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Journal of Cleaner Production 201 (2018) 974e987
Contents lists available at ScienceDirect
Journal of Cleaner Production
journal homepage: www.elsevier.com/locate/jclepro
Carbon footprint reduction of acid gas enrichment units in hot
climates: A techno-economic simulation study
Satyadileep Dara a, Aisha A. AlHammadi a, Abdallah S. Berrouk a, b, e, Fadi Al Khasawneh c,
Abdulla Al Shaiba d, Yasser F. AlWahedi a, e, *
a
Chemical Engineering Department, Khalifa University of Science and Technology, PO Box 2533, Abu Dhabi, United Arab Emirates
Mechanical Engineering Department, Khalifa University of Science and Technology, PO Box 2533, Abu Dhabi, United Arab Emirates
Research & Development Department, Abu Dhabi National Oil Company, PO Box 898, Abu Dhabi, United Arab Emirates
d
Projects and Engineering Division, Al Yasat Petroleum Operations Company Ltd., PO Box 44476, Abu Dhabi, United Arab Emirates
e
Center for Catalysis and Separation, Khalifa University of Science and Technology, PO Box 127788, Abu Dhabi, United Arab Emirates
b
c
a r t i c l e i n f o
a b s t r a c t
Article history:
Received 16 April 2018
Received in revised form
10 July 2018
Accepted 7 August 2018
Available online 8 August 2018
In natural gas processing plants, acid gas enrichment (AGE) units play a vital role in increasing H2S purity
in the acid gas feed of sulfur recovery units (SRUs). Moreover, AGE units also produce a CO2-rich gas
stream that is often vented to atmosphere. If CO2 purity is sufficiently high, this stream can be used as an
injection gas for enhanced oil recovery (EOR) or for sequestration. In hot climates, AGE units operate at
significantly low efficiencies owing to the exothermic nature of their operation. Any enhancement in the
efficiency can reap significant benefits. In this work, we study the economic and environmental impact of
a process scheme wherein a RanqueeHilsch vortex tube (RHVT) is used as a cooling system for a lean
solvent in an AGE unit located in a hot region of the United Arab Emirates. A simulation model is built
using the process simulator ProMax® and is validated using plant design data. It is found that reducing
the lean solvent temperature increased the purity of H2S and CO2 product streams. At temperatures
lower than 25 C, the inverse occurs as CO2 absorption becomes favorable thermodynamically. Consequently, a lean solvent temperature of 25 C is identified to be optimal, thus achieving the lowest energy
consumption and carbon footprint, while maintaining high purities of the product gases. At the optimal
temperature, the proposed scheme results in steam savings of 13 kg/s (equivalent to 40% reduction in
total steam rate). This reduced energy consumption leads to an annual CO2 footprint reduction of 83.7
million kg (equivalent to 40% reduction in total CO2 footprint). The optimal lean solvent temperature
increases the purity of the H2S-rich gas stream (acid gas) to 67.3 mol% compared to its base case value of
45.7 mol%. Further, the purity of CO2-rich gas stream increases to 97 mol% compared to its base case value
of 89 mol%, thus making it suitable for EOR or sequestration. Economically, the evaluated annual energy
savings translate to 11.2 million USD, at a crude oil price of 50 USD. The computed payback period is 1.3
years, thus showing the potential of the proposed process. The process scheme proved to be superior to
other commercial alternatives from economic and environmental perspectives.
© 2018 Elsevier Ltd. All rights reserved.
Keywords:
Acid gas enrichment
RanqueeHilsch vortex tube
CO2 footprint
Process simulation
1. Introduction
Acid gas enrichment (AGE) units play a vital role in any natural
gas complex as they separate undesired acid gases (CO2 and H2S) to
ensure compliance with safety and operational requirements.
These units produce two major product streams: (i) a H2S-rich gas
stream, which is used as a feed gas for sulfur recovery units (SRUs)
* Corresponding author. Chemical Engineering Department, Khalifa University of
Science and Technology, PO Box 2533, Abu Dhabi, United Arab Emirates.
E-mail address: yaalwahedi@pi.ac.ae (Y.F. AlWahedi).
https://doi.org/10.1016/j.jclepro.2018.08.067
0959-6526/© 2018 Elsevier Ltd. All rights reserved.
and (ii) a CO2-rich gas stream, which is often vented to the
atmosphere.
1.1. Motivation
CO2-rich gas stream, which is the major product stream of the
AGE, has significant potential applications. If the purity of the
former is sufficiently high, this stream can be used as an injection
gas for enhanced oil recovery (EOR) or can be sequestrated. This has
triggered the investigation of various schemes and configurations
to increase the purity of the CO2-rich gas stream to allow its use in
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
the aforementioned applications. Moreover, any enhancements in
the operational efficiency and reduction in carbon footprint are
desired. To this end, this paper focuses the attention on a commercial AGE plant, which operates on HIGHSULF process (a
patented process that involves recycling a portion of acid gas to the
gas absorber), based in United Arab Emirates and attempts to
improve the performance in the aspects of operating costs and
carbon footprint.
1.2. Literature survey
Several studies were reported that assessed various techniques
to improve the performance of AGE units in terms of product gas
purity and energy consumption. Weiland and Khanmamedov
(2010) carried out a process simulation study of three AGE processes: 1) the HIGHSULF process, 2) the AGE process with two
absorbers, and 3) the conventional AGE process. They concluded
that the HIGHSULF process results in a more concentrated (19%
higher) acid gas stream compared with the conventional AGE
process and the AGE process with two absorbers. Khanmamedov
(2013) also recommended the HIGHSULF process over the other
processes due to a potential reduction of 8% in capital cost of the
downstream SRU besides the increased acid gas purity. To this end,
the above two studies conclude that HIGHSULF process is the most
efficient scheme of the available AGE processes. On the other hand,
it is worth to note that the commercial AGE unit being studied in
this paper operates on the HIGHSULF process thus leaves no scope
for any potential performance improvement by means of retrofitting with a different AGE scheme. Parks et al. (2010) mentioned that
the use of FLEXSORB SE and FLEXSORB SE PLUS can replace MDEA
(Methyldiethanolamine)-based solvents owing to their higher
selectivity for H2S, thereby increasing the purity of the H2S-rich gas
stream from 30 mol% to 65 mol%. However, this also leads to a
major increase in the operating expenses due to high cost of proprietary solvents. In general, such schemes require major modifications to the plant hardware rather than simple changes in
operating conditions or minor modifications. The latter are often
more attractive as they do not involve any significant downtime in
unit operation. However, studies on the impact of major process
variables are lacking in the field of AGE. In contrast, in the field of
natural gas sweetening, several studies have been conducted on the
impact of major operating variables. A lean solvent temperature has
been reported to have the strongest influence on plant performance. For instance, Lunsford and McIntyre (1999) showed that
decreasing the absorber temperature from 74 C to 30 C leads to a
decrease in sweet gas H2S concentration from 6000 ppm to 10 ppm.
Addington and Ness (2009) performed a simulation study of a
Russian gas plant and showed that a reduction in lean solvent
temperature from 50 C to 10 C resulted in a decrease in the concentration of sweet gas H2S from 28 ppm to 4 ppm. Pandey (2005)
carried out an on-site parametric study of a commercial amine
sweetening unit (Hazira Gas Processing Plant, Oil and Natural Gas
Corporation Limited, India) and demonstrated that H2S absorption
selectivity over CO2 can be increased by 25% by decreasing the lean
solvent temperature from 50 C to 39 C. Dara and Berrouk (2017)
performed a process optimization study on a middle-east-based
commercial amine sweetening unit and identified the optimum
conditions of major operating variables. They concluded that low
solvent temperatures are suitable for better performance of the
absorber owing to the increased solubility of acid gas governed by
vapor liquid equilibrium. They showed that steam consumption
rate can be reduced by 13.5% as a result of changing the lean solvent
temperature from 65 C to 50 C. They reported that the concentration of sweet gas CO2 decreases with lean amine temperature
until 54 C and then follows a reverse trend for further reduction
975
until 50 C. They concluded that the optimum choice of lean solvent
temperature is not straightforward and should be investigated
thoroughly in order to achieve optimum energy consumption and
appropriate balance of absorption degrees of H2S and CO2. The
aforementioned studies in natural gas sweetening units show that
optimizing the lean solvent temperature can achieve two benefits:
(i) energy savings that would thus lead to cost savings and (ii)
increased purity levels of H2S and CO2 streams.
In hot climates where refrigeration costs are high, lowering the
solvent temperature can be achieved by means of a RanqueeHilsch
vortex tube (RHVT) provided a high pressure inert gas is available.
The RHVT is a static device that uses vortex motion to separate a
compressed gas flow into a hot stream and a cold stream. Saberi
et al. (2016) reported the improvement of surface grinding process performance through the minimum quality lubrication (MQL)
technique coupled with a simple and inexpensive vortex tube by
achieving a cold air jet temperature of /4 C compared with the
conventional coolant temperature of 15.5 C. Jozic et al. (2015)
investigated the application of the RHVT to produce cold air in
the end milling process. They reported that the cutting speed
increased by 50% upon using the RHVT for cooling, which enhanced
process efficiency.
Also, many novel techno-economic studies were carried out for
investigating various innovative methods involving utilization of
ambient conditions for energy conservation purposes. Esen et al.
(2006) performed performance experiments and economic analysis of a horizontal ground source heat pump (GSHP) system for
space heating. They proved that this system offers several economic
advantages over many other conventional techniques. Esen et al.
(2007) reported a techno-economic comparison between a
ground-coupled heat pump (GCHP) system and an air-coupled heat
pump (ACHP) system for space cooling. They concluded that GCHP
systems are economically preferable to ACHP systems for the purpose of space cooling. Esen and Yuksel (2013) carried out experimental investigation of greenhouse heating by biogas, solar and
ground energy in Elazig, Turkey climate conditions. They reported
the effects of climatic conditions and operating parameters on the
system performance parameters.
1.3. Contribution
In this work, we propose the application of RHVT for cooling the
lean solvent in an AGE unit based in the United Arab Emirates.
Pressurized nitrogen gas, which is readily available in the plant as
the by-product of the air separation unit in the same plant complex,
is used as the RHVT feed. A simulation model is built for the
combined system by using the process simulator ProMax® and is
validated using the AGE plant design data. The model is used for
studying the impact of lean solvent temperature on the product
streams. Then, potential energy savings of the proposed scheme are
evaluated. The economic feasibility and environmental impact of
the RHVT-based AGE unit are assessed by considering achievable
reductions in operating cost, energy consumption and carbon
footprint. To the best knowledge of the authors, the integration of
RHVT with AGE and the evaluation of the associated benefits have
not been studied in the reported literature. Also, the proposed
process scheme is compared to other commercial processes in order to benchmark its potential benefits against commercial options.
2. Process description
2.1. Acid gas enrichment unit
Fig. 1 depicts the process flow diagram of the commercial AGE
unit considered in this study. In the commercial AGE unit, feed gas
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S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
Fig. 1. Process flow diagram of the commercial AGE unit.
at 180 kPa and 57 C enters the absorber bottom, while the lean
MDEA solvent is fed to the top of the absorber. There are three
provisions for solvent entry at the 9th, 11th, and 13th trays. During
normal operation, the 13th tray is used. In the absorber, the downflowing MDEA solution absorbs H2S and CO2 from the ascending
sour gas to produce a sweetened gas stream. The ascending gas
enters the wash water section of the column. In this section, the gas
is washed with water in order to minimize any solvent carryover in
the gas phase. CO2-rich gas leaves the absorber overhead at 160 kPa
and 55.1 C. Then, it enters the absorber off-gas knock out drum,
where any condensed water is separated. Rich MDEA solvent leaves
the absorber from the lowest tray. Then, it is pumped at 900 kPa to
the lean/rich solvent heat exchanger where the stream is heated to
117 C by exchanging heat with the hotter lean solvent exiting the
regenerator. The hot rich MDEA enters the regenerator at the top
section (tray 18). Low pressure (LP) steam is supplied to the
reboiler, which vaporizes a portion of water associated with the
MDEA solvent. Steam rising through the column regenerates the
MDEA by stripping out the evolved H2S and CO2. The overhead gas
leaving from the regenerator at 195 kPa and 112 C passes through
an air fan cooler, which cools the gas to 55 C, thus resulting in the
partial condensation of water. The two phases are separated in the
reflux drum, where the water is pumped to the regenerator as
reflux. H2S-rich gas at 185 kPa and 55 C splits into two streams.
Major split feeds the acid gas header while the other is recycled to
the absorber. The lean MDEA from the regenerator bottom is
pumped to the lean-rich heat exchanger, where it is cooled to 80 C.
A colder lean solvent is recombined and cooled further across a lean
solvent air fan cooler to 55 C before it enters the absorber again.
2.2. RanqueeHilsch vortex tube (RHVT)
The RHVT or the vortex tube is a static device used to separate a
pressurized gas into a colder stream and a hotter stream (Fig. 2).
Eiamsa-ard and Promvonge (2008) explained that this separation
of streams is known as the temperature separation effect or the
energy separation effect. Energy separation is the process of
redistribution of the total fluid energy flow without additional
Fig. 2. Schematic representation of the RanqueeHilsch vortex tube.
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
work or heat. As a result, a part of the fluid leaves with high energy,
whereas the other part leaves with low energy compared with the
inlet fluid. When the compressed gas is introduced into the tube
through the tangential nozzle, the nozzle forces the gas to flow in a
centrifugal motion through the tube. This type of motion increases
the pressure of the gas along the outer edge of the tube and a
reduction in pressure in the middle of the tube. The temperature
separation within the tube is induced by the friction experienced by
gas with the tube wall. Because the speed of the gas in the outer
edges is lower than that at the center, energy is transferred from the
center region to the tube wall region, thus generating a hot gas
stream near the wall and a cold gas stream at the center.
3. Methodology
This study proposes employing a RHVT for producing a cold
nitrogen stream, which can be used as a coolant stream in the lean
solvent-air cooler. Fig. 3 depicts the proposed scheme that integrates the RHVT and AGE unit. High-pressure nitrogen enters the
RHVT and splits into hotter and colder streams. The latter then
mixes with ambient air in an air-nitrogen mixer to provide a
coolant stream at sufficiently lower temperatures such that it cools
down the lean solvent to the desired levels.
The methodology employed in this study is as follows (see
Fig. 4):
1. Develop the simulation model for the commercial AGE unit
considered in this study and validate the model using plant data
to ensure an prediction-error level below 15%
2. Perform sensitivity analysis to examine the effect of lean solvent
temperature on plant performance and thereby identify the
optimum temperature
3. Evaluate the nitrogen flow rate required to meet the optimum
lean solvent temperature, through mass and energy balance
across the air-nitrogen mixer
4. Evaluate the carbon footprint reduction achieved via energy
savings
5. Perform economic analysis to assess the profitability of the
proposed scheme
977
To this end, the following four models are developed for carrying out the aforementioned methodology.
1
2
3
4
AGE simulation model
Lean solvent cooling system model
Carbon footprint model
Economics model
3.1. AGE simulation model
This section discusses the phenomena of acid gas absorption in
MDEA followed by the reaction kinetics based model employed to
simulate it.
The absorption of H2S and CO2 by MDEA from the gas phase
entails three steps:
1. Mass transfer of H2S and CO2 molecules from the bulk gas phase
to the gas/liquid interface
2. Dissolution of H2S and CO2 molecules from the gas phase into
the liquid phase, which is governed by thermodynamic
equilibrium
3. Reaction of H2S and CO2 molecules with MDEA in the liquid
phase
Kohl and Nielsen (1997) described the reaction chemistry of
MDEA with H2S and CO2 by using the following chemical reactions:
1) The MDEA reaction with H2S is given by
CH3N(C2H4OH)2 þ H2S )/ CH3(C2H4OH)2NHþ þ HS
(1)
2) The MDEA reaction with CO2 is given by
CO2 þ H2O þ CH3N(C2H4OH)2 )/ CH3(C2H4OH)2NHþ þ HCO
3 (2)
The reaction between H2S and MDEA occurs via a direct proton
transfer mechanism as shown in equation (1). This reaction is fast
Fig. 3. Proposed scheme involving the integration of the RHVT with the AGE unit.
978
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
Fig. 4. Methodology of the study on the proposed scheme.
and proceeds to equilibrium instantaneously. Furthermore, mass
transfer from the gas phase to the gas/liquid interface is fast.
Therefore, the H2S absorption is predominantly controlled by the
vapor/liquid equilibrium at the gas/liquid interface. On the contrary, CO2 reaction with MDEA is relatively slow. CO2 undergoes a
hydrolysis reaction to form carbonic acid first, which then undergoes acidebase reaction with MDEA to yield the overall reaction
shown in equation (2). Kidnay and Parrish (2007) explained the CO2
absorption phenomena with MDEA using the following reactions:
1) The CO2 hydrolysis step is given by
CO2 þ H2O )/ H2CO3
(3)
2) The Carbonic acid dissociation step is given by
H2CO3 )/ Hþ þ HCO
3
(4)
3) The MDEA protonation step is given by
CH3N(C2H4OH)2 þ Hþ )/ CH3(C2H4OH)2NHþ
(5)
CO2 dissolves in water and hydrolyses to form carbonic acid
(equation (3)), which in turn slowly dissociates to bicarbonate and
hydrogen ion (equation (4)). The latter then undergoes an
acidebase reaction with the amine to yield a protonated amine
(equation (5)). It should be noted that the dissociation of carbonic
acid to bicarbonate (equation (4)) is the rate limiting step. Vapor/
liquid equilibrium at the interface also plays a major role in controlling CO2 absorption.
For simulating the absorption of acid gases, this study employs
TSWEET kinetics package of ProMax simulator, licensed by Bryan
Research and Engineering Inc. This method accounts for the mass
transfer of acid gases (H2S and CO2) and liquid phase reaction kinetics through a combination of reaction kinetic models and ideal/
real tray ratios.
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
Below are the generalized rate equations of the TSWEET package
to model the relatively slow CO2 absorption.
1) The rate of forward reaction of the controlling step of CO2 absorption is given by
r ¼ k½H2 CO3 (6)
2) The rate constant of the forward reaction of CO2 absorption is
given by
k ¼ Aexp E=RT T b
(7)
where r is the rate of forward reaction, T is the reaction temperature, k is the rate constant, A is the frequency factor, E is the activation energy, and b is a constant quantifying temperature effect.
3) The rate of backward reaction of the controlling step of CO2
absorption is given by
i
h i h
r 0 ¼ k0 Hþ HCO
3
(8)
4) The rate constant of the backward reaction of the controlling
step of CO2 absorption is given by
0
k ¼ A exp E0=RT T b
0
0
(9)
where r 0 is the rate of backward reaction, k0 is the rate constant, A0 is
the frequency factor, E0 is the activation energy, and b’ is the constant quantifying temperature effect.
Further, R represents the gas constant and [H2CO3], [Hþ], &
[HCO3] represent the concentrations of the reactants.
979
5) Hence, the complete reaction rate of CO2 absorption is given by
i
h i h
rnet ¼ k½H2 CO3 k0 Hþ HCO
3
(10)
where rnet is the net rate of reaction of the controlling step of CO2
absorption.
The net absorption levels of CO2 directly influence the reaction
medium temperature, and hence, affect the H2S equilibrium solubility and impact the product gas H2S.
The ideal/real tray ratios are obtained based on industrial
experience. The ratio of ideal to real stages is assumed to be ~3 in
the absorber, as recommended by Bryan Research and Engineering
Inc. Similarly, the regenerator is modeled using the ‘TSWEET
alternate stripper’ package that assumes a ratio of 2. Therefore, the
absorber and regenerator are modeled with 8 and 11 ideal stages,
respectively.
To predict the vapor liquid equilibrium, the ProMax simulator
adopts a refined Soave Redlich Kwong (SRK) model. Soave (1972)
reported that this model uses the SRK equation of state for
modeling vapor phase interactions and the electrolytic non random
two liquid (NRTL) for predicting liquid phase properties. The ProMax simulation model for the AGE unit is depicted in Fig. 5.
3.2. Lean solvent cooling system model
To calculate the air and nitrogen flowrates required for generating the coolant stream, a standalone ProMax model for the air
mixture-lean solvent cooler is used. The following assumptions are
made in the model:
1. The heat capacity of the gas mixture (Cpmix Þ is equal to that of
nitrogen (CpN2 Þ.
2. The minimum end temperature approach across the cooler is
5 C.
Fig. 5. Process simulation model of the commercial AGE unit.
980
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
3. The supply temperature of nitrogen from the RHVT system
(TN2 in ) is 10 C.
4. The supply temperature of air ðTAir in Þ is 50 C, in order to
simulate a summer case in the UAE.
The governing model for the lean solvent cooling system is as
follows:
1) The energy balance across the lean solvent cooler is given by
Qsolvent cooler ¼ m_ mix Cpmix DTmix
(11)
where DTmix ¼ ðTmix in Tmix out Þ
After determining the total flow rate of the mixture, the mass
flow rates of cold nitrogen gas and air are computed by imposing
mass and energy balances across the cooler, as shown in Fig. 6:
2) The mass balance across the mixer is given by
m_ mix ¼ m_ Air þ m_ N2
(12)
3) The energy balance across the mixer is given by
m_ Air CpAir DTAir þ m_ N2 CpN2 DTN2 ¼ m_ mix Cpmix DTmix
(13)
where DTAir ¼ ðTAir in Tmix in Þ and DTN2 ¼ ðTN2 in Tmix in Þ
Based on the targeted mixture temperatures, the required
flowrates of air and nitrogen can be computed as follows:
m_ N2 ¼ m_ mix m_ mix Cpmix DTmix CpN2 DTN2
CpAir DTAir CpN2 DTN2
(15)
where m_ Air , m_ N2 , and m_ mix are the mass flow rates of the air
stream, nitrogen stream, and coolant mixture, respectively; CpAir ,
CpN2 , and Cpmix are the heat capacities of air, nitrogen, and the coolant
mixture, respectively; DTAir , DTN2 , and DTmix are the changes in
temperature for air, nitrogen, and the coolant mixture across the
heat exchanger, respectively; Tmix in and Tmix out are the mixture
inlet and outlet temperatures, respectively. TAir in and TN2 in are the
inlet temperatures of air and N2, respectively, and Qsolvent cooler is the
cooling duty of the lean solvent cooler.
3.3. Carbon footprint model
The carbon footprint reduction corresponding to given steam
savings is evaluated using equations (16) to (18). First, the theoretical energy required for producing the low pressure steam
(Qsteam ), at 600 kPa is calculated using equation (16) based on the
mass of water (mwater ), heat capacity (Cpwater ), saturation temperature ðTSaturation ), reference temperature (TRef ), and latent heat of
vaporization (l). The net heat required (Qnet ) is calculated using
equation (17) by assuming an overall efficiency (h) for the steam
generation process. Finally, the carbon footprint reduction
(mCarbon ) that corresponds to the reduction in steam consumption
rate is calculated based on the study by U.S. Energy Information
Administration, which reported the amount of CO2 (C ) emitted
per 1 kJ energy of natural gas.
4) The mass flow rate of air required is given by
m_ Air
m_ mix Cpmix DTmix CpN2 DTN2
¼
CpAir DTAir CpN2 DTN2
5) The mass flow rate of nitrogen required is given by
1) The theoretical energy required for steam generation is given by
(14)
Qsteam ¼ mwater Cpwater TSaturation TRef þ mwater l
2) The net energy required for steam generation is given by
Fig. 6. Illustration of mass and energy balance across air-nitrogen mixer.
(16)
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
Qnet ¼ Qsteam =h
(17)
981
Table 1
Summary of inputs for AGE simulation model.
AGE simulation model inputs
3) The carbon footprint reduction corresponding to the energy of
given amount of steam saved is given by
mCarbon ¼ Qnet C
(18)
3.4. Economics model
The total operating cost of an AGE unit is comprised of
1 Cost of steam consumption in the regenerator reboiler
2 Cost of electrical power consumed by the solvent pumps
3 Cost of power consumption of cooling fans of the lean solventair cooler
The power consumption across the solvent cooler is unaffected
by the change in the cooling mixture (air þ nitrogen) since the
cooling fans draw a constant volumetric flow rate. Furthermore, the
changes in solvent properties such as density and viscosity are
minor across the studied solvent temperatures, and hence, they do
not affect the pumping load. Therefore, the costs of electrical power
consumed by the solvent pump and cooling fans remain constant
with solvent temperature, and hence, are not considered as operating costs. Also, this study did not take the price of nitrogen into
account since the latter is readily available as the by-product of the
air separation unit in the same plant complex. To summarize,
operating cost savings are estimated solely based on the decreased
steam consumption of the regenerator reboiler. Monetary savings
due to steam reduction are evaluated based on the cost of natural
gas required for producing the steam. The required quantity of
natural gas (mnaturalgas ) is calculated using equation (19) based on
the heating value of natural gas (Н natural gas ) and the net heat
required (Qnet ) for producing the steam. The cost of the former is
evaluated using the correlation provided by the Abu Dhabi National
Oil company (equation (20)), which is a function of crude oil price
(CP).
1) The amount of natural gas corresponding to energy of given
amount of steam is given by
mnaturalgas ¼ Qnet
.
Н natural gas
Feed gas temperature, ºC
Feed gas pressure, kPa
Feed gas flow rate, kg/s
Feed H2S flow rate, kg/s
Feed CO2 flow rate, kg/s
Feed H2O flow rate, kg/s
Feed hydrocarbons flow rate, kg/s
Absorber pressure drop, kPa
Absorber overhead pressure, kPa
Actual number of absorber trays
Ratio of real to ideal trays (absorber)
Lean solvent flow rate, kg/s
Lean solvent concentration, wt %
Lean solvent temperature, oC
Number of regenerator trays
Regenerator pressure drop, kPa
Regenerator overhead pressure, kPa
Reboiler steam rate, kg/s
Reboiler steam pressure, kPa
Acid gas condenser temperature, oC
Ratio of real to ideal trays (regenerator)
57
180
33.4
4.6
27.4
1.3
0.1
20
160
18
3
277.8
40
55
22
55
195
32.5
600
55
2
Table 2
Summary of inputs for the lean solvent cooling system model, carbon footprint
model, and economics model.
Lean solvent cooling system model inputs
Heat capacity of air (CpAir Þ, kJ/kg. oC
Heat capacity of nitrogen (CpN2 Þ, kJ/kg. oC
Heat capacity of coolant mixture (Cpmix Þ, kJ/kg. oC
Minimum end temperature approach, oC
Cold nitrogen temperature ðTN2 in Þ, oC
Cold air temperature ðTair in Þ, oC
Carbon footprint model inputs
Heat capacity of water ðCpwater Þ, kJ/kg. oC
Steam pressure, kPa
Saturation temperature of steam ðTSaturation ), oC
Reference temperature (TRef ), oC
Latent heat of vaporization (l), kJ/kg
Overall process efficiency of steam generation (h )
Cold air temperature (Tair in ), oC
Amount of CO2 (C) emitted per 1 kJ of energy of natural gas, kg
Economics model inputs
Heating value of natural gas (Н naturalgas ), kJ/kg.
Costing parameter1 (a)
Costing parameter2 (b)
Costing parameter3 (g)
Unit conversion factorðCF)
1
1.039
1.039
5
10
50
4.184
600
158.8
20
2085
0.65
50
5 E5
42300
13.5
50
0.5084
1000
(19)
model.
2) The cost of processed natural gas per unit mass is given by
Cnaturalgas ðUSD=kgÞ ¼ ðða*ðCPÞ þ bÞ*gÞ = CF
3) The total cost of natural gas corresponding to energy of given
amount of steam is given by
TCnaturalgas ðUSDÞ ¼ Cnaturalgas * mnaturalgas
4. Results and discussion
(20)
(21)
where a, b, and g are the cost parameters for estimating the cost of
processed natural gas in the industry and CF is a conversion factor.
3.5. Case study inputs
Table 1 presents the inputs and parameters used for the AGE
simulation model, while Table 2 presents those of the lean solvent
cooling system model, carbon footprint model, and economics
The plant design data of the AGE unit (shown in Table 1) are
used to validate the simulation model. A comparison of the model
predictions and plant data is presented in Table 3, which shows that
all the process variables are predicted with an error less than 15%. It
should be noted that all the parameters are predicted with an error
significantly lower than 10%, except for the CO2 concentration in
the H2S-rich gas stream and the regenerator overhead. The error in
this case is 10.4%. The TSWEET package is more optimized for
predicting H2S absorption in MDEA solvents when compared with
CO2. This is because the accurate prediction of CO2 within an error
of 10% is usually not required in natural gas sweetening systems.
4.1. Parametric analysis
This section also discusses various results (Figs. 7a and 10)
982
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
Table 3
Summary of model validation with plant data.
Stream
Parameter
Model prediction
Plant data
% Error
CO2-rich gas (Product gas from absorber overhead)
Temperature (ºC)
Pressure (kPa)
Flow rate (kg/s)
H2S flow rate (kg/s)
CO2 flow rate (kg/s)
Temperature (ºC)
Pressure (kPa)
Flow rate (kg/s)
Temperature (ºC)
Pressure (kPa)
Flow rate (kg/s)
H2S flow rate (kg/s)
CO2 flow rate (kg/s)
Temperature (oC)
Pressure (kPa)
MDEA flow rate (kg/s)
H2S flow rate (kg/s)
CO2 flow rate (kg/s)
Temperature (ºC)
Pressure (kPa)
Flow rate (kg/s)
H2S flow rate (kg/s)
CO2 flow rate (kg/s)
55.1
160
21.7
0.0037
20.7
67.0
157
289.7
112.0
195
26.9
4.643
6.7
55
900
111.1
0.0166
0.0166
55
185
11.8
4.62
6.7
53.6
160
22.8
0.0038
21.4
68.5
157
289.1
111.8
195
28.6
4.640
6.0
55
900
111.0
0.0170
0.171
55
185
11.1
4.61
6.0
2.7
0
4.8
2.6
3.3
2.2
0
0.2
0.2
0
5.9
0.1
10.4
0
0
0.1
2.3
2.9
0
0
5.9
0.2
10.4
Rich solvent
Regenerator overhead
Lean solvent
H2S-rich gas (Acid gas)
produced using the models explained in section 3. The lean solvent
temperature is varied between 55 C and 21 C to investigate its
impact on the following process parameters: 1) CO2 purity of CO2rich gas stream (product gas from absorber overhead stream), 2)
H2S purity of H2S-rich gas stream (acid gas from regenerator
overhead), and 3) the steam consumption rate. It should be noted
that the concentration of H2S in the absorber overhead is maintained at ~500 ppmv throughout the analysis by varying the steam
rate in the regenerator reboiler. Fig. 7a depicts the effect of lean
solvent temperature on the CO2 content in the product gas. The
latter increases from 89 mol% to 97 mol% as the lean solvent temperature decreases from 55 C to 25 C. Afterward, the CO2 content
in the product gas follows an opposite trend and decreases to
96 mol% as the lean solvent temperature decreases to 21 C. Fig. 7a
also shows the effect of lean solvent temperature on H2S content of
the H2S-rich gas stream. The latter increases from 45 mol% to
69.3 mol% as the lean solvent temperature decreases from 55 C to
25 C. Thereafter, the H2S content in the H2S-rich gas stream follows an opposite trend and decreases to 65.5 mol% as the lean
solvent temperature decreases to 21 C. Within the temperature
range between 55 C and 25 C, H2S absorption process is governed
by vapor/liquid equilibrium at the phase interface, while CO2 absorption is governed by reaction kinetics in the liquid phase.
Because the dissolution of H2S from the gas phase to the liquid
phase is exothermic, lower temperatures favor higher H2S concentrations in the liquid phase, thus leading to higher absorption of
H2S. Within the same temperature range, CO2 absorption is dominated by the reaction kinetics in the liquid phase, which tends to
slow down with decreasing temperature. The combined effect of
both phenomena leads to higher purity levels in the CO2-rich gas
stream and H2S-rich gas stream. Below 25 C, the vapor/liquid
equilibrium of CO2 causes higher CO2 absorption in the solvent,
thus reducing the driving force available for H2S absorption. This
leads to a higher concentration of CO2 in the H2S-rich gas stream
and higher concentration of H2S in the CO2-rich gas stream, which
leads to lower purity levels.
For applications such as EOR and carbon sequestration in
depleted oil and gas reservoirs, Abbas et al. (2013) reported that a
CO2 purity of 95 mol% is essential. At 25 C, the proposed scheme
results in a product gas with a CO2 purity of 97 mol%, which is
significantly higher than the base case CO2 purity in the AGE unit
(89 mol%), thus making it suitable for the aforementioned
applications.
Fig. 7b depicts the effect of the lean solvent temperature on the
steam consumption rate. The steam consumption rate decreases
from 32.6 kg/s to 19.6 kg/s as the lean solvent temperature decreases from 55 C to 25 C, and then increases to 20 kg/s as the lean
solvent temperature decreases to 21 C. Between 55 C and 25 C,
the higher equilibrium driving force for H2S absorption combined
with slower CO2 reaction kinetics enhance the solvent affinity for
H2S. This implies that a lean solvent with a lower level of regeneration (higher remnant concentrations of absorbed H2S) can
achieve the desired product gas H2S content (500 ppmv). Therefore,
the steam consumption rate decreases owing to the lower level of
regeneration required. Similarly, for temperatures lower than 25 C,
the increased rate of CO2 absorption adversely affects the driving
force available for H2S absorption. This in turn demands a higher
level of solvent stripping to meet the desired H2S content in the
product gas. Therefore, the required steam rate increases.
To summarize, the optimum lean solvent temperature is 25 C
owing to the lowest steam consumption rate and the highest purity
levels of H2S-rich gas streams and CO2-rich gas streams.
Fig. 8 presents the required flow rates of air and nitrogen for
achieving various lean solvent temperatures, which are determined
using equations (11) to (15) mentioned in section 3.2. It is found
that 864 kg/s of nitrogen is needed to reduce the lean solvent
temperature to the optimum value of 25 C. Fig. 9 shows the net
steam savings and the corresponding reduction in carbon footprint
as a function of nitrogen flow rate. Steam savings and the corresponding carbon footprint increase with an increase in the nitrogen
flow rate until the optimum value of 864 kg/s and follow an
opposite trend thereafter. Steam savings of 13 kg/s and an annual
carbon footprint reduction of 83.7 million kg are achieved as a
result of the proposed scheme.
Fig. 10 depicts the potential annual monetary benefits for
various crude oil prices as a function of nitrogen flow rate used in
the proposed scheme. At a crude oil price of 50 USD, the evaluated
annual energy savings translate to 11.2 million USD.
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
983
Fig. 7. a)Effect of lean solvent temperature on the H2S content of H2S-rich gas (acid gas) and CO2 content of CO2-rich gas (product gas) (All other model parameters are as shown in
Tables 1 and 2). b): Effect of lean solvent temperature on the steam consumption rate (All other model parameters are as shown in Tables 1 and 2).
4.2. Economic analysis
4.3. Comparison with conventional processes
To investigate the profitability of the proposed scheme, it is
important to incorporate the capital cost of the RHVT for the given
process conditions. The cost of the RHVT is 2.38 million USD, as
stated by Tunkel (2017) (Universal Vortex, Inc., Designer of industrial vortex tubes). Further, a Lang factor (defined as the ratio of the
total cost of installation to the cost of the equipment) of 6.0 is
considered, as recommended by Peters et al. (2003). This leads to a
total installation cost of 14.3 million USD for the RHVT. Thus, the
proposed scheme can return its capital investment in 1.3 years.
This demonstrates the potential of the proposed scheme in
enhancing AGE unit economics in addition to reducing carbon
footprint. Furthermore, the improvement in the purity of CO2-rich
gas stream to the levels adequate for EOR application and carbon
sequestration justifies its implementation.
This section presents the performance of various conventional
processes (reported by Weiland and Khanmamedov (2010)) and
Khanmamedov (2013)) listed below.
1) HIGHSULF process, which is the process employed in the AGE
unit being studied
2) AGE process with separate feeds
3) AGE process with two absorbers
The simulation model shown in Fig. 5 is used for analyzing the
HIGHSULF process. However, the temperature of the solvent-air
cooler is maintained at 55 C, which is adequate for the ambient
temperature of the UAE. For the other two processes i.e. AGE process with two separate feeds and the AGE process with two absorbers, process simulation models shown in Figs. 11 and 12 are
984
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
Fig. 8. Effect of lean solvent temperature on the required mass flow rates of nitrogen and air (All other model parameters are as shown in Tables 1 and 2).
Fig. 9. Effect of nitrogen mass flow rate on the reductions in steam consumption rate and annual carbon footprint (All other model parameters except lean solvent temperature are
as shown in Tables 1 and 2).
developed using the same kinetic and thermodynamic methods as
that of Fig. 5.
These processes are compared with the proposed method in
various aspects such as product and acid gas purities, steam consumption rate, annual operating costs and capital investments. As
can be seen from Table 4, AGE process with two separate feeds and
the AGE process with two absorbers do not even meet the product
gas H2S concentration of 500 ppm despite increasing the reboiler
steam rate to 32.6 kg/s, which is the design limit. This clearly concludes that these two processes are not the feasible retrofit options.
The HIGHSULF process, which is the base case, achieves the product
H2S purity level of 500 ppm at a steam consumption of 32.6 kg/s.
On the other hand, the proposed method achieves the same at a
steam consumption of 19.6 kg/s. Furthermore, the acid gas CO2
purity of the proposed case is also significantly higher than that of
the base case as delineated in section 4.1. Also, as explained in
section 4.2, this scheme incurs a capital investment of 14.3 million,
which can be recovered within 1.3 years.
5. Conclusion
This study proposed a configuration involving the integration of
the RHVT cooling system in the AGE unit. A process simulation
model of a commercial middle-east-based AGE unit was developed
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
985
Fig. 10. Effect of nitrogen mass flow rate on the monetary benefits for various crude oil price scenarios (All other model parameters except lean solvent temperature are as shown in
Tables 1 and 2).
Fig. 11. Process simulation model of the AGE process with separate feeds.
using the process simulator ProMax 4.0, which uses data-driven
kinetic models to address gas absorption. A parametric analysis
was carried out to quantify the reduction in operational energy as a
function of lean solvent temperature. It was found that reducing the
lean solvent temperature to an optimal temperature of 25 C
resulted in increased purity levels of H2S and CO2 in the two
product streams. Energy balance calculations revealed that a nitrogen flow rate of 864 kg/s is required as feed for the RHVT to
achieve the optimum lean solvent temperature. A 40% reduction in
both steam consumption and annual carbon footprint can be achieved using the proposed scheme. Economic analysis estimated an
annual operating cost savings of 11.2 million USD when using the
integrated system. The computed payback period for the RHVT is
1.3 years. This highlights the potential of the RHVT integrated
system in reducing the operating cost and carbon footprint of the
AGE unit while producing a CO2-rich gas stream with adequate
purity for EOR and sequestration applications. Also, the potential
benefits of the proposed method were compared with those of the
other alternate commercial processes. The proposed scheme
manages to achieve the required specifications of low H2S content
and suitable CO2 gas for EOR at a substantially lower operating cost
and carbon footprint.
986
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
Fig. 12. Process simulation model of the AGE process with two absorbers.
Table 4
Comparison of the proposed method with conventional processes.
Case
Product gas
HIGHSULF - Combined feeds scheme
(Base Case/Current AGE process)
Separate feeds scheme
Multiple absorbers scheme
Base case þ RHVT(Proposed scheme)
a
Steam consumption rate
Annual operating cost savings
Capital cost
H2S, ppm
CO2, %
Acid Gas
H2S, %
CO2, %
kg/s
USD
USD
500
89
45
48.2
32.6
0
0
17263
2362
500
86.1
92.4
97
46.1
32.8
69.3
45.2
58.5
22
32.6
32.6
19.6
0
0
11.2 million
NAa
NAa
14.3 million
NA e Not applicable; Economic analysis not required since these processes do not meet the product specifications.
Nomenclature
A
A’
b
b’
C
CF
CP
Cnaturalgas
CpAir
Cpmix
CpN2
Cpwater
E
E0
Н naturalgas
k
k0
Frequency factor of H2 CO3 dissociation reaction, 1/s
Frequency factor of reaction between [Hþ] and [HCO3],
1/mol. 1/s
Constant quantifying temperature effect of
H2 CO3 dissociation reaction
Constant quantifying temperature effect of reaction
between [Hþ] and [HCO3]
Amount of CO2 emitted per 1 kJ energy of natural gas, kg
Unit conversion factor
Price of crude oil, USD
Cost of processed natural gas per unit mass, USD/kg
Heat capacity of air, kJ/kg. oC
Heat capacity of mixture, kJ/kg. oC
Heat capacity of nitrogen, kJ/kg. oC
Heat capacity of water, kJ/kg. oC
Activation energy of H2 CO3 dissociation reaction, kJ/mol
Activation energy of reaction between [Hþ] and
[HCO3], kJ/mol
Heating value of natural gas, kJ/kg
Rate constant of H2 CO3 dissociation reaction, 1/s
Rate constant of reaction between [Hþ] and [HCO3], 1/
mol. 1/s
mCarbon
Carbon footprint reduction, kg
mnaturalgas Natural gas quantity required for steam generation, kg
Mass of water, kg
mwater
Mass flow rate of air, kg/s
m_ Air
Mass flow rate of mixture, kg/s
m_ mix
Mass flow rate of nitrogen, kg/s
m_ N2
Qnet
Net heat required for steam generation, kJ
Qsolvent cooler Duty across the lean solvent cooler, kJ
Qsteam
Theoretical heat required for steam generation, kJ
R
Gas constant, kJ/mol. oC
r
Rate of H2 CO3 dissociation reaction, mol/m3.s
Rate of reaction between [Hþ] and [HCO3], mol/m3.s
r0
rnet
Net rate of CO2 absorption reaction, mol/m3.s
T
Reaction temperature, oC
TAir in
Air inlet temperature, oC
TCnaturalgas Total cost of processed natural gas, USD
Air-N2 mixture inlet temperature, oC
Tmix in
Tmix out
Air-N2 mixture outlet temperature, oC
TN2 in
N2 inlet temperature, oC
TRef
Reference temperature, oC
TSaturation Saturation temperature of water, oC
½H2 CO3 Concentration of H2 CO3 , mol/m3
[Hþ]
Concentration of Hþ, mol/m3
3
[HCO ] Concentration of HCO3, mol/m3
S. Dara et al. / Journal of Cleaner Production 201 (2018) 974e987
AGE
EOR
MDEA
MQL
NRTL
RHVT
SRK
SRU
Acid Gas Enrichment
Enhanced Oil Recovery
Methyldiethanolamine
Minimum Quality Lubrication
Non Random Two Liquid
RanqueeHilsch Vortex Tube
Soave Redlich Kwong
Sulfur Recovery Unit
Greek Symbols
a
Parameter for estimating the cost of processed natural
gas
b
Parameter for estimating the cost of processed natural
gas
g
Parameter for estimating the cost of processed natural
gas
DTAir
Temperature change of air, oC
DTmix
Temperature change of mixture, oC
DTN2
Temperature change of nitrogen, oC
l
Latent heat of vaporization, kJ/kg
h
Overall process efficiency for steam generation
References
Abbas, Z., Mezher, T., Abu-Zahra, M.R.M., 2013. Evaluation of CO2 purification requirements and the selection of processes for impurities deep removal from the
CO2 product stream. Energy Procedia 37, 2389e2396.
Abu Dhabi National Oil Company, 27th Mar 2016. Personal Communication.
Addington, L., Ness, C., 2009. An evaluation of general “rules of thumb” in amine
sweetening unit design and operation. In: Proceedings of the GPA Europe Sour
Gas Processing Conference, Sitges, Spain.
Bryan Research and Engineering, Inc., Powerful Process Simulation Technology
(ProMax) http://www.bre.com/(Accessed on 14th Jan 2017)
Dara, S., Berrouk, A.S., 2017. Computer-based optimization of acid gas removal unit
using modified CO2 absorption kinetic models. Int. J. Greenhouse Gas Control
59, 172e183.
Eiamsa-ard, S., Promvonge, P., 2008. Review of RanqueeHilsch effects in vortex
987
tubes. Renew. Sustain. Energy Rev. 12, 1822e1842.
Esen, M., Yuksel, T., 2013. Experimental evaluation of using various renewable
energy sources for heating a greenhouse. Energy Build. 65, 340e351.
Esen, H., Inalli, M., Esen, M., 2006. Technoeconomic appraisal of a ground source
heat pump system for a heating season in eastern Turkey. Energy Convers.
Manag. 47, 1281e1297.
Esen, H., Inalli, M., Esen, M., 2007. A techno-economic comparison of groundcoupled and air-coupled heat pump system for space cooling. Build. Environ.
42, 1955e1965.
Jozi
c, S., Baji
c, D., Celent, L., 2015. Application of compressed cold air cooling:
achieving multiple performance characteristics in end milling process. J. Clean.
Prod. 100, 325e332.
Khanmamedov, T.K., 2013. Strategies for acid gas enrichment and tail gas treatment.
In: Proceedings of AIChE Spring Meeting and Global Congress on Process Safety,
Houston, Texas, USA.
Kidnay, A.J., Parrish, W.R., 2006. Fundamentals of Natural Gas Processing. Taylor &
Francis Group, Boca Raton.
Kohl, A.L., Nielson, R.B., 1997. Gas Purification, fifth ed. Gulf Publishing Company,
Houston.
Lunsford, K., McIntyre, G., 1999. Decreasing contactor temperature could increase
performance. In: Proceedings of the Seventy-eighth GPA Annual Convention.
Nashville, Tennessee, USA, pp. 121e127.
Pandey, M., 2005. Process optimization in gas sweetening unit-a case study. In:
Proceedings of the International Petroleum Technology Conference, Doha,
Qatar.
Parks, L.E., Perry, D., Fedich, R., 2010. FLEXSORB ®SE a proven reliable acid gas
enrichment solvent. In: Proceedings of the 2nd Annual Gas Processing Symposium, vol. 2, pp. 229e235. Amsterdam, Netherlands.
Peters, M.S., Timmerhaus, K.D., West, R.E., 2003. Plant Design and Economics for
Chemical Engineers, fifth ed. McGraw-Hill higher education, New York.
Saberi, A., Rahimi, A.R., Parsa, H., Ashrafijou, M., Rabiei, F., 2016. Improvement of
surface grinding process performance of CK45 soft steel by minimum quantity
lubrication (MQL) technique using compressed cold air jet from vortex tube.
J. Clean. Prod. 131, 728e738.
Soave, G., 1972. Equilibrium constants from a modified Redlich-Kwong equation of
state. Chem. Eng. Sci. 27, 1197e1203.
Tunkel, L., Universal Vortex Inc, 10th Aug, 2017. Personal Communication.
Universal Vortex, Inc., Designer of industrial vortex tube, http://www.universalvortex.com/(Accessed on 16th Jul 2017)
U.S. Energy Information Administration, https://www.eia.gov/environment/
emissions/co2_vol_mass.php (Accessed on 9th Aug 2017)
Weiland, R.H., Khanmamedov, T.K., 2010. Acid gas enrichment flow sheet selection.
Oil Gas J.
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