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Energy Conversion and Management 174 (2018) 489–503
Contents lists available at ScienceDirect
Energy Conversion and Management
journal homepage: www.elsevier.com/locate/enconman
Integration of calcium looping technology in combined cycle power plants
using co-gasification of torrefied biomass and coal blends
T
⁎
Po-Chih Kuoa,b, Jhao-Rong Chena, Wei Wua, , Jo-Shu Changa,c
a
Department of Chemical Engineering, National Cheng Kung University, Tainan 70101, Taiwan
Hawaii Natural Energy Institute, University of Hawaii at Manoa, Honolulu, HI 96822, USA
c
Research Center for Energy Technology and Strategy, National Cheng Kung University, Tainan 70101, Taiwan
b
A R T I C LE I N FO
A B S T R A C T
Keywords:
CaL technology
SE-WGS reaction
Kinetic modelling
Torrefaction
Co-gasification
To comprehensively understand the impact of calcium looping (CaL) technology on a co-gasification of torrefied
biomass and coal power plant, two kinds of proposed configurations (pre- and post-CaL), including a co-gasification system, a CaL system, and a combined heat and power system, are studied and compared with each other
in the present work. In the CaL process, the kinetic modeling of the carbonation reaction and CaO sorptionenhanced water gas shift (SE-WGS) reaction are developed in a fast fluidized bed reactor (carbonator) to predict
the performance of hydrogen production and CO2 capture. The influences of torrefied biomass blending ratios
(BRs) and CaO to fuel mass flow rate ratios (CaO/F) on various performance indicators such as hydrogen enhancement factor, hydrogen thermal efficiency (HTE), CO2 capture efficiency, specific CO2 emissions, and
overall system efficiency are evaluated. A comparison of pre- and post-CaL schemes reveals that the SE-WGS
reaction has a markedly profound effect on the former, causing the hydrogen production and HTE to be higher
than those in the latter, whereas the latter is much more conducive to CO2 capture and specific CO2 emissions.
Under optimal operating conditions (BR = 40 wt%, CaO/F = 3.5), the values of CO2 capture efficiency and
overall system efficiency of both schemes are higher than 90% and 50%, respectively. Overall, the pre-CaL case
is suitable to design as a highly efficient co-generation of hydrogen production and electricity plant with low CO2
emissions, whereas the post-CaL case is recommended for a co-gasification power plant with nearly zero CO2
emissions.
1. Introduction
CaO sorption-enhanced water gas shift reaction:
Carbon capture technologies are currently being widely developed
to reduce CO2 emissions from power plants and industrial sectors. In
recent years, several carbon capture approaches have become available
to capture CO2 from flue gases. In general, they can be classified into
pre-combustion, post-combustion, and oxy-fuel combustion capture
types [1]. In addition to the above conventional methods, various
chemical looping technologies are now being investigated for integration into power plants, one of which is the calcium looping (CaL)
carbon capture process, which utilizes solid CaO particles to carry out
sorption and desorption cycles for the purpose of capturing and concentrating CO2. A schematic illustration of the CaL process is shown in
Fig. 1, which contains a carbonator and a calciner. The major reactions
occurring in the CaL process are given as follows [2,3]:
Carbonation reaction:
CaO(s) + CO2(g) ↔ CaCO3(s) ΔH∘298 = −178 kJ mol −1
⁎
(R1)
CO(g) + H2 O(g) ↔ CO2(g) + H2(g) ΔH∘298 = −42.4 kJ mol −1
(R2)
Calcination reaction:
CaCO3(s) ↔ CaO(s) + CO2(g) ΔH∘298 = 178 kJ mol −1
(R3)
In the carbonator, emitted CO2 in an exhaust gas produced from a
power plant is reacted with solid CaO sorbents to form CaCO3 through
an exothermic carbonation reaction (R1) at around 650 °C. Furthermore, if the flue gas contains syngas (H2 and CO), H2O, and CO2, according to Le Chatelier’s principle, a water gas shift reaction (R2) also
occurs due to the removal of CO2 from carbonation, which is the socalled CaO sorption-enhanced water gas shift reaction (SE-WGS).
Therefore, syngas can be converted into a hydrogen-rich product gas
through the CaO SE-WGS reaction with simultaneous in situ CO2 removal. In the calciner, the endothermic calcination (R3) of CaCO3 takes
place at around 900 °C to decompose concentrated CO2 and regenerate
Corresponding author.
E-mail address: weiwu@mail.ncku.edu.tw (W. Wu).
https://doi.org/10.1016/j.enconman.2018.08.044
Received 16 June 2018; Received in revised form 8 August 2018; Accepted 11 August 2018
0196-8904/ © 2018 Elsevier Ltd. All rights reserved.
Energy Conversion and Management 174 (2018) 489–503
P.-C. Kuo et al.
Nomenclature
A
Ar
CA
Cb, i
Cc, i
Ce, i
umf
ϕ
dp
dp*
h
g
ρs
ρg
εmf
εb
u
μ
uT
kelu
kcat
kb,i
k c, i
k e, i
Rep
νf
fb
fc
fd
fe
fl
fe
fcore
fwall
Wd
Wl
WT
Hd
Hl
HT
Rb, i
R c, i
R e, i
cross sectional area of fluidized bed (cm2)
Archimedes number (–)
concentration of reactant gas (mol L−1)
concentration of reactant i in the bubble phase (mol L−1)
concentration of reactant i in the cloud phase (mol L−1)
concentration of reactant i in the emulsion phase (mol
L−1)
minimum fluidize velocity (cm s−1)
particle sphericity
particle diameter (cm)
dimensionless particle dimeter
height of fluidized bed (cm)
gravity (980 cm s−2)
density of particle (kg m−3)
density of reactant gas (kg m−3)
porosity of the fixed bed at minimum fluidization
local bubble volume fraction
gas velocity (cm s−1)
viscosity of reactant gas (kg m−1s−1)
terminal velocity (cm s−1)
elutriation velocity (cm s−1)
intrinsic rate constant (s−1)
rate constant of reactant i in bubble phase (s−1)
rate constant of reactant i in cloud phase (s−1)
rate constant of reactant i in emulsion phase (s−1)
Kbc
K ce
ηd
ηl
Reynolds number (–)
kinematic viscosity of the gas (cm2 s−1)
volume fraction of solids in bubble phase
volume fraction of solids in cloud phase
volume fraction of solids in dense region
volume fraction of solids in emulsion phase
volume fraction of solids in lean region
volume fraction of solids for the exit location
volume fraction of solids in the core region
volume fraction of solids sliding down the wall
weight of catalyst in dense region (kg)
weight of catalyst in lean region (kg)
total weight of catalyst (kg)
height of dense region (cm)
height of lean region (cm)
total height of fluidized bed (cm)
reaction rate of reactant i in bubble phase (mol kg−1 s−1)
reaction rate of reactant i in cloud phase (mol kg−1 s−1)
reaction rate of reactant i in emulsion phase (mol kg−1
s−1)
volume rate of gas exchange coefficients between bubble
and cloud phases per unit bubble volume (s−1)
volume rate of gas exchange coefficients between cloud
and emulsion phases per unit bubble volume (s−1)
contact efficiency in the dense region
contact efficiency in the lean region
Fig. 1. Schematic illursation of the calcium looping technology.
example, Cormos and Cormos [6] evaluated a coal integrated gasification combined cycle (IGCC) power plant with the CaL process in both
pre- and post-combustion configurations. They concluded that the net
efficiency of their designed systems was in the range of 34–37% with a
carbon capture rate greater than 95%, and the pre-combustion configuration had higher energy plenty when compared to that of the postcombustion one. Zhu et al. [7] simulated a coal IGCC system with the
CaL process using the Aspen Plus simulator and pointed out that higher
energy and exergy efficiencies are obtained in the CaL based IGCC
system compared to those of a physical absorption-based system. In
addition, they also found that the CaL based system is more economical
than the physical absorption-based technology. Hu and Ahn [8]
CaO particles. Based on the reversible carbonation and calcination cycles of the CaL process, its advantages include the following [2–5]: (1) a
high concentration of CO2 is obtained without using solvents, and there
are lower energy and cost penalties when compared to MEA- or MDEAbased process; (2) the natural sorbents used for CaL process are low cost
and widely available; (3) the excess waste heat produced from the
carbonator can be recovered for electricity generation, leading to
higher system energy efficiency, and (4) the waste spent from the CaL
process can be reutilized and integrated into the cement industry.
Reviewing recent studies concerning the applications of CaL technology, many researchers have studied the process integration of the
CaL process in a coal gasification combined cycle power plant. For
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Energy Conversion and Management 174 (2018) 489–503
P.-C. Kuo et al.
highest system thermal efficiency value, but hydrogen generation was
slightly less than the iron based chemical looping system. Mehrpooya
et al. [13] designed a co-generation of hydrogen and electricity plant by
combining biomass gasification, a CaL process, and a power generation
cycle with CO2 capture. They pointed out that the overall system efficiency was 57.3% with zero CO2 emissions, and thus this new integrated system could replace conventional gasification systems.
Shahbaz et al. [14] conducted a steady state Aspen Plus simulation to
investigate biomass steam gasification with CaL CO2 capture and found
that the presence of CaO sorbents can enhance hydrogen formation.
From these earlier studies, it can be concluded that biomass steam
gasification incorporating CaL technology not only can substantially
reduce CO2 emissions, but also can enable the production of a hydrogen-rich product gas as a result of the SE-WGS reaction [15,16].
Despite the numerous advantages of integrating biomass gasification with CaL technology, raw biomass is generally characterized by a
relatively low calorific value and a high moisture content, which results
in low energy efficiency [5,17,18]. Furthermore, significantly more
energy is required for grinding raw biomass into small particles due to
its fibrous nature as compared to coal, resulting in higher costs for
transport and storage. Therefore, to this end, a few studies focusing on
investigated a natural gas IGCC power plant associated with the CaL
process. They found that the net efficiency of the designed plant could
be improved by introducing exhaust gas recirculation. Based on the
foregoing review of the literature, it is evident that CaL technology is
playing a vital role in applications to coal gasification power plants and
other industrial processes.
Another prospective possibility and attractive approach to reducing
CO2 emissions and simultaneously providing economical and efficient
use of energy resources is the development of biomass energy.
Considerable efforts have been made in developing biomass thermal
conversion technologies, including pyrolysis, torrefaction, gasification,
and combustion, to convert raw materials into high-quality biofuels,
biogas, heat, and electricity [9]. Among them, biomass gasification is
gaining significant attention due to the higher flexibility of fuels and
greater energy efficiency associated with a combined cycle power plant
[10,11]. Over the past several years, a number of studies have been
conducted on integrated biomass gasification with the CaL technology.
For example, Yan et al. [12] proposed a biomass gasification system
with a CaL process in Aspen Plus and compared it with two iron based
chemical looping systems in terms of hydrogen generation and system
thermal efficiency. They indicated that the CaL based system had the
Fig. 2. Block diagram of the (a) pre-calcium looping process and (b) post-calcium looping.
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Energy Conversion and Management 174 (2018) 489–503
P.-C. Kuo et al.
Syngas flow rate (kmol h-1)
performance indicators. The optimal operating conditions of each
configuration will also be outlined. Finally, two optimal designs for
both pre-and post-CaL cases are presented, which can provide useful
insight into the development of a process for co-gasification of torrefied
biomass and coal using CaL technology.
H2
CO
Syngas
120
100
2. Process design and simulation
80
2.1. System description
60
Two configurations of a co-gasification power plant associated with
a calcium looping carbon capture are designed in Aspen Plus V8.8 and
Matlab 2015. The Peng-Robinson Boston Mathias (PR-BM) equation of
state was used as the thermodynamic property. As a whole, the entire
system consists of three major units, a co-gasification unit, a calcium
looping (CaL) unit, and a heat recovery steam generator (HRSG) unit.
The first designed process is called a pre-calcium looping (pre-CaL)
power plant, whereas the second one is called a post-calcium (post-CaL)
looping power plant. The key difference lies in the position of the CaL
unit. In the pre-CaL scheme, the CaL unit is installed behind the cogasification system, while the CaL unit is installed downstream of the
syngas combustion system in the post-CaL scheme. The overall process
flow diagram for both schemes is shown in Fig. 2. Crucial sub-systems
such as the co-gasification unit and the CaL unit are described in detail
below, whereas the HRSG system process is explained in a previous
study [17].
40
20
0
BR=0%
S/C=1
BR=20%
S/C=0.9
BR=40% BR=60% BR=80%
S/C=0.78 S/C=0.68 S/C=0.55
BR=100%
S/C=0.45
Fig. 3. The syngas formation from the co-gasification system under optimal
operating conditions [5].
blending of biomass and coal combined with a CaL process have been
published. Yan and He [19] indicated that the biomass mixing ratio has
an obvious impact on the carbon sequestration rate in a combined cycle
power plant. Above all, negative net carbon emissions of the system was
observed when the biomass mixing ratio was greater than 0.6. More
recently, Schakel et al. [20] analyzed the impact of adopting various
fuels such as coal, natural gas, biomass, and coal/biomass blends as
feedstocks on the environmental performance of the post-combustion
CaL process in a cement plant. They reported that the CaL process
performance could be enhanced by replacing coal with biomass and a
negative life cycle for CO2 emissions was presented.
In contrast to the use of coal/biomass blends, torrefied biomass or
coal/torrefied biomass has recently been extensively used as a feedstock
in the gasification process to evaluate its potential for replacing traditional coal gasification. Torrefied biomass is produced from a pretreatment process called torrefaction, which is a mild pyrolysis carried
out at temperatures ranging from 200 to 300 °C [5,21]. This pretreatment can improve the aforementioned intrinsic drawbacks of raw biomass [5,17,18]. In other words, using torrefied biomass as a fuel has
potential benefits for gasification [15,21,22], combustion [23,24], cofiring [18,25–27], and co-gasification [5,17,28]. However, an examination of the recent literature indicates that no research has been
conducted to explore an integrated system incorporating co-gasification
of torrefied biomass and coal with CaL capture. In addition, most of the
aforementioned studies adopted a thermodynamic analysis to evaluate
the performance of the carbonation reaction (R1) and even neglected
the importance of the CaO SE-WGS reaction (R2). For these reasons, the
present work proposes two integrated co-gasification of torrefied biomass and coal systems with the CaL process, called pre-CaL and postCaL, which are compared with each other. In particular, kinetic models
of the carbonation and CaO SE-WGS reactions are developed in a fast
fluidized bed reactor (carbonator) to investigate the effect of CaO sorbents on the hydrogen production performance and CO2 capture efficiency in detail. Emphasis is particularly placed on the impact of various operating conditions, including torrefied biomass blending ratios
(BRs) and CaO to fuel mass flow rate ratios (CaO/F), on the
2.2. Co-gasification process
In this study, co-gasification of torrefied biomass and coal using
steam as the gasifying agent is conducted for the purpose of producing
syngas. The physical and chemical properties of the torrefied biomass
and coal. Regarding the biomass steam gasification process, three operation units, including a drying unit, a decomposition unit, and a gasifier, are taken into account. The detailed gasification process are
presented in our previous study [5,17,29,30]. In addition, the validation of developed gasification model has been carried out in the studies
of Kuo et al., [29] and Kuo and Wu, [30], where the predicted results
were in good agreement with the experimental results.
Two significant variables, the steam-to-fuel mass flow rate ratios (S/
F) and the torrefied biomass blending ratios (BRs), are taken into account to investigate the performance of the co-gasification system, as
follows:
S /F =
ṁ steam
ṁ fuel
BR (wt%) =
(1)
ṁ torrefied biomass
ṁ torrefied biomass + ṁ coal
× 100%
(2)
where ṁ steam and ṁ fuel indicate the mass flow rate of the steam (gasifying agent) and blending fuels (kgh−1), respectively, and ṁ torrefied biomass
and ṁ coal represent the mass flow rate of input the torrefied biomass and
coal.
In this work, six BRs are considered: 0%, 20%, 40%, 60%, 80%, and
100%. Therefore, the higher the BR is, the more torrefied biomass is
blended in the fuel. Furthermore, according to our previous study [5],
where a constrained optimization algorithm was conducted to maximize the syngas yield in a co-gasification process subject to the
minimum energy requirements for its system, it is shown that the optimal S/F ratio for each BR is located at the carbon boundary point
(CBP). Based on the previous optimization algorithm, Fig. 3 shows the
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Energy Conversion and Management 174 (2018) 489–503
P.-C. Kuo et al.
Table 1
The reactor specifications and operating conditions for the input parameters
used in validation of the present model [31].
Unit specification/
Inlet composition/
Performance
indicator
Carbonator
Height (cm)
Diameter (cm)
Temperature (°C)
Superficial
velocity (m s−1)
Inventory weight
(kg m2)
Value
(experimental
operating
conditions)
Value (input
parameters in
simulation)
650
10
568–722
3
650
10
650
3
33–357
178
20
80
20
80
Experimental data
Simulation data
∼4
4.86
> 70%
78.98
Inlet gas composition (mol%)
CO2
N2
Comparisons between simulation and
experimental results
Capture rate (mol
m−2 s−1)
CO2 capture
efficiency (%)
Fig. 4. Schematic diagram of the carbonator and the calciner.
optimal syngas production values for different BRs. Consequently, the
following investigation of two co-gasification power plants schemes for
six BRs from 0 to 100 wt% is conducted at those optimal S/F ratios (i.e.
1.0, 0.9, 0.78, 0.68, 0.55, and 0.45).
CO2 partial pressure or driving force , the reaction order is first order,
whereas the reaction order is approximately zero at higher CO2 partial
pressures. Consequently, the reaction rate (ks ) can be expressed as
follows based on the CO2 partial pressure driving force [32]:
For
2.3. Calcium looping process
The calcium looping carbon capture unit, which includes a carbonator and a calciner, is another important part of the proposed co-gasification power plant. First of all, the mathematical model of the carbonator is built in Matlab 2015a to simulate a gas-solid (CaO-CO2)
reaction, namely carbonation reaction. In this work, the carbonator is
modelled as a fast fluidized bed reactor under steady-state operation.
The design equations for the fast fluidized bed reactor are provided in
the Appendix A. Meanwhile, the specifications of the carbonator are
based on the experimental system established by Rodriguez et al. [31].
In their experiments, the height of the carbonator was 650 cm with an
internal diameter of 10 cm. Fig. 4 shows a schematic diagram of the
dual fluidized bed CaL process. Initially, a sepcific amount of the bed
inventory (CaO) is packed inside the carbonator. The syngas (pre-CaL
case) or exhaust gas (post-CaL case) is then fed into the bottom of the
carbonator and reacted with the regenerated CaO from the calciner.
Therefore, the CO2 contained in the syngas or exhaust gas reacts with
the active CaO in the carbonator through the carbonation reaction (R1).
For simplicity, the calcination reaction (R3) is modeled in Aspen Plus as
an RGibbs reactor (calciner), in which the chemical and phase equilibrium calculations are performed by minimizing the Gibbs free energy.
(PCO −PCO ,eq) > 10 kPa,
2
(
)
ks
−E
⎞ (mol m−2 s−1),
= 1.67 × 10−3exp ⎛
⎝ RT ⎠
E = 29 kJ mol−1
(4)
100
CO2 capture efficiency (%)
80
2.3.1. The kinetics of carbonation reaction
The intrinsic kinetic reaction of carbonation reaction (i.e. CaO-CO2
reaction) is established according to Sun et al. [32]. In general, a CaOCO2 reaction model can be expressed as follows:
dX
R =
= MCaO ks PCO2−PCO2, eq n S
dt (1−X )
n = 0,
2
60
40
20
Present model
Experimental data [31]
(3)
−1
0
where R is the specific reaction rate (s ); X is the conversion of CaO;
MCaO is the molar weight of CaO (56 g mol−1); ks is the rate constant;
PCO2 and PCO2, eq are the partial pressure of CO2 (kPa) and equilibrium
partial pressure of CO2 (kPa), respectively; n is the reaction order, and S
is the specific surface area (29 m2 g−1)
Notably, the reaction rate (ks ) and reaction order (n ) are affected by
the CO2 partial pressure driving force, namely (PCO2−PCO2, eq) . At a low
0
50
100
150
200
250
300
350
400
Inventory weight (kg m-2)
Fig. 5. A comparison of the CO2 capture efficiency between experimental and
simulated results [31].
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Energy Conversion and Management 174 (2018) 489–503
P.-C. Kuo et al.
(PCO −PCO ,eq) ⩽ 10 kPa,
2
2
n = 1,
ks
−E
⎞ (molm−2s−1 (kPa)−1),
= 1.67 × 10−4exp ⎛
⎝ RT ⎠
= 29 kJ mol−1
E
(5)
100
It is thus obvious that a critical value of the CO2 partial pressure
driving force is the most influential factor for the carbonation reaction.
In addition, the equilibrium partial pressure of CO2 (PCO2, eq ) is influenced by the reaction temperature, which can be described by the
following relation [33,34]:
8307.83
T (k )
80
CO2 capture efficiency (%)
log PCO2, eq (kPa) = 9.079−
(6)
In order to calculate the CO2 molar concentration, Eq. (3) can be
further rearranged to the following equation on a concentration basis:
dCCO2
= MCaO ks RTρbed CCO2−CCO2, eq n S (1−X )
dt
(
)
90.00
80.00
70.00
60.00
50.00
40.00
30.00
20.00
10.00
90
70
60
50
40
30
20
800
10
(7)
0
400
where CCO2 and CCO2, eq are the concentration of CO2 and equilibrium
concentration of CO2, respectively; R is the ideal gas constant (8.314 L
kPa mol−1 K−1), and ρbed is the bulk density (kg m−3).
The variations in the CO2 concentration as a function of the carbonator height can thus be obtained by solving a set of ordinary differential equations (ODEs) (see Appendix A).
T em
600
pera
ture o
(C
)
-2
m
(kg
400
ht
g
i
e
200
yw
tor
n
e
Inv
600
800 0
)
Fig. 6. Distributions of CO2 capture efficiency as a function of reactor temperature and inventory weight.
2.3.3. The kinetics of the CaO sorption-enhanced water gas shift (SE-WGS)
reaction
In addition to developing a kinetic model of the carbonation reaction, the CaO sorption-enhanced water gas shift (SE-WGS) reaction is
also considered in the carbonator. However, due to the lack of kinetic
data for the CaO SE-WGS reaction in the literature, an approach using a
nonlinear regression of the WGS reaction is carried out to estimate the
kinetic parameters (activation energy, pre-exponential factor, and reaction order) of the CaO SE-WGS reaction. A kinetic expression for the
general WGS reaction on a Ni/MgAl2O4 catalyst has been considered as
the basis of the present model, as given by Xu and Froment [35]:
2.3.2. Model validation and performance
The carbonator model validation has been carried out by considering the experimental study from Rodríguez et al. [31]. Table 1
shows the reactor specifications, operating conditions, feed gas composition, and a comparison between the simulation results and available experimental data. It is clearly shown that the experimental results
are in good agreement with the predicted results. Furthermore, Fig. 5
compares the CO2 capture efficiency as a function of inventory weight
between the simulated and experimental results. It can be also seen that
the experimental data fit the predicted data very well. Thus, it is verified that the proposed model is appropriate to evaluate the performance of the actual carbonator.
In order to obtain the optimal operating conditions for the carbonator, two manipulated variables, operating temperature and inventory
weight, are taken into account in optimizing the CO2 capture efficiency.
The reaction temperature of carbonator is in the range of 400–775 °C,
while the inventory weight is controlled between 25 and 725 kg m−2.
Fig. 6 shows the three-dimensional profiles of CO2 capture efficiency as
a function of temperature and inventory weight. It can be seen that the
CO2 capture efficiency is in the range of 4.6–96.9% and is sensitive to
changes in inventory weight and reaction temperature, particularly to
the latter. Moreover, the values of CO2 capture efficiency are higher
than 90% when the operational temperature range is between 600 and
650 °C and the inventory weight is greater than 375 kg m−2. Notably,
CO2 capture efficiency descends rapidly as the reaction temperature
increases or decreases from a reaction temperature of 650 °C. This is
due to the fact that the carbonation reaction rate significantly decreases
when the operating temperature is below 550 °C, whereas the thermodynamic limitation causes lower CO2 capture efficiency when the operating temperature is greater than 700 °C. It is thus determined that
the optimal reaction temperature is 650 °C, which is consistent with
other studies [4], while the optimum inventory weight should be controlled to be greater than 375 kg m−2 in order to maintain a higher
carbonation reaction rate.
rWGS =
1
kWGS
CCO CH2 O−CH2 CCO2
(DEN )2 CH2
(
)
DEN = 1 + K CO CCO + K H2 CH2 + K CH4 CCH4 +
(8)
K H2 O CH2 O
CH2
(9)
The rate constants can be expressed by:
kWGS = 7.558 exp ⎡−
⎣
67310 1
1 ⎤
⎛ −
⎞ ·(1−0. 2X )(kmol kg−1h−1bar −1)
R ⎝ T 648 ⎠ ⎦
(10)
Muller et al. [36] established an experimental system to observe the
CaO SE-WGS reaction behavior by considering a gas mixture containing
7 mol% CO, 30 mol% H2O, and 63 mol% N2 at the temperature of
600 °C. Aa a result, based on the experimental data reported by Muller
et al. [36], Eq. (10) can be further modified by comparing it with the
experimental data via a nonlinear regression analysis. In this way, the
kinetic rate constant of CaO SE-WGS reaction can be estimated as follows:
kSE
- WGS
= 1. 2841 exp ⎡−
⎣
67310 1
1 ⎤
⎛ −
⎞ ·(1−0.2X )(kmol kg−1h−1bar −1)
R ⎝ T 648 ⎠ ⎦
where X is the conversion of CaO.
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Energy Conversion and Management 174 (2018) 489–503
P.-C. Kuo et al.
0.1
calcium looping technology, namely, pre-calcium looping (pre-CaL) and
post-calcium looping (post-CaL), are evaluated first. The biomass
blending ratio (BR) is taken into account to determine the optimal BR of
the two proposed schemes. Subsequently, an optimal design of each
configuration is recommended in terms of overall system energy efficiency, carbon capture efficiency, CO2 emissions, and so forth.
Mole fraction (%)
0.08
Simulation CO
Simulation H 2
Simulation CO2
Experiment CO [36]
Experiment H 2 [36]
Experiment CO2 [36]
0.06
0.04
3.1. Comparisons of the calcium looping process’ performance
As discussed above, the CaL unit is installed behind the co-gasification system in the pre-CaL case (Fig. 2a), while the CaL unit is placed
at the downstream of a combustion system in the post-CaL case
(Fig. 2b). Initially, in order to compare the performance of the CaL unit
for the two configurations, it is important to define the system
boundaries for two cases. The system boundary considered in the preCaL case is the co-gasification/CaL system, whereas it covers the cogasification/combustion/CaL in the post-CaL case. The detailed operating conditions for each sub-system are presented in Table 2. Fig. 8
presents the hydrogen production rate at the carbonator exit (Fig. 8a)
and the performance of the carbonator, including the hydrogen formation enhancement (Fig. 8b), hydrogen thermal efficiency (Fig. 8c),
CO2 capture efficiency (Fig. 8d), and the energy consumption of the
calciner (Fig. 8e) at various BRs (0–100 wt%) for the two schemes.
Meanwhile, the performance indexes are obtained by
0.02
0
0
200
400
600
800
1000
1200
Time (s)
Fig. 7. Plots for experimental data and fitted curves of the CaO sorption-enhanced WGS reaction [36].
Table 2
Design assumptions and operating conditions used in each sub-system
[5,17,31].
Process sub-systems
Parameters
Value
Co-gasification system
Temperature (°C)
Pressure (atm)
Fuel inlet flow rate (kg h−1)
BR ratio (–)
S/F ratio (–)
Carbon conversion (%)
900
1
1000
0–100
0.45–1
∼99
Calcium looping system
Carbonator
Hydrogen enhancement factor (%) =
CO2 capture efficiency (%) =
XCO (%) =
650
1
650
138.59–160.4
3
742.32–994.25
3
Calciner
Temperature(°C)
Pressure (atm)
875
1
HRSG system
HPT outlet pressure (atm)
IPT outlet pressure (atm)
LPT outlet pressure (atm)
Isentropic efficiency (%)
Pump efficiency (%)
Approach point and pinch point (°C)
110
80
40
72
85
10
FHin2
× 100%
in
FCO
2
in
+ XCO FCO
−FCout
O2
in
in
FCO2 + XCO FCO
(12)
× 100%
(13)
in
out
FCO
−FCO
in
FCO
(14)
Hydrogen thermal efficiency (% )=
Temperature (°C)
Pressure (atm)
Height (cm)
Diameter (cm)
Superficial velocity (m s−1)
Inventory weight (kg m2)
CaO/F ratio (–)
FHout
−FHin2
2
ṁ H2 LHVH2
̇
̇
ṁ fuel LHVfuel + QCOG
+ QCalciner
× 100%
(15)
and
the mole flow rate of hywhere
drogen, CO2, and CO at the inlet and outlet carbonator (kmol h−1),
respectively; XCO is the conversion of the WGS reaction, which is defined by Eq. (14); ṁ H2 is the mass flow rate of hydrogen at the outlet
carbonator (kg h−1); LHVH2 and LHVfuel are the lower heating values of
̇
̇
the hydrogen and fuel (MJ kg−1), respectively, and QCOG
and QCalciner
represent the required heat rate of the co-gasification system and the
calciner (kW), respectively.
As shown in Fig. 8a, for the pre-CaL case, the hydrogen production
rate is significantly affected by the BR within in the ranges investigated
in this work and is in the range of 89.93–115.18 kmol h−1. A higher BR
means a lower amount of coal blended in the feedstock, resulting in less
carbon content in the feedstock for the co-gasification system. In contrast, for the post-CaL case, the hydrogen production rate is insensitive
to the BR and is much lower than the pre-CaL case for all BRs. This
arises from the fact that syngas produced from the co-gasification is first
combusted, resulting in the generation of less hydrogen in the CaL unit.
Next, in regard to examining the influence of the CaO sorption-enhanced WGS reaction on hydrogen production, Fig. 8b reveals that for
the pre-CaL case, the hydrogen formation enhancement is greater than
65%, regardless of what BR is tested, implying a substantial increase in
hydrogen production after passing through the CaL unit. Specifically, a
higher BR leads to better enhancement. However, similar to the postCaL case in Fig. 8a, almost all of the syngas is consumed in the combustor, and thus the hydrogen enhancement is not taken into account
FHin2 ,
The profile of the curve fitting results is displayed in Fig. 7, where
the modified reaction rate expression (Eq. (11) completely fits the experimental data [36] regardless of what the gas component is, meaning
that the modified kinetic equation and parameters can reasonably
predict the CaO SE-WGS reaction in the carbonator.
3. Results and discussion
In this work, two designed configurations of the co-gasification of
torrefied biomass and coal power generation associated with the
495
FHout
,
2
in
FCO
,
2
FCout
O2 ,
in
FCO
,
out
FCO
are
Energy Conversion and Management 174 (2018) 489–503
P.-C. Kuo et al.
(b)
120
(a)
100
Pre-CaL
Hydrogen enhancement factor (%)
Pre-CaL
Post-CaL
H 2 flow rate (kmol h-1)
100
80
60
40
20
0
0%
20%
40%
60%
80%
80
60
40
20
0
100%
0%
20%
BR(%)
(c)
(d)
80%
100%
120
Pre-CaL
Post-CaL
100
80
CO2 capture efficiency (%)
Hydrogen thermal efficiency (%)
Pre-CaL
Post-CaL
60
40
20
80
60
40
20
0%
20%
40%
60%
80%
0
100%
0%
20%
40%
60%
80%
100%
BR(%)
BR(%)
(e)
60%
BR(%)
100
0
40%
3500
Pre-CaL
Post-CaL
Energy requirements (kW)
3000
2500
2000
1500
1000
500
0
0%
20%
40%
60%
80%
100%
BR(%)
Fig. 8. The performance of the calcium looping unit for pre- and post-CaL cases: (a) H2 production rate, (b) enhancement factor of H2 production, (c) hydrogen
thermal efficiency, (d) CO2 capture efficiency, and (e) energy requirements of the calciner.
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Table 3
The performance of each sub-system for pre- and post-CaL cases.
BR (wt%)
0
20
40
60
80
100
S/F
1
0.9
0.78
0.68
0.55
0.45
Co-gasification system
Thermal energy of fuel (kW)
Co-gasification duty (kW)
7450
4251.82
7193.89
3845.07
6937.78
3415.64
6681.67
3008.94
6425.56
2564.27
6169.44
2156.84
Calcium looping system
Calciner duty (kW)
Pre-CaL/post-CaL
2674.50/2912.57
2658.82/2873.47
2620.57/2828.94
2580.04/2777.56
2534.94/2720.19
2489.28/2658.82
HRSG system
Net power output (kW)
Pre-CaL/post-CaL
7163.62/7533.87
6923.66/7260.13
6563.74/6822.92
6203.86/6399.82
5963.97/6113.10
5484.23/5689.49
60
here. As a consequence, the CaO sorption-enhanced water gas shift (SEWGS) reaction plays an important role in the pre-CaL case, whereas it is
less applicable to the post-CaL case. As far as the hydrogen thermal
efficiency is concerned (Fig. 8c), it is worth noting that the HTE values
for the pre- and post-CaL cases vary from 63 to 65.5% and from 9 to
11.1%, respectively, when the BR ratio is raised from 0 to 100 wt%.
Obviously, blending the torrefied biomass with coal can slightly amplify
system energy efficiency. The profiles of the CO2 capture efficiency
from the CaL unit for the two schemes at various BRs is displayed in
Fig. 8d. As a whole, the CO2 capture efficiencies for the post-CaL case
are in the range of 93–97%, which is greater than that of the pre-CaL
case (86–91%). Since combustion system is installed in front of the CaL
unit, (Fig. 2), this observation can be attributed to the fact that the
concentration of CO2 for the post-CaL case is greater than that for the
pre-CaL case. Furthermore, it is also noteworthy that the CO2 capture
efficiency can be improved with increasing BR, regardless of which
configuration is tested. With attention paid to the energy demand for
regeneration of the CaO sorbent from the calciner, Fig. 8e shows that
the energy consumption of the calciner decreases from 2694.5 to
2489.28 and from 2912.57 to 2658.82 kw for pre- and post-CaL cases,
respectively, as the BR ratio rises from 0 to 100 wt%. Apparently, the
post-CaL case requires more energy for the calcination as a result of the
increased CO2 produced from the combustor as compared to the preCaL case.
Pre-CaL
Post-CaL
Overall system efficiency (%)
50
40
30
20
10
0
0%
20%
40%
60%
80%
100%
BR(%)
(a)
45
Pre-CaL
Post-CaL
Specific CO2 emissions (g/kW h)
40
3.2. Comparisons of the co-gasification power plant performance
35
30
Subsequently, the variations in the overall system efficiency and
specific CO2 emissions (ECO2 ) with respect to the BR for both configurations are compared with each other. Each index is defined by
25
Overall system efficiency (%) =
Wnet
× 100%
̇
̇
+ QCalciner
ṁ fuel LHVfuel + QCOG
(16)
20
Wnet (kW) = WHPT + WIPT + WLPT −WPump
15
Specific CO2 emissions (kg/kWh) =
10
0%
20%
40%
60%
80%
out
ECO
2
Wnet
(18)
where Wnet is the net power output of steam turbines from the HRSG
system (kW), which is calculated by using Eq. (17); WHPT , WIPT , and
WLPT are the power outputs from each steam turbine, corresponding to
the high-pressure turbine, intermediate-pressure turbine, and lowpressure turbine (kW), respectively; WPump is the energy requirement of
out
the water pump (kW), and ECO
is the mass flow rate of CO2 emitted
2
−1
into the environment (kg h ).
Table 3 shows the performance at various BRs for the pre- and postCaL cases. It has been shown that there is a linear relationship between
the BR and net electricity production. Again, a lower BR means more
coal content in the feedstock, thus producing more syngas from the cogasification system, which provides more heat for power generation
5
0
(17)
100%
BR(%)
(b)
Fig. 9. Distributions of (a) overall system efficiency and (b) specific CO2
emissions at various BR for pre- and post-CaL cases.
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Energy Conversion and Management 174 (2018) 489–503
P.-C. Kuo et al.
120
100
115.00
110.00
105.00
100.00
95.00
90.00
85.00
80.00
100
90
90
CO2 capture efficiency (%)
H 2 flow rate (kmol hr-1 )
110
80
70
60
0
20
40
60
BR (w
t% )
80
1001
1.5
2
2.5
3.5
3
CaO
4
4.5
90.00
85.00
80.00
75.00
70.00
65.00
60.00
55.00
50.00
45.00
40.00
35.00
80
70
60
50
40
30
20
10
5
0
0
20
/F
BR 40
(wt% 60
)
80
100 1
Fig. 10. Effects of CaO/F ratio on the hydrogen production at various BR for
pre-CaL case.
1.5
2
2.5
3
3.5
CaO
4
4.5
5
/F
(a)
during the HRSG process. This is thus the reason why the net electricity
production decreases linearly with an increase in BR. Fig. 9 demonstrates the profiles for overall system efficiency and specific CO2
emissions under various BRs for two types of co-gasification power
plants. In Fig. 9a, the values of the overall system efficiency for the preand post-CaL cases range from 49.76 to 51.74% and 51.55 to 52.22%,
respectively. Apparently, the overall system efficiency of the former
case is always higher than that of the latter case, regardless of kind of
BR is examined. In addition, for each scheme, it is noteworthy that a
slight increase in overall system efficiency is observed with increases in
BR. This arises from the fact that a higher BR contributes to lower energy input (i.e. the higher heating value of fuel) in the co-gasification
power plant, resulting in higher overall system efficiency. This observation is consistent with the results of Adnan and Hossain [37],
where they evaluated the system performance of co-gasification of coal
and microalgae. As a whole, blending torrefied biomass with coal can
promote the overall system efficiency of a co-gasification power plant
incorporating CaL process. With regard to specific CO2 emissions for the
two configurations, Fig. 9b indicates that the higher the BR, the lower
the specific CO2 emissions. However, it is not surprising that the values
of specific CO2 emissions for the pre-CaL case is more than two times
higher than that seen for the post-CaL case, stemming from the combustion system implemented behind the CaL unit for pre-CaL design.
Most importantly, as compared to Fig. 8c for the pre-CaL case, it can be
clearly seen that if the hydrogen generated from the CaL unit is further
combusted for the purpose of generating power, the system efficiency
will decrease from the range of 63 to 65.5% (HTE) to the range of 49.76
to 51.74% (electric efficiency). This implies, in turn, that the pre-CaL
case is less effective with regard to generating electricity and simultaneously suppressing CO2 emissions. It is thus concluded that the preCaL design is not proper to integrate with a combustion system intended to produce electricity, but the proposed pre-CaL scheme is recommended for co-generation of hydrogen and electricity. This will be
discussed in detail in the following section.
100
CO2 capture efficiency (%)
90
95.00
90.00
85.00
80.00
75.00
70.00
65.00
60.00
55.00
50.00
45.00
40.00
35.00
80
70
60
50
40
30
20
10
0
0
20
BR 40
(wt% 60
)
80
100 1
1.5
2
2.5
3
3.5
CaO
4
4.5
5
/F
(b)
Fig. 11. Effects of CaO/F ratio on the CO2 capture efficiency: (a) pre-CaL, and
(b) post-CaL.
carried out, the effect of the amount of CaO sorbent used in the CaL
process on the performance of the aforementioned indexes is discussed.
To obtain the optimal conditions for the CaL process, the CaO to fuel
mass flow rate ratios (CaO/F) are investigated in the range of 1–5 and
are calculated as follows:
CaO/ F =
ṁ CaO
ṁ fuel
(19)
where ṁ CaO is the mass flow rate of the CaO recirculated to the carbonator (kg h−1).
The three-dimensional profile of the CaL process performance in
terms of hydrogen production, CO2 capture efficiency, and energy
consumption as a function of the BR and CaO/F ratio is shown in
Figs. 10–12. It can be seen in Fig. 10 for the pre-CaL case that hydrogen
3.3. Optimal operating conditions and performance
By virtue of the fact that the amount of CaO sorbent is a significant
parameter affecting the reaction extent of CaO when carbonation is
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P.-C. Kuo et al.
BR = 40% when the CaO/F ratio varies from 1 to 3.5, accounting for a
13.7 and 13.97% enhancement in hydrogen production. This can be
elucidated by the fact that the SE-WGS reaction (R(1) and R(2) is intensified with increases in the amount of CaO. According to Le Chatelier’s principle, the removal of CO2 from carbonation (R1) can drive the
forward WGS reaction (R2), resulting in favorable conditions in which
to produce more hydrogen. However, when the CaO/F ratio is higher
than 3.5, i.e. CaO/F ≥ 3.5, the hydrogen formation remains almost
constant at each BR ratio. With attention paid to the impact of CaO/F
on the CO2 capture efficiency, Fig. 11a (pre-CaL case) and b (post-CaL
case) indicate that the CO2 capture efficiency has a markedly increasing
trend at first as the CaO/F ratio rises from 1 to 4, at which over a 90 and
97% of CO2 capture efficiency for pre- and post-CaL cases, respectively,
can be achieved. Nevertheless, with further increases in CaO/F to 5, the
CO2 capture efficiencies are insensitive to the variations in the CaO/F
and BR. With regard to the influence of CaO/F on the energy consumption from the calciner, Fig. 12a and b show that the energy requirements for the post-CaL case are obviously higher than those for the
pre-CaL case as a result of the larger amount of CO2 generated from the
combustion unit. It also should be addressed that the energy consumption is highly related to the CaO/F. Specifically, the higher the
CaO/F, the higher the energy penalty. Therefore, the CaO/F must be
minimized in order to achieve higher overall system energy efficiency.
From the above observations, from the viewpoint of hydrogen
production (only for pre-CaL case), overall energy efficiency, CO2
capture efficiency, and specific CO2 emissions, the BR of 40 wt% is a
recommended operating condition for both configurations. Similar
observations were also found in the previous study [17], where a cogasification of torrefied wood and coal power generation system
without CO2 capture was investigated. With regard to the optimum
amount of CaO sorbent, the condition of CaO/F = 3.5 for both cases is
suggested, at which the maximum CO2 capture efficiency of pre- and
post- CaL cases can reach 90.46 and 97.84, respectively. The results of
all performance indicators under the recommended optimal operating
conditions are summarized in Table 4. Under the optimal operating
conditions, it can be seen that the hydrogen production rate and hydrogen thermal efficiency are as high as 105.81 kmol hr−1 and 52.33%,
respectively, with a CO2 capture efficiency of 90.34% for the pre-CaL
case, while the CO2 capture efficiency is up to 97.69% with an overall
system energy efficiency of 51.41% for the post-CaL case.
3500
2800.00
2600.00
2400.00
2200.00
2000.00
1800.00
1600.00
1400.00
1200.00
Energy requirements (kW)
3000
2500
2000
1500
1000
500
0
0
20
BR 40
(wt% 60
)
80
100 1
1.5
2
2.5
3
3.5
CaO
4
4.5
5
/F
(a)
3500
3000.00
2800.00
2600.00
2400.00
2200.00
2000.00
1800.00
1600.00
1400.00
1200.00
Energy requirements (kW)
3000
2500
2000
1500
1000
500
0
0
20
BR 40
(wt% 60
)
80
100 1
1.5
2
2.5
3
3.5
CaO
4
4.5
5
/F
3.4. Optimal design of integrated plants
(b)
As can be observed in Fig. 12, it is apparent that regeneration of
CaO sorbent (calcination) is an extensively endothermic process.
Therefore, in order to supply a great amount of the energy requirements
for the calciner (Fig. 12), the optimal design of the integrated plants for
pre- and post-CaL is eventually proposed. As mentioned previously, it is
adequate for the pre-CaL case to be designed as a co-generation of
hydrogen and electricity plant, and thus the performance of such a
plant is evaluated here. Under the optimal conditions (BR = 40% and
CaO/F = 3.5), the energy duties from the endothermic calcination for
pre- and post- CaL are 2709.62 and 2918.55 kw, respectively. As a
consequence, a splitter is adopted to split a specific amount of hydrogen-rich gas (pre-CaL) or syngas (post-CaL) to an additional combustor, providing enough energy for the calciner. Figs. 13 and 14 demonstrate the layout of the optimal design of integrated plants for a cogeneration of hydrogen and electricity plant with CaL CO2 capture (i.e.
the pre-CaL case) and a co-gasification power plant with CaL CO2
capture (i.e. the post-CaL case), respectively. Tables 6 and 7 give the
stream results of the main flows shown in the flow diagram for the preCaL case (Fig. 13) and the post-CaL case (Fig. 14), respectively.
Fig. 12. Effects of CaO/F ratio on the energy requirements of the calciner: (a)
pre-CaL, and (b) post-CaL.
Table 4
Comparisons of performance indicators under the optimal operating conditions
for pre- and post-CaL cases. (BR = 40 wt% and CaO/F = 3.5).
−1
Hydrogen production rate (kmol h
Hydrogen thermal efficiency (%)
Overall energy efficiency (%)
CO2 capture efficiency (%)
Specific CO2 emissions (g/kWh)
)
Pre-CaL
Post-CaL
105.81
63.81
50.24
90.34
26.25
16.12
9.51
51.41
97.69
6.48
formation depends strongly on the CaO/F, regardless of the BR ratio.
For instance, the hydrogen production increases from 97.55 to 110.91
kmol hr−1 at BR = 20% and from 92.84 to 105.81 kmol hr−1 at
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Energy Conversion and Management 174 (2018) 489–503
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6 H2-rich gas: H2 105.81 kmol hr-1
CO 3.75 kmol hr-1
CO2 2.88 kmol hr-1
7 Concentrated CO2:
46.21 kmol hr-1
5 Raw syngas: H2 60.91 kmol hr-1
CO 48.65 kmol hr-1
CO2 1.80 kmol hr-1
3 H2O: 42.04 kg hr-1
B8-Calciner
B7-Carbonator
B5-Cooler
4 Gasifying agent:
Steam 780 kg hr-1
1 Fuels:
Torrefied biomass 400 kg hr-1
Acid gases
2 Fuels:
Coal 600 kg hr-1
Heat
B1-Drying
B2-Decomposition
B6-Separator
B3-Gasifier
B4-SSplit
8 Regenerated CaO:
62.41 kmol hr-1
9 Steam
17 Concentrated CO2:
46.21 kmol hr-1
10 O2 :24.38 kmol hr-1
B8-Combuster
HRSG
11 H2-rich gas: H2 .37.03 kmol hr-1
CO 1.31 kmol hr-1
-1
18 H2-rich gas: H2 .68.78 kmol hr
CO 2.44 kmol hr-1
CO2 2.88 kmol hr-1
13 Steam
14
HP
15
G
LP
IP
B-10 Condenser
12 H2-rich gas: H2 .68.78 kmol hr-1
CO 2.44 kmol hr-1
CO2 2.88 kmol hr-1
B9-Splitter
16 Feed water to HRSG:
2198.14 kg hr-1
B-11 Pump
Fig. 13. The optimal design of a co-generation of hydrogen and electricity power plant with CaL process (pre-CaL scheme).
Notably, Fig. 13 reveals a split fraction of 0.35 (B9) for the hydrogenrich gas stream (stream 6), which means a mole flow rate of 38.34 kmol
hr−1 (stream 11), fed to the combustor can completely recover the
energy requirements of the calciner for the pre-CaL case. Similarly,
Fig. 14 illustrates that a split fraction of 0.29 (B7) for the syngas stream,
which represents a mole flow rate of 31.77 kmol hr−1 (stream 7), can
be sent to the combustor for the purpose of recovering the duties consumed in the calciner for the post-CaL case. The performance of each
optimal design is outlined in Table 5. From the viewpoint of hydrogen
production and HTE, the pre-CaL design is significantly superior to that
5 Raw syngas: H2 60.91 kmol hr-1
CO 48.65 kmol hr-1
CO2 1.80 kmol hr-1
3 H2O: 42.04 kg hr-1
Performance indicator
Pre-CaL
Post-CaL
Split fraction (stream to combustor)
Hydrogen production rate (kmol h−1)
Hydrogen thermal efficiency (%)
Electric efficiency (%)
CO2 capture efficiency (%)
Specific CO2 emissions (g/kWh)
0.35
68.78
52.33
4.41
91.46
180.49
0.29
11.51
–
46.23
97.22
8.59
6 Syngas:H2 43.25 kmol hr-1
CO 34.51 kmol hr-1
CO2 1.27 kmol hr-1
B5-Cooler
4 Gasifying agent:
Steam 780 kg hr-1
1 Fuels:
Torrefied biomass 400 kg hr-1
2 Fuels:
Coal 600 kg hr-1
Table 5
The performance of optimal design of integrated plants.
B7-Splitter
8 Air: N2 135.25 kmol hr-1
O2 35.95 kmol hr-1
Acid gases
B1-Drying
B9-Combuster
B6-Separator
B2-Decomposition
B3-Gasifier
B4-SSplit
7 Syngas:H2 17.66 kmol hr-1
CO 14.11 kmol hr-1
CO2 0.52 kmol hr-1
11 CO2-lean gas:H2 11.51 kmol hr-1
CO 0.044 kmol hr-1
CO2 0.93 kmol hr-1
10 Flue gas:
H2 2.42 kmol hr-1
CO 9.14 kmol hr-1
CO2 26.82 kmol hr-1
H2O 43.15 kmol hr-1
B10Carbonator
12 Concentrated CO2
49.67 kmol hr-1
B11-Calciner
B8-Combuster
9 O2 :17.65 kmol hr-1
18 Exaust gas:H2 11.51 kmol hr-1
CO 0.044 kmol hr-1
CO2 0.93 kmol hr-1
HRSG
19 Concentrated CO2
49.67 kmol hr-1
14 Steam
HP
15
Heat
16
IP
LP
13 Regenerated CaO:
62.41 kmol hr-1
G
B-12 Condenser
17 Feed water to HRSG:
20982.19 kg hr-1
B-13 Pump
Fig. 14. The optimal design of a co-gasification of torrefied biomass and coal power plant with CaL process (post-CaL scheme).
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Energy Conversion and Management 174 (2018) 489–503
P.-C. Kuo et al.
environment as compared to the post-CaL design. From the viewpoint
of electric efficiency and CO2 capture efficiency, the post-CaL design
can ensure the lowest CO2 emissions and simultaneously achieve high
electric efficiency. However, assuming biofuels are potentially carbonneutral, blending 40 wt% torrefied biomass with coal as a fuel in the
two types of proposed systems might achieve a nearly zero CO2 emissions power plant, especially in the case of post-CaL design.
Table 6
Simulation results of the pre-CaL scheme (Fig. 13).
Stream
number
Type
Temperature (°C)
Pressure
(atm)
Mass flow
(kg h−1)
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
Torrefied biomass
Coal
H2O
Steam
Syngas
H2-rich gas
Concentrated CO2
Regenerated CaO
Steam
O2
H2-rich gas
H2-rich gas
Steam
Steam
Steam
H2O
Concentrated CO2
H2-rich gas
25
25
105
200
900
575
816.6
816.6
575
25
575
575
450
450
500
25
138.8
127.3
1
1
1
1
1
1
1
1
1
1
1
1
110
80
40
1
1
1
400
600
42
780
1738
445.2
2033.6
3500
757.5
780
155.8
289.4
2200
2200
2200
2198.1
2033.6
289.4
4. Conclusions
Two types of co-gasification of torrefied biomass and coal plant with
a calcium looping CO2 capture (pre-CaL and post-CaL) were proposed
and compared with each other in this work. The SE-WGS reaction was
found to have a greater impact on the pre-CaL scheme, leading to much
more hydrogen formation and higher hydrogen thermal efficiency as
compared to the post-CaL scheme. Considering the CO2 capture efficiency, specific CO2 emissions, and overall energy efficiency, the postCaL scheme is significantly superior to the pre-CaL scheme. Specifically,
the CO2 capture efficiency values can reach as much as 93–97%, which
is greater than that of the pre-CaL scheme (86–91%). From the viewpoint of specific CO2 emissions and overall system efficiency, the preCaL design is recommended for co-generation of hydrogen and electricity. Overall, the recommended operating conditions are BR = 40 wt
% and CaO/F = 3.5 for both the pre- and post- CaL cases. A split
fraction of 0.35 for the pre-CaL case and 0.29 for the post-CaL case can
be used to split a specific amount of hydrogen-rich gas (pre-CaL) or
syngas (post-CaL) into an additional combustor to completely cover the
energy demand. It is thus summarized that the pre-CaL case represents
as a highly efficient co-generation of hydrogen and electricity plant
with low CO2 emissions, whereas the post-CaL case is appropriate for
use in a co-gasification of torrefied biomass and coal power plant with
nearly zero CO2 emissions.
Table 7
Simulation results of the post-CaL scheme (Fig. 14).
Stream
number
Type
Temperature (°C)
Pressure
(atm)
Mass flow
(kg h−1)
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
Torrefied biomass
Coal
H2O
Steam
Syngas
Syngas
Syngas
Air
O2
Flue gas
CO2-lean gas
Concentrated CO2
Regenerated CaO
Steam
Steam
Steam
H2O
Exhaust gas
Concentrated CO2
25
25
105
200
900
600
600
25
25
600
600
862.56
862.56
570
570
521
25
128.3
132.7
1
1
1
1
1
1
1
1
1
1
1
1
1
110
80
2
1
1
1
400
600
42.04
780
1738
1157.7
472.9
4939.3
564.7
6097.0
4557.4
2577.2
3500
21,000
21,000
21,000
20,982
4557.4
2577.2
Acknowledgments
The authors would like to thank the Ministry of Science and
Technology of the Republic of China (Taiwan) for its partial financial
support of this research under grant MOST 107-3113-E-006-009.
of the post-CaL design, but a greater amount of CO2 is emitted into the
Appendix A
Minimum fluidize velocity (cm s−1) [38]:
umf = 0.00075
(ρs −ρg ) gdp2
μ
Ar < 103
(1)
Terminal velocity (cm s−1):
-1
2.335−1.744φ ⎤
18
uT = ⎡
⎢ (d ∗)2 +
⎥ 0.5 < φ < 1
(dp∗)0.5
p
⎣
⎦
Elutriation velocity (cm s
(2)
−1
) [39]:
2
u
kelu = 1.1ρs ⎛1− T ⎞
u⎠
⎝
(3)
Reynolds number (Re)
Rep =
dp u
νf
(4)
Archimedes number (Ar)
501
Energy Conversion and Management 174 (2018) 489–503
P.-C. Kuo et al.
Ar =
gρg (ρs −ρg ) dp3
μ2
(5)
The solid faction in the lower dense region [40]:
fd = (1−εb)(1−εmf )(1−0. 14Rep 0.4Ar −0.13)
(6)
The solid faction in the upper lean region [41]:
fl = fe + (fd −fe )−0.02h
(7)
The height of the dense and lean region [40]:
Wd =
∫0
Wl =
∫H
Hd
Ht
d
Aρs fd dh
(8)
Aρs fl dh
(9)
WT = Wd + Wl
(10)
Hl = Ht −Hd
(11)
For a first-order irreversible reaction, mass balance of reactant i over the emulsion, bubble, and cloud phases [42]:
dCb, i
= −(fb Rb, i )−(Cb, i−Cc, i )
dt
(12)
Kbc (Cb, i−Cc, i ) = fc R c, i + K ce (Cc, i−Ce, i )
(13)
K ce (Cc, i−Ce, i ) = fe R e, i
(14)
Rb, i = −kb, i Cb, i kb, i = fb kcat
(15)
R c, i = −kc, i Cc, i
kc, i = fc kcat
(16)
R e, i = −ke, i Ce, i
ke, i = fe kcat
(17)
⎛
dCb, i
= −kcat Cb, i ⎜fb +
⎜
dt
⎜
⎝
1
k cat
Kbc
+
1
fc +
1
1 / fe + k cat / K ce
⎞
⎟
⎟
⎟
⎠
(18)
The contact efficiency in the dense region [41]:
ηd =
fcore +
1
k cat / KCW + 1 / fwall
fcore + fwall
(19)
For the dense region of the fast fluidized bed reactor, concentration of reactant A:
dCAd, i
1
⎞ CAd, i
= ⎜⎛fcore kcat +
⎟
dh
1/ K CW + 1/ fwall kcat ⎠ u
⎝
(20)
with the boundary condition:
CAd, i = Ci,0
at h = 0
For the lean region of the fast fluidized bed reactor, concentration of reactant A:
CAl, i
dCAl, i
= ηl fl kcat
u
dh
(21)
with the boundary condition:
CAl, i = Ci,exit at h = Hd
The contact efficiency in the lean region [41]:
ηl = 1−(1−ηd )0.0662h
(22)
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