating rate of 20 °C/min with nitrogen flow rate of 50
ml/min. As for the measurements of non-catalytic co-pyrolysis, biomass (cellulose/Douglas
fir sawdust) with LDPE powder was first blended as the mass ratio of 4:1. And then, about 5
g mixture was loaded into the crucible to conduct the TGA test. In addition, the same
thermogravimetric analyzer was used to perform the thermogravimetric analysis for catalytic
co-pyrolysis of feedstock mixture with ZMS-5 powder. Herein, the mixture to catalyst mass
ratio was also set as 4:1.
400
12.3.4 Kinetic study
For the TGA assays, the evolution with temperature of weight loss (TG) and the weight loss
rate (DTG) were gained co-pyrolysis.23-27 The DTG was calculated by the expression:
=−
1
'
h
d
f
(1)
To illuminate the synergistic effect between the co-reactants, the difference of weight loss
(∆
) was often defined on the basis of the synergistic effect of each material during
pyrolysis.23-25, 27, 28
∆
=
ijkgl
− (m
+m
)
(2)
By using integral method, the kinetic parameters, activation energy, and pre-exponential
factor of co-pyrolysis of biomass with polymers can be determined.
23-31
Moreover, it was
assumed that the co-pyrolysis is a first order reaction. 23-29 The kinetics calculation was based
on the Arrhenius equation. Therefore, the reactions regarding the co-pyrolysis of biomass
with polymers can be explained as the following formula:
m
= n mE d−
#
f (1 − m )
op
(3)
Where A is pre-exponential factor; E is activation energy; T is temperature; t is time; m is
weight loss fraction or co-pyrolysis conversion which can be calculated by the equation.
m=
−
'−
'
h
(4)
q
Where W0 is the initial mass of the test sample; Wt is the mass at time t and Wf is final mass at
401
the end of co-pyrolysis. For a constant heating rate H during co-pyrolysis, H = p ⁄
,
rearranging the above equations and integrating gives:
− (1 − m )
s
t=
p
u
no
2op
#
d1 −
fw −
v#
#
op
(5)
For most values of E and for the temperature range of co-pyrolysis, it is manifested that the
expression
[no/v# (1 − 2op/#)] in the proposed equation is essentially constant.
Hence, if the left side is plotted against 1/T, a straight line will be achieved. From the slope,
-E/R, the activation energy E can be figured out. In addition, by taking the temperature at
which Wt = (W0 + Wf)/2 in the place of T in the intercept term of the above-mentioned
equation, the pre-exponential factor A can also be determined.
12.4 Results and discussion
12.4.1 Pyrolysis of materials and their blends
Thermogravimetric analysis showed the relationship between the weight change of a sample
and temperature, playing a significant role in future understanding of the thermal
decomposition and reaction mechanism during the pyrolysis. TG curves indicates the mass
loss of the sample versus temperature change of the thermal degradation, and DTG presents
the corresponding rate of mass loss of TG curves. Fig. 12.1 shows TG curves of isolated
materials (biomass and LDPE) and their mixture. The TG curves of cellulose with LDPE are
sketched in Fig. 12.1 (A). It can be seen that the thermal degradation of cellulose took place
in 300 – 400 °C; while the thermal decomposition of LDPE occurred at 400 – 500 °C. It was
402
found that LDPE was completely degraded without the formation of solid residues. When 20
wt% LDPE was mixed with cellulose, the degradation temperature of the blends decreased to
300 °C, closing to the degradation of cellulose. However, the mixture with the catalyst was
quite different from the others. The decomposition mostly occurred at 400 – 450 °C. These
results indicate that the presence of both LDPE and catalyst have a strong effect on the
degradation temperature.
The TG curves of Fig. 12.1 (B) is extremely similar to the Fig. 12.1 (A). Likewise, Douglas
fir sawdust mainly decomposed at 300 – 400 °C. In general, the degradation curve can be
divided into three stages. The first stages (20 – 250 °C) resulted from moisture evaporation.
The second stage (300 – 400 °C) caused the main weigh loss when the main pyrolysis
occurred by devolatilization, and most organic compounds (cellulose, hemicellulose and
partial lignin content) were degraded at this stages. The third stage (over 400 °C)
corresponded to the continuous devolatilization with small partial lignin and charring. As
such, the TG curve of DF and LDPE mixture is comparable with that of cellulose and LDPE
mixture. Additionally, the curve of DF-LDPE-Catalyst was identical to that of
Cellulose-LDPE-Catalyst.
403
Fig. 12.1 TG curves of cellulose with LDPE (A); and Douglas fir sawdust (DF) with LDPE
(B).
Fig. 12.2 shows the DTG curves of the different samples. From Fig. 12.2 (A), both the
404
decomposition of cellulose or LDPE alone showed an independent peak at 300 – 400 °C or
400 – 500 °C, respectively. Interestingly, the DTG curve of cellulose and LDPE mixture
presented two peaks, which is attributed to the different characteristics of cellulose and LDPE;
while, with the presence of ZSM-5 catalyst, the high peck shifted to a lower temperature. It
can be seen that the high peak shifted to around 450 °C, which is low than that of cellulose
and LDPE mixture. It is affirmed that the catalyst could significantly decrease the
decomposition temperature.
The first small peak on the DTG of DF decomposition curve as shown in Fig 12.2 (B)
appeared at about 320 °C, which should be attributed to the time when the maximum
decomposition rate of hemicellulose occurred. The large peak on the DTG of DF
decomposition emerged at about 380 °C, corresponding to the time when maximum
decomposition rate of cellulose reached. However, the thermal degradation of lignin occurred
in a wide temperature range from about 200 – 500 °C. Thus, it did not have an obvious
behavior of lignin as the other two major compositions of biomass. This is in good agreement
with previous results that the temperature range of hemicellulose decomposition is from 190
– 380 °C, and the temperature range for cellulose decomposition is from 250 to 380°C, with a
maximum mass weight loss at about 350 °C.32 It was also observed that the DTG curve of DF
and LDPE mixture slightly shifted to a lower temperature zone in comparison with DTG of
DF or LDPE alone, which can be seen from the DTG peaks positions. These results
suggested that there is a synergistic effect between biomass and plastics during pyrolysis.
405
In order to investigate the effect of catalyst on thermal decomposition of DF and LDPE
mixture, the blends was mixed with catalyst as the mass ratio of 4:1. The DTG curve of
DF-LDPE-Catalyst curve can also been seen in Fig. 12.2 (B). It was observed that the DTG
curve was similar to that of Cellulose-LDPE-Catalyst obtained in the same condition. More
importantly, the two peaks derived from the decomposition of DF and LDPE respectively
both shifted to a lower temperature when using the catalyst. It was evidenced that the catalyst
have a significantly influence in both the degradation of DF and LDPE, which could
dramatically reducing the initial decomposition temperature. As a result, the catalyst used
could be both beneficial to the decomposition of DF and LDPE.
406
Fig. 12.2 DTG curves of cellulose with LDPE (A); and Douglas fir sawdust (DF) with LDPE
(B).
Fig. 12.3 Variation of ∆
for biomass (cellulose and DF) and LDPE blends.
407
In order to determine the synergistic effect between biomass (cellulose/DF) and LDPE, we
defined the difference of weight loss (ΔW). Taking ΔW on the basis of the synergistic effect of
each material during pyrolysis. Fig. 12.3 shows the variation of ΔW with the temperature for
the different biomass/plastics blends. It can be seen that for the two blends, ΔW is less than
±1% before 300 °C. That is because at these temperature LDPE was still decomposed,
obviously there is no interaction between cellulose/DF and LDPE. However, ΔW is not equal
to zero in the two cases, which may be caused by the experiment errors such as the different
initial weight of samples and thermal conductive conditions. Both for the Cellulose-LDPE
and DF-LDPE, ΔW is more than 1% at the pyrolysis temperature higher than 500 °C, which
indicates that the synergistic effect during co-pyrolysis occurred mainly in the high
temperature. Furthermore, it was observed that ΔW of Cellulose-LDPE (~2.5%) is much
higher than that of DF-LDPE, suggesting that cellulose and LDPE have a more positive
synergistic effect than DF and LDPE. It is also noticed that ΔW of the two samples almost did
not change above 500 °C. In contrast, ΔW of the two samples first declined and then
increased sharply at 300 – 500 °C, peaking at 400 °C. These special pyrolysis behavior of the
blends can be explained by the fact that LDPE first soften at about 300 °C, and then produced
a plastic state which would inhibit the evolution of volatile matter. With further increasing
temperature, LDPE began to decompose quickly; therefore ΔW increased considerably
increased. Above 500 °C, the devolatilization processes of the blends have been essentially
completed. Hence ΔW is stable in this stage. On the other hand, woody biomass is more
stable than cellulose during pyrolysis, this is favorable for the fact that evolution of volatile
matter was remarkably affected by the soften LDPE at 300 – 500 °C.
408
Table 12.1 Representative results regarding kinetic parameters for the co-pyrolysis of cellulose
with LDPE.a
Mixture
Temp. (°C)
Conv. (%)
E (KJ/mol)
A (min-1)
R2
LDPE
400 – 520
2 – 100
213.78
3.95 × 1014
0.96
5 – 96
157.32
7.89 × 1012
0.96
Cellulose
300 – 400
Cellulose–LDPEb
300 – 500
3 – 92
134.59
5.37 × 107
0.96
Cellulose–LDPE–Catalyst
300 – 500
4 – 99
89.51
6.73 × 105
0.95
a Reaction
condition: plastics to biomass mass ratio, 1:4; heating rate, 20 °C/min.
12.4.2 Kinetic parameters
Linear model were determined by plotting ln[-ln(1-x)T2] versus 1/T for co-pyrolysis of
cellulose and LDPE as shown in Fig. 12.4; results suggested that the data were well fitted by
the gained linear models. The kinetic parameters including apparent activation energy (E) and
pre-exponential factor for LDPE, Cellulose, Cellulose-LDPE, and Cellulose-LDPE-Catalyst
during pyrolysis were determined by TGA as shown in Table 12.1. Since the R2 of the Fig.
12.4 (A), (B), (C), and (D) are all larger than 0.95, it was proved that the first-order reaction
mechanism
fits
well
for
LDPE,
Cellulose-LDPE,
and
Cellulose-LDPE-Catalyst
corresponding to Fig. 12.4 (A), (B), (C), and (D) respectively. It was obtained that apparent
activation energy (E) of LDPE and cellulose are 213.78 and 157.32 KJ/mol, respectively.
However, the apparent activation energy (E) of cellulose and LDPE mixture is only 134.59
KJ/mol, which is both less than that of LDPE and cellulose individually. This result indicates
that there was a positive synergistic effect between cellulose and LDPE on the lowering the
activation energy. In addition, with the presence of catalyst, the activation energy (E)
409
decreased to 89.51 KJ/mol comparing with activation energy (E) of the mixture without the
addition of catalyst. It was reaffirmed that the catalyst had a vital influence in the decrease of
activation energy (E).
-11
A
-12
ln (-ln (1-x)/T2)
-13
-14
y = -25713x + 20.644
R² = 0.963
-15
-16
-17
-18
0.00125
0.0013
0.00135
0.0014
1/T (K-1)
410
0.00145
0.0015
-11
B
ln (-ln (1-x)/T2)
-12
y = -18922x + 16.697
R² = 0.9602
-13
-14
-15
-16
0.00145 0.0015 0.00155 0.0016 0.00165 0.0017 0.00175 0.0018
1/T (K-1)
-13
C
ln (-ln (1-x)/T2)
-14
y = -16188x + 11.51
R² = 0.9608
-15
-16
-17
0.0015
0.00155
0.0016
0.00165
1/T (K-1)
411
0.0017
0.00175
0.0018
-11
D
ln (-ln (1-x)/T2)
-12
-13
y = -10766x + 2.2049
R² = 0.9458
-14
-15
-16
-17
0.0012
0.0013
0.0014
0.0015
1/T (K-1)
0.0016
0.0017
0.0018
Fig. 12.4 Plots of ln(-ln(1-x)/T2) vs 1/T of LDPE (A), Cellulose (B), Cellulose-LDPE (C),
Cellulose-LDPE-Catalyst calculated by using TGA.
As such, Fig. 12.5 shows the plots of ln[-ln(1-x)T2] versus 1/T for co-pyrolysis of DF and
LDPE; results indicated that the data were favorably fitted by the linear models as well. The
kinetic parameters including apparent activation energy (E) and pre-exponential factor for
LDPE, DF, DF-LDPE, and DF-LDPE-Catalyst during pyrolysis were determined by TGA as
shown in Table 12.2. Since the R2 of the Fig. 12.5 (A), (B), (C), and (D) are all larger than
0.95, it was evidenced that the first-order reaction mechanism fits well for LDPE, DF-LDPE,
and DF-LDPE-Catalyst corresponding to Fig. 12.4 (A), (B), (C), and (D) respectively. It was
also obtained that apparent activation energy (E) of LDPE and cellulose are 213.78 and 71.70
KJ/mol, respectively. However, the apparent activation energy (E) of cellulose and LDPE
mixture is only 89.63 KJ/mol, which is both less than that of LDPE and DF individually.
412
Additionally, with the presence of catalyst, the activation energy (E) decreased to 54.51
KJ/mol comparing with activation energy (E) of the mixture without the addition of catalyst.
It was reaffirmed that the catalyst had an important influence in the decrease of activation
energy (E).
Table 12.2 Representative results regarding kinetic parameters for the co-pyrolysis of Doulas fir
sawdust with LDPE.a
Mixture
Temp. (°C)
Conv. (%)
E (KJ/mol)
A (min-1)
R2
LDPE
400 – 520
2 – 100
213.78
3.95 × 1014
0.96
DF
300 – 400
15 – 89
71.70
2.79 × 105
0.98
DF–LDPE
300 – 500
4 – 99
89.63
1.86 × 106
0.95
DF–LDPE–Catalyst
300 – 450
11 – 93
54.51
5.19 × 103
0.99
a Reaction
condition: plastics to biomass mass ratio, 1:4; heating rate, 20 °C/min.
-11
A
-12
ln (-ln (1-x)/T2)
-13
-14
y = -25713x + 20.644
R² = 0.963
-15
-16
-17
-18
0.00125
0.0013
0.00135
0.0014
-1
1/T (K )
413
0.00145
0.0015
-12
B
ln (-ln (1-x)/T2)
-12.5
-13
y = -8624x + 0.4292
R² = 0.9822
-13.5
-14
-14.5
-15
0.00145 0.0015 0.00155 0.0016 0.00165 0.0017 0.00175 0.0018
1/T (K-1)
-13
C
ln (-ln (1-x)/T2)
-13.5
y = -6556x - 3.4157
R² = 0.9964
-14
-14.5
-15
0.00145 0.0015 0.00155 0.0016 0.00165 0.0017 0.00175 0.0018
1/T (K-1)
414
-11
D
ln (-ln (1-x)/T2)
-12
y = -10781x + 2.2258
R² = 0.9459
-13
-14
-15
-16
-17
0.0012
0.0013
0.0014
0.0015
1/T (K-1)
0.0016
0.0017
0.0018
Fig. 12.5 Plots of ln(-ln(1-x)/T2) vs 1/T of LDPE (A), DF (B), DF-LDPE (C),
DF-LDPE-Catalyst calculated by using TGA.
12. 5 Conclusions
Thermogravimetric analysis (TG) was used to study the thermal decomposition behavior and
kinetics of biomass (cellulose/DF) and plastics and catalytic co-pyrolysis of their mixture
over ZSM-5 catalysts. It was found that there was a positive synergistic effect between
biomass and plastics in the non-catalytic pyrolysis according to the ΔW at the end of the
experiment. In addition, the TG results showed that DTG curves slightly shifted to the left in
the presence of catalysts, which means adding catalysts tends to slightly lower the
temperature of thermal degrading process. The kinetic study of non-catalytic and catalytic
co-pyrolysis indicted that the first order well matched the non-catalytic and catalytic
co-pyrolysis of biomass with plastics. From the activation energy (E) and pre-exponential
factor (A) points of view, the addition of plastics in the non-catalytic pyrolysis of biomass
could appreciably decrease the activation energy (E). Moreover, the presence of ZSM-5
415
catalyst could further decrease the activation energy (E) based on the result of the blends
without catalyst. The kinetics analysis showed that adding catalyst doesn’t change the
decomposition mechanism. As a result, the kinetic study on catalytic co-pyrolysis of biomass
with plastics was implied that the catalytic co-pyrolysis is a promising technique that can
significantly reduce the energy input.
416
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418
CHAPTER THIRTEEN
GENERAL CONCLUSIONS AND FUTURE WORK SUGGESTIONS
13.1 General conclusions
The main purpose of this research was to investigate the influence of plastics in the catalytic
microwave-induced pyrolysis of biomass using zeolite-based catalysts on the product yield
and chemical selectivity. The reaction conditions were first optimized via catalytic
microwave pyrolysis of biomass using zeolite-based catalyst for product yield and selectivity.
The catalyst was then modified to improve the liquid selectivity from the catalytic microwave
pyrolysis of biomass. In the light of the aromatic-enriched bio-oil, the bio-oil was thereafter
hydrogenated into highly advanced biofuels by using Raney nickel-type catalysts. Based on
the proposed route from biomass to renewable cycloalkanes for jet fuels proposed, plastics
was added into the catalytic microwave pyrolysis of biomass to improve the carbon yield of
aromatics. It was found that with the plastics introduced in the catalytic microwave pyrolysis,
the overall carbon efficiency was strongly enhanced. Last but not least, the model compounds
(cellulose and low-density polyethylene) represented lignocellulosic biomass and waste
plastics were selected to determine the interactions between lignocellulosic biomass and
plastics during the catalytic microwave co-pyrolysis. The major conclusions form this
doctoral research are summarized below:
419
1. A novel pathway was investigated to produce gasoline-range aromatics and
hydrogen-enriched fuel gas by microwave-induced pyrolysis of cellulose integrated
with packed-bed catalysis in the presence of solid phase catalyst. The employed
catalyst was well-promoted ZSM-5 after the couplings of hydrothermal and calcined
treatments, completely converting volatile vapors derived from microwave pyrolysis
into aromatics and non-condensable gases. A central composite experimental design
(CCD) was employed to investigate the effects of catalytic temperature and inverse
weight hourly space velocity (WHSV)-1 on the pyrolysis-oils composition. It was
observed that the chemical compounds of the upgraded bio-oils from catalytic
microwave pyrolysis of cellulose were aromatic hydrocarbons, phenols, and aromatic
oxygenates. Aromatic hydrocarbons accounted for the largest selectivity of these
compounds were in the range from 82.93 to 96.60% in bio-oils depending on
alterations of catalytic conditions. Up to 48.56% selectivity towards aromatics in the
upgraded bio-oil belongs to gasoline-range aromatics at the mild condition. The
maximum selectivity of aromatic hydrocarbons (96.60%) was gained at packed-bed
temperature of 500 °C and WHSV-1 of 0.067 h. Gaseous results show that hydrogen
was the dominant composition, occupying approximately 40 vol.%. The high amounts
of gasoline-range aromatics and valuable hydrogen is attributed to the technologies of
microwave-assisted pyrolysis and ex-situ catalysis. These findings from this study
pave a new route for biorefinery industries to produce developed products (aromatics
and hydrogen-rich gas) through microwave-induced technologies.
420
2. The microwave-induced pyrolysis of low-density polyethylene (a model of
waste plastics) towards its conversion into biofuels was investigated using
ZSM-5 as a catalyst, generating significant amounts of gasoline-range
hydrocarbons. A central composite experimental design (CCD) was done to
investigate the effects of catalytic temperature and reactant to catalyst ratio on
the pyrolysis-oils composition and to achieve the maximum liquid yield. The
optimized condition for maximizing the yield of upgraded oil (32.58 wt. %)
was at 450°C and reactant to catalyst ratio of 2. GC-MS analysis showed that
mono-ring aromatic hydrocarbons were enriched and became the most
abundant compounds which varied from 74.73% to 88.49% in upgraded
pyrolysis-oils, depending on the catalytic pyrolysis conditions. Both low
temperature and high reactant to catalyst ratio gave rise to the formation of less
desirable polycyclic aromatic hydrocarbons whereas high temperature and high
ratio contributed to mono-ring aromatic hydrocarbons. The primary reaction
competing with aromatic hydrocarbon production was the formation of coke
which was negligible even at low catalytic temperatures. A plausible reaction
mechanism was also proposed in order to shed light on the overall catalytic
microwave pyrolysis of LDPE for aromatic hydrocarbons.
3. A novel pathway was investigated to produce jet fuel range cycloalkanes from intact
biomass. The consecutive processes for converting lignocellulosic biomass into jet
fuel range cycloalkanes principally involved the use of the well-promoted ZSM-5 in
421
the process of catalytic microwave-induced pyrolysis and Raney nickel catalysts in
the hydrogen saving process. Up to 24.68% carbon yield of desired C8 – C16 aromatics
was achieved from catalytic microwave pyrolysis at 500 °C. We observed that
solvents could assist in the hydrogenation reaction of naphthalene; and the optimum
result for maximizing the carbon selectivity (99.9%) of decalin was from the reaction
conducted in the n-heptane medium. The recovery of organics could reach ~94 wt. %
after the extracting process. These aromatics in the n-heptane medium were
eventually hydrogenated into jet fuel range cycloalkanes. Various factors were
employed to determine the optimal result under mild conditions. Increasing catalyst
loading, reaction temperature, and prolonged time could enhance the hydrogenation
reactions to improve the selectivity of jet fuel range cycloalkanes. Three types of
hydrogenation catalysts (NP Ni, Raney-Ni 4200, home-made Raney Ni) were chosen
to evaluate the catalytic performance. Results indicated that the home-made Raney
nickel was the optimal catalyst to obtain the highest selectivity (84.59%) towards jet
fuel range cycloalkanes. These cycloalkanes obtained can be directly used as additives
to make the desired jet fuels by blending with other hydrocarbons. Hence integrating
catalytic processes and conversion of lignocellulosic biomass pave a new avenue for
the development of green bio-jet fuels over inexpensive catalysts under mild
conditions.
4. Enhanced carbon yields of renewable alkanes for jet fuels were obtained through a
novel pathway from co-feeding of cellulose and low-density polyethylene (LDPE).
422
The consecutive processes proceeded via the catalytic microwave-induced pyrolysis
over well-promoted ZSM-5 catalyst and hydrogenation process by using home-made
Raney Ni catalyst. It was found that parent ZSM-5 modified by hydrothermal and
calcined treatments resulted in the increase of surface area as well as the formation of
more mesopores. Interestingly, the well-promoted ZSM-5 catalyst had high selectivity
toward C8 – C16 aromatic hydrocarbons in the co-feed catalytic microwave pyrolysis.
The raw organics with improved carbon yield (~ 44%) were more principally lumped
in the jet fuel range at the catalytic temperature of 375 °C with the LDPE to cellulose
ratio of 0.75. As the Ni-Al alloy dissolved by 20 wt.% NaOH solution, the BET
surface area, pore volume, and pore surface area of home-made Raney Ni catalyst
were appreciably improved. Based on XRD analysis, the diffractions regarding
metallic Ni were achieved as amorphous nature for home-made Raney Ni catalyst.
SEM analysis confirmed that the home-made Raney Ni catalyst was constituted by the
typical fractured and angular particles, and the small particles of home-made Raney
Ni catalyst were dispersive. The home-made Raney Ni catalyst was assayed for
hydrogenation of diverse organics species derived from co-feed catalytic microwave
pyrolysis under a low-severity condition. It was observed that the four species of raw
organics in the n-heptane medium were almost completely converted into saturated
hydrocarbons. The overall carbon yield (with regards to co-reactants of cellulose and
LDPE) of hydrogenated organics that mostly match jet fuels was sustainably
enhanced, reaching above 39%. Meanwhile, ~ 90% selectivity toward renewable
alkanes for jet fuels was attained. These enhanced hydrogenated organics with high
423
amounts of renewable cycloalkanes can be potentially used as high-density jet fuels or
additives for blending with commercial jet fuels.
5. Enhanced carbon yields of jet fuel range alkanes were manufactured from co-feeding
of lignocellulosic biomass with plastics. The consecutive processes proceeded via the
co-feed catalytic microwave-induced pyrolysis and hydrogenation process. In the
co-feed catalytic microwave pyrolysis by using ZSM-5 as the catalyst, parent ZSM-5
fabricated by hydrothermal and calcined treatments contributed to the increase of
surface area as well as the formation of more mesopores. Liquid organics with
enhanced carbon yield (40.54%) were more principally lumped in the jet fuel range
from the co-feed catalytic microwave pyrolysis performed at the catalytic temperature
of 375 °C with the plastics to biomass ratio of 0.75. To manufacture home-made
Raney Ni catalyst, the BET surface area, pore surface area, and pore volume of the
home-made Raney Ni catalyst were considerably improved when the Ni-Al alloy was
dissolved by the NaOH solution. In the hydrogenation process, we observed the three
species of raw organic derived from the co-feed catalytic microwave pyrolysis were
almost completely converted into saturated hydrocarbons under a low-severity
condition. The improved carbon yield (38.51%) of hydrogenated organics regarding
co-reactants of biomass and plastics predominantly match jet fuels. In the
hydrogenated organics, over 90% selectivity toward jet fuel range alkanes was
attained. In this respect, these hydrogenated organics with high amounts of renewable
424
cycloalkanes can be potentially served as high-density jet fuels or additives for
blending with civilian jet fuels.
13.2 Future research suggestions
According to the studies in the dissertation, several suggestions for future research may be
attempted and are summarized below:
1. Development of bimetallic catalysts for the ring opening of cycloalkanes to produce
aliphatic alkanes
Research question: will the developed catalysts significantly convert double-ring
cycloalkanes into aliphatic alkanes?
2. To
scale
up
our
proposed
catalytic
microwave-induced
co-pyrolysis,
a
techno-economic analysis from lignocellulosic biomass and waste plastics to jet fuels
should be conducted.
Research question: how to design a more efficient catalyst to minimize coke formation in the
catalytic co-pyrolysis of lignocellulosic biomass with plastics which is still a big challenge
for catalytic fast pyrolysis?
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