Fuel 266 (2020) 117111
Contents lists available at ScienceDirect
Fuel
journal homepage: www.elsevier.com/locate/fuel
Full Length Article
Techno-economic assessment of process integration models for boosting
hydrogen production potential from coal and natural gas feedstocks
T
Usman Hamida,b, Ali Raufb, , Usama Ahmedc, , Md. Selim Arif Sher Shahd, Nabeel Ahmade
⁎
⁎
a
Department of Chemical Engineering, Texas Tech University, Lubbock, TX 79409-3121, USA
Department of Chemistry and Chemical Engineering, Lahore University of Management Sciences, Lahore 54792, Pakistan
c
Chemical Engineering Department, King Fahad University of Petroleum & Minerals, Dhahran, Saudi Arabia
d
Department of Chemical and Biomolecular Engineering, Yonsei University, Republic of Korea
e
Department of Chemical Engineering, COMSATS University Islamabad, Lahore Campus, Pakistan
b
G R A PHICA L A BSTR A CT
A R TICL E INFO
A BSTR A CT
Keywords:
Gasification
Steam methane reforming
Heat integration
H2 production
CO2 emissions
The elevated energy demands from past decades has created the energy gaps which can mainly be fulfilled
through the consumption of natural fossil fuels but at the expense of increased greenhouse gas emissions.
Therefore, the need of clean and sustainable options to meet energy gaps have increased significantly.
Gasification and steam methane reforming are the efficient technologies which resourcefully produce the syngas
and hydrogen from coal and natural gas, respectively. The syngas and hydrogen can be further utilized to
generate power or other Fischer Tropsch chemicals. In this study, two process models are developed and
technically compared to analyze the production capacity of syngas and hydrogen. First model is developed based
on conventional entrained flow gasification process which is validated with data provided by DOE followed by
its integration with the reforming process that leads to the second model. The integrated gasification and reforming process model is developed to maximize the hydrogen production while reducing the overall carbon
dioxide emissions. Furthermore, the integrated model eradicates the possibility of reformer’s catalyst deactivation due to significant amount of H2S present in the coal derived syngas. It has been seen from results that
updated model offers 37% increase in H2/CO ratio, 10% increase in cold gas efficiency (CGE), 25% increase in
overall H2 production, and 13% reduction in CO2 emission per unit amount of hydrogen production compared to
Abbreviations: GEE, GE Energy is a type of entrained flow gasifier; SMR, Steam methane reforming; DMR, Dry methane reforming; POX, Partial oxidation; WGS,
Water gas shift; HHV, Higher heating value; LHV, Lower heating value; CGE, Cold gas efficiency; IGCC, Integrated gasification combined cycle; IPCC,
Intergovernmental Panel on Climate Change; EIA, Energy Information Administration; NG, Natural gas; MFK, Mass flowrate in kg/hr; TIC, Total investment cost;
HCR, Hydrogen to carbon monoxide ratio
⁎
Corresponding authors.
E-mail addresses: ali.rauf@lums.edu.pk (A. Rauf), usama.ahmed@kfupm.edu.sa (U. Ahmed).
https://doi.org/10.1016/j.fuel.2020.117111
Received 1 August 2019; Received in revised form 11 January 2020; Accepted 14 January 2020
0016-2361/ © 2020 Elsevier Ltd. All rights reserved.
Fuel 266 (2020) 117111
U. Hamid, et al.
base case model. Furthermore, economic analysis indicated 8% reduction in cost for case 2 while presenting 7%
enhanced hydrogen contents.
1. Introduction
the syngas. The major drawbacks of this process involve endothermic reactions that increases both operational and production costs compared to
SMR process. Moreover, the H2/CO ratio is 1:1 that indicates less H2 production with respect to SMR [18]. Furthermore, DMR process causes catalyst deactivation that results in operational problem [19]. Similarly, partial
oxidation (POX) is a noncatalytic process which upon combustion with pure
oxygen produce syngas. The main drawbacks are high oxygen separation
cost, high operating temperature i.e. 1300–1500℃, and pressure of 3–8 MPa
to proceed reactions [10]. In addition, syngas produced from the POX
process offer H2/CO ratio of 1:1 which is much less than the ratio offered by
SMR [10,20]. Conclusively, considering prime advantages of gasification
and SMR over DMR and POX, both gasification and SMR are best suited to
maximize the H2 production while minimizing the greenhouse gas emissions.
Most of attentions are devoted towards the standalone gasification
development with and without CO2 capture assessment using coal and
other feedstocks to produce H2, electricity, poly-generation, and syngas
[21–24]. Cormos [25] developed the overall integrated gasification
combined cycle (IGCC) with CO2 capture and storage. IGCC processes
have also shown promising integration with other processes to produce
H2 through various gasifier types i.e. Shell, GEE, etc. upon utilization of
coal and biomass [26–29]. In addition, Cormos et al. [14] presented
multi-fuel and multi-product operation for IGCC with carbon capture
and storage. On the other hand, economic analysis was also being
performed for H2 production through natural gas reforming with CO2
capture [30]. Szima et al. studied the DMR process for H2 production
along with CO2 capture [31]. Qian et al. [32] presented an integrated
model for efficient utilization of coke-oven gas and coal through gasification using tri-reforming. Yi et al. [33] studied the integration of
reforming and coal gasification for CO2 utilization to minimize the
water resources. Conclusively, most of the research conducted so far
dealt either standalone reforming process or coal gasification.
However, a few studies have been carried out recently to investigate
the harmonious process integration and intensification between gasification and SMR. Adams and Barton [34] proposed techno-economic
studies for several processes to produce efficient polygeneration from
coal and natural gas. Cormos et al. [35] assessed the techno-economic
calculation for flexible H2 and power generation through coal gasification. Ahmed et al. [36] suggested the integrated model of coal gasification and SMR to utilize syngas heat and steam from gasification
process to increase H2 production. Ahmed et al. [37] also performed
techno-economic assessment for both conventional coal and integrated
coal gasification and SMR process, concluding the integrated model as a
more efficient process. However, the earlier process integrations of
gasification and reforming units in series experienced operational issues
in the reforming section as the coal derived syngas contains higher
sulfur content then the allowable limits of the reforming catalyst.
Keeping in view the limitations in the harmonious integration between gasification and reforming technologies, new pragmatic integration approach has been demonstrated in this work to indirectly
utilize the heat from the gasification derived syngas in the SMR process
with the aid of radiant cooler without any risk of catalyst poisoning and
deactivation. The proposed integration exploits the parallel process
integration network approach to utilize the key technical and economic
benefits of different technologies that can enhance the overall process
sustainability without relying on the single fuel. The current article
consists of four sections. The first section provide a through introduction followed by the process simulation methodology containing base
case model (case 1) and the updated model (case 2). In subsequent
section, a detailed analysis based on critical process variables against
Since the last few decades, abrupt increase in population, industrial
development, and rising living standards resulted in an enormous escalation in the energy demands. According to BP energy outlook 2019,
the overall energy demands across the globe is expected to grow by
1.2% per annum [1]. Consequently, creating a pressing problem to
explore new technologies and energy resources to overcome the energy
gap. Although, the avid consumption of natural resources meet energy
gap but at the expense of rising level of greenhouse gases especially
CO2, which is a prime root cause of global warming [2]. For instance,
the overall growth rate of CO2 emission is 2.1% after three years of
stability from 2014 to 2016, leading to temperature rise of 2℃ [3,4].
According to Intergovernmental Panel on Climate Change (IPCC), it is
expected that global temperature may rise from 1.1 to 6.4℃ by 2100
[5]. Therefore, it is need of time to explore such options that efficiently
utilize the energy resources while minimizing the greenhouse gas.
Presently, the primary fossil fuels available to generate power and
overcome energy gap includes natural gas, crude oil, and coal [3]. Natural
gas yields high efficiencies for power plants with low greenhouse gas
emissions and showed a potential of growth as compared to oil [6]. According to U.S. Energy Information Administration (EIA), the natural gas
prices for the projected period till 2050 are anticipated to remain low [7,8].
Therefore, it is expected to gain priority in future due to cost effectiveness
and efficient production of H2 with minimum greenhouse gas emissions.
Importantly, advancements through environmentally benign coal and natural gas technologies coupled with carbon capture and sequestration (CCS)
will gain priority and will open doors for coal utilization [9,10]. On the
other hand, natural resources survey conducted in 2013 showed an approximated 890 billion tonnes of coal reserves available across the globe
[11]. Furthermore, consumption of coal indicates increasing trend from
2012 to 2040 with an average increase of 0.6% per annum [12]. Considering the fossil fuel reserves, coal reserves are much more than that of
natural gas. Therefore, the abundantly available coal reserves across the
globe at relatively low cost with prime applications in power and hydrogen
(H2) production, coal is expected to remain potential fossil fuel in future
[12]. Due to ever increasing energy demands and economic instability
around the globe, it is unlikely to meet demands while solely relying on
single fuel. Therefore, the need of hybrid systems to handle multiple fuels
generating multiple products in environmentally benign manner will enhance the overall process sustainability.
Recent developments in H2 and electricity generation processes using
coal and natural gas as primary fuels have devoted more attentions to meet
energy requirements while mitigating greenhouse gases emissions [13].
Some of the commercial H2 production technologies from fossil fuels includes coal gasification, steam methane reforming (SMR), dry methane reforming (DMR), and partial oxidation (POX) [10]. The gasification of low
and high-quality fuels in both standalone and blending manner have been
evaluated as one of the cost-effective approaches for future power and H2
generation technology with low carbon footprint. Importantly, the diversity
of products that can be produced from the gasification mainly H2 make
gasification technology even more prominent [14,15]. Moreover, the flexibility of gasification process for integration with other technologies signifies its future potential. SMR process uses natural gas and steam to generate syngas containing H2 and carbon monoxide with ratio of 3:1 that
indicate higher H2 production potential. It is among the most developed,
widespread, and least expensive technology for H2 production [16]. Moreover, the high efficiencies, low operational, and production cost along with
readily available natural gas strengthen this technology [17]. On the other
hand, dry methane reforming (DMR) uses CO2 and natural gas to produce
2
Fuel 266 (2020) 117111
U. Hamid, et al.
product composition, cold gas efficiency (CGE), H2/CO (HCR) ratio,
higher heating value (HHV), and CO2 emissions has been discussed. The
last section concludes the article along with some future directions.
this point, the outlet of water gas shift process contains high contents of
water, which upon temperature reduction removed as a condensate
with the help of flash drum in gas cooling. Subsequently, acid gas removal section is installed which separate CO2 from H2. The validation
[43] of gasification and WGS models of the base case with the literature
has been done and represented in the Fig. 2.
2. Simulation methodology and model development
Aspen Plus V9 is used in this study as a simulation tool where the
Peng-Robinson with Boston-Mathias alpha function (PR-BM) is used as
an effective thermodynamic property package for gasification,
scrubber, steam methane reforming (SMR), acid gas removal, and water
gas shift sections [38–40]. Due to varying composition of coal from one
region to another, proximate analysis, ultimate analysis, and heating
value data is provided in Table 1. Prior to perform the simulation studies and convert the coal into syngas, RYield model of Aspen Plus is
used to breakdown coal into its constituent elements. Subsequently,
RGibbs reactor is used to model gasification process which results in
syngas production and slag formation. Similarly, RGibbs model is also
used to model the SMR process. The composition [36,42] of the natural
gas used in the process is provided in Table 2. Concisely, RGibbs model
of Aspen Plus works on the principle of Gibbs free energy minimization
to calculate the equilibrium with phase splitting. Likewise, REquil
model of Aspen Plus is used to calculate the equilibrium of phases when
stoichiometry of the reaction is known. Therefore, water gas shift reactors are modeled based on REquil reactors. Some of the design assumptions and process streams data is presented in Tables 3 and 4.
Before investigation of integration in the case 2, the base case model
(case 1) is validated with the data published in the literature.
2.2. Case 2 – Integration of GEE radiant gasifier with steam methane
reforming process
In case 2, gasification and steam methane reforming (SMR) processes are coupled using parallel design configuration as shown in
Fig. 3. Natural gas [36,42] is fed to the reforming unit where reforming
reactions mentioned in Table 5 occurs to produce syngas at ~912℃ and
32 bar. In the reforming process, heavier hydrocarbons are first converted into methane, which subsequently undergoes methane reforming
in the presence of steam. The developed SMR model is also validated
with the published data [36] as represented in Fig. 4.
Case 2 presents the heat integration of SMR section with the GEE
radiant type gasifier. In GEE radiant type gasifier, significant heat can
be extracted from the syngas leaving the gasifier that can efficiently be
utilized to proceed SMR process. To make this heat integration possible,
heat exchangers (HEX301 & MHT) are installed in such a way that heat
extracted from the syngas is used for pre-heating and sustaining the
endothermic reactions involved in SMR process. Subsequently, heavier
hydrocarbon is converted into methane and eventually to syngas at
912°C in RS701 and RG701, respectively. The detailed flowsheet built
using Aspen Plus is presented in Fig. 5. Due to high temperature of
resulting syngas from reforming process, the temperature is reduced to
325°C by water quenching followed by the water gas shift section. The
flowsheet of case 2 for Scrubber, water gas shift, and cool gas section
remain same as depicted for case 1. A selected comparison of process
streams data of case 1 (Fig. 1) and case 2 (Fig. 3) is presented in Table 4.
2.1. Case 1 (Base Case) – Syngas production from GEE entrained flow
gasifier
The complete process is divided into multiple sections i.e. gasification, scrubber, water gas shift, and gas cooling section. The schematic
diagram of case 1 is presented in Fig. 1. The base case of gasification
section, scrubber, water gas shift, gas cooling section, and steam methane reforming section has been validated [36,43]. The gasifier section
is the most important section in any gasification process where syngas is
produced from the coal at high pressure and temperature in the presence of water and oxygen. After decomposition of coal, resulting
product is being mixed with oxygen and slurry water. Subsequently,
this mixture undergoes gasification process where reactions mentioned
in Table 5 resulted in syngas generation and slag leaving at 1370°C. The
GE Energy (GEE) radiant entrained flow gasifier is considered in this
work due to which heat is recovered from syngas to decrease the
temperature from 1370°C to 760°C. Afterward, the quenching is carried
out with the help of liquid water to decrease the temperature to 216°C
follow by slag removal.
After gasification, scrubber is used to remove particulate matters
and chlorides through water to avoid fouling in the subsequent unit
operations. The gas product leave from the top while waste leave from
the bottom of scrubber at 211° C . The syngas is then fed to the water gas
shift (WGS) reactor to convert carbon monoxide into CO2 in the presence of steam to maximize the H2 production. Also, carbonyl sulfide
(COS) in the syngas is difficult to remove as compared to hydrogen
sulfide (H2 S ). Therefore, in WGS section COS is converted into H2 S in
the presence of steam resulting in H2 and CO2 . Steam to carbon monoxide ratio of 2.0 is maintained at the inlet of water gas shift reactor to
maximize the CO conversion. First water gas shift reactor (SHFT501)
decreases the CO mole fraction from 24.5% to 5% while increases the
H2 mole fraction from 18.8% to 38.2%. An intermediate cooler
(HEX501) is used to remove the heat generated due to the exothermic
reactions before introducing to the second water gas shift reactor.
Ultimately, CO mole fraction reduced to 0.8% while H2 mole fraction
become as high as 42.4% and leave the shift reactor (SHFT502) at258°C.
Subsequently, heat is being recovered from the product stream containing 0.8% CO, 29.9%CO2 , 42.4%H2 , 25.3%H2 O, and 1.8 ppm COS. At
3. Results and discussion
3.1. Heat integration analysis
Two process models have been developed and compared in this
study to analyze the syngas production and the heating requirements.
Case 1 is the conventional process for generating syngas from coal
Table 1
Proximate and ultimate analysis of coal [41].
Coal Type
Proximate analysis (weight %)
Moisture
Ash
Volatile Matter
Fixed Carbon
Total
Sulfur
HHV, kJ/kg (Btu/lb)
LHV, Btu/lb (Btu/lb)
Bituminous Illinois No. 6
As Received
Dry Basis
11.12
9.7
34.99
44.19
100
2.51
27,113 (11,666)
26,151 (11,252)
0
10.91
39.37
49.72
100
2.82
30,506 (13,126)
29,544 (12,712)
As Received
11.12
63.75
4.5
1.25
0.29
2.51
9.7
6.88
100
Dry Basis
0
71.72
5.06
1.41
0.33
2.82
10.91
7.75
100
Ultimate Analysis (weight %)
Moisture
Carbon
Hydrogen
Nitrogen
Chlorine
Sulfur
Ash
Oxygen
Total
3
Fuel 266 (2020) 117111
U. Hamid, et al.
gasification process, whereas, case 2 represents the integration of gasification and reforming technologies for enhanced syngas generation.
The process flow diagram developed in Aspen Plus for case 2 showing
the heat exchanger network for recovering the heat from coal derived
syngas and supplying it for NG reforming is represented in Fig. 6. The
syngas generated from gasification unit is typically at a temperature of
1370℃ which is being cooled enough to carry out WGS reactions.
Therefore, two heat exchangers are installed in the downstream of gasification section to reduce the temperature from 1370℃ to 760℃. The
outlet temperature of first heat exchanger is adjusted in such a way that
the extracted heat from the coal derived syngas can efficiently be utilized in the reformer. The methane and water that will be utilized in the
SMR section entered at 25℃ which undergoes pre-heating to meet reformer temperature requirement. The calculator blocks are used to
control the flowrates of both methane and water in such a way that all
heat recovered from the syngas of gasifier efficiently be utilized in feed
pre-heating and sustaining reformer temperature.
For efficient heat recovery between hot and cold streams, Aspen
Energy Analyzer has been used to perform the pinch analysis and
construct grid and composite diagrams. Thermal data of the process
streams has been extracted through the simulation models followed by
approach temperature difference (ΔTmin) of 10℃ between hot and cold
streams. Moreover, the composite curves are developed for both the
cases to analyze the heat recovery between hot and cold utilities. The
composite diagram depicting the enthalpy vs temperature graphs are
presented for case 1 and case 2 in Figs. 7 and 8, respectively.
The comparative analysis showed that, an energy saving of
156.8 MW is possible in case 2 design through the heat integration
which otherwise require significant flowrates of external utilities to
drive steam methane reforming. Subsequently, the syngas from reformer leaves at 912℃ which is at much higher temperature and require cooling before injecting to water gas shift section. Therefore, two
heat exchangers are installed to pre-heat the water to meet the
quenching water requirement and steam required in water gas shift
section. This heat integration resulted in reducing the utilities load of
the overall process to 47 MW.
Table 2
Natural gas composition [36,42].
Natural gas composition (Mole Fraction)
CH4
C2H6
C3H8
C4H10
CO2
N2
Total
Lower heating value (LHV)
0.939
0.032
0.007
0.004
0.010
0.008
1.000
47.76 MJ/kg
Table 3
Design assumptions.
Equipment
Aspen model
Description
Coal flowrate
Mixer
Gasifier (GEE)
RYield, RGibbs
(Reactor)
Pre-reformer
RStoic (Reactor)
Reformer
RGibbs (Reactor)
Air Separation Unit
(ASU)
Separator,
Multistage
compressor
Water Gas Shift
(WGS)
Gas Cooling
REquil (Reactor),
Heat exchangers
Flash drum
Coal = 62.20 kg/sec (70% coal, 30%
water)
Entrained flow gasifier GE Energy
(GEE) Temperature = 1370℃
Pressure = 5.6 MPa;
Carbon conversion = 98%
Hydrocracking of heavier
hydrocarbon
Natural gas flowrate = 5.5 kg/sec;
Temperature = 912℃;
Pressure = 3.2 MPa;
Steam/CH4 ratio = 3.0;
Nickel based catalyst
Oxygen flowrate = 50.57 kg/sec;
Temperature = 90.9℃;
Pressure = 6.76 MPa;
94.97% pure oxygen (mole %)
Adiabatic reactors = 2;
Steam/CO ratio = 2.0
Isothermal flash vessels = 3
Table 4
Comparison of Case 1 and Case 2 process streams data.
Air
Temperature [°C]
Pressure [MPa]
Mass Flows [kg/hr]
CO
CO2
H2
H2O
CH4
H2S
N2
COS
O2
Water
Case 1
Case 2
Case 1
Case 2
Case 1
Case 2
Case 1
Case 2
15.0
0.1
794,515
15.0
0.1
794,515
90.9
6.8
182,048
90.9
6.8
182,048
60.0
6.0
93,188
60.0
6.0
93,188
0.0003
0.0003
0.0099
0.0099
1.0000
1.0000
0.7732
0.7732
216.0
5.5
662,874
0.2724
0.0695
0.2087
0.4323
0.0001
0.0049
0.0061
0.0002
205.6
3.2
663,234
0.2722
0.0694
0.2086
0.4326
0.0001
0.0049
0.0061
0.0002
0.0188
0.0188
Scrubber
Temperature [°C]
Pressure [MPa]
Mass Flows [kg/hr]
CO
CO2
H2
H2O
CH4
H2S
N2
COS
Syngas
WGS
Gas Cooling
Case 1
Case 2
Case 1
Case 2
Case 1
Case 2
211.0
5.5
619,159
0.2931
0.0747
0.2245
0.3899
0.0001
0.0053
0.0065
0.0002
189.4
3.2
638,507
0.2835
0.0722
0.2172
0.4098
0.0001
0.0051
0.0063
0.0002
150.0
5.4
730,977
0.0083
0.2993
0.4243
0.2531
0.0001
0.0046
0.0054
150.0
3.2
807,548
0.0079
0.2790
0.4547
0.2418
0.0037
0.0040
0.0049
39.0
5.4
559,018
0.0111
0.3999
0.5673
0.0017
0.0001
0.0061
0.0073
39.0
3.2
616,163
0.0103
0.3670
0.5984
0.0024
0.0049
0.0052
0.0064
4
Fuel 266 (2020) 117111
U. Hamid, et al.
Fig. 1. Case 1. Syngas production from coal-based gasification process.
Table 5
Chemical reactions involved in the process.
Gasification Reactor
C(s) + H2 O CO + H2
C(s) + CO2
2CO
CO
H= +131MJ/kmol
H = + 172 MJ/kmol
H=
111MJ/kmol
CO2
H=
283MJ/kmol
H2 O
H=
242MJ/kmol
Water gas shift (WGS) reactor
COS + H2 O H2 S+ CO2
CO + H2 O H2 + CO2
H=
H=
34MJ/kmol
41MJ/kmol
1
2
1
CO + O2
2
1
H2 + O2
2
C(s) + O2
Steam methane reforming reactor
3 C2 H6 + H2 O 5CH 4 + CO
3C3 H8 + 2H2 O 7CH 4 + 2CO
3C4 H10 + 3H2 O 9CH 4 + 3CO
CH 4 + 2O2
CO2 + 2H2 O
CH4 + H2 O CO + 3H2
ΔH
ΔH
ΔH
ΔH
ΔH
=
=
=
=
=
+3.6460
+16.607
+41.116
−802.54
+206.12
MJ/kmol
MJ/kmol
MJ/kmol
MJ/kmol
MJ/kmol
Fig. 4. Base case validation of steam methane reforming section [36].
Base Case Validation
CO
CO2
H2
H2O
0.45
0.40
CH4
3.2. Effect of slurry temperature and oxidant requirements on syngas
composition
0.4244
0.4212
0.3931
0.3893
Coal slurry and moisture is one of the main constituents of gasification process and the temperature of slurry also plays an important
role in the syngas composition and cold gas efficiency (CGE). Coal
based power plants usually have large amount of waste heat which can
be utilized for coal slurry pre-heating. In this study, the sensitivity
analysis has been carried out at five different temperature i.e. 25℃,
60℃, 250℃, 278℃, 300℃ to evaluate its effect on syngas composition.
The results presented in the Fig. 9 reveal the mole fraction, O2/Coal
ratio, and CGE at five different temperature of slurry water. The results
showed that increasing coal slurry temperature increases the CO and H2
mole fraction while decreases the steam and CO2 content in the syngas.
CGE is an important parameter to analyze the performance of the gasifier and it typically represents the heating value of the syngas as given
in Eq. (1). The results showed the CGE of the syngas increase with an
increase in the temperature of coal slurry.
Mole Fraction
0.35
0.30
0.3026
0.3012
0.2942
0.2993
0.2531
0.2526
0.25
0.1862
0.20
0.1807
0.15
0.1051
0.1002
0.10
0.05
0.00
0.0002
Reference
0.0002
Base Case
Gasifer Product
0.0082
0.0001
Reference
0.0083
0.0001
Base Case
WGS Product
Fig. 2. Base case validation [43].
Fig. 3. Case 2: Heat integrated gasification and steam methane reforming.
5
Fuel 266 (2020) 117111
U. Hamid, et al.
RY301
COAL(IN)
H301
CH4(IN)
PUMP301
S311
H2O(IN)
S301
MIX301
OXY102(IN)
RG301
S302
HEX301
HOTGAS
VLV301
HEX302
MHT
S304
S303
S312
SPL302
SNG101(OUT)
S313
SSPLIT
WATER(IN)
SLAG
S310
S701
S702
MIX701
RS701
RG701
S703
S704
S705
SNG103(OUT)
H701
HEX702
S305
S306
HEX703
S307
S308
HEX704
S309
S101(OUT)
SPL301
S706
Fig. 5. Integrated gasification and steam methane reforming flowsheet using Aspen Plus®.
Fig. 6. Heat integration between Gasification and SMR for Case 2 Design.
CGE =
msyngas × LHVsyngas
m fuel × LHVfuel
coal slurry temperature as shown in Fig. 9. As we know that O2 production is an energy intensive process and it affects the overall process
performance. The analysis showed that the O2/coal ratio can significantly be reduced at elevated coal slurry temperature. Moreover, the
results indicate that with an increase in slurry water temperature, the
overall load of gasifier to increase the syngas temperature to 1370°C
(1)
Moreover, the effect of coal water slurry’s temperature has been also
observed on the consumption of oxidizing agent in the gasifier. The
results showed that oxygen consumption decreases with an increase in
Fig. 7. Composite curve of case 1.
6
Fuel 266 (2020) 117111
U. Hamid, et al.
Fig. 8. Composite curve of case 2.
Fig. 9. Effect on slurry water temperature on gasifier syngas and CGE.
Fig. 11. Effect of steam to carbon monoxide ratio on WGS product.
Fig. 10. Effect of reformer temperature on syngas composition and methane
conversion.
7
Fuel 266 (2020) 117111
U. Hamid, et al.
Table 8
Estimation of operational and maintenance expenditures.
Fig. 12. Case comparison at the outlet of water gas shift reactor.
O&M Cost
Units
Case 1
Case 2
Fixed O&M Cost
Maintenance Cost
Labor Cost
Administrative, support & overhead cost
Total Fixed O&M Cost
MM
MM
MM
MM
€/year
€/year
€/year
€/year
24.70
6.00
1.80
32.50
25.20
6.00
1.80
33.00
Variable O&M Cost
Natural Gas
Coal
Boiler Feed Water (BFW)
WGS Catalyst
Solvent (Selexol)
Reforming Catalyst
Total variable O&M Cost
MM
MM
MM
MM
MM
MM
MM
€/year
€/year
€/year
€/year
€/year
€/year
€/year
–
48.64
0.48
1.36
0.80
–
51.29
15.94
48.64
1.15
1.38
0.94
0.02
68.06
Total Fixed and Variable Cost
Total Fixed and Variable Cost (net H2)
Total Lifetime Cost
Lifetime Hydrogen Cost
MM €/year
MM €/tonne H2
MM €
€/tonne H2
83.79
2.59
2941.55
484.71
101.06
2.49
3390.52
446.14
Table 6
Case comparison against selected parameters.
Parameters
Case 1
Case 2
CO
CO2
H2
H2O
N2
H2S
Others
Higher Heating Value (HHV) [MJ/kmol]
Lower Heating Value (LHV) [MJ/kg]
Hydrogen Flowrate [tonne/hr]
Hydrogen/Carbon Monoxide (HCR)
CO2/H2 (CO2 emissions per unit H2 production)
Cold Gas Efficiency (CGE)
0.0083
0.2993
0.4244
0.2531
0.0054
0.0046
0.0049
137.77
7.13
32.37
51.10
0.71
67.0%
0.0079
0.2790
0.4547
0.2418
0.0049
0.0040
0.0077
148.65
8.14
40.53
57.89
0.61
73.7%
Table 9
Results summary.
CO2/H2 (CO2 emissions per unit H2)
Energy saving [MWth]
Hydrogen production [tonne/hr]
Cold gas efficiency (CGE)
Higher heating value (HHV) [MJ/kmol]
Total Investment Cost [MM €]
Hydrogen Cost [€/tonne H2]
0.7
CO
H2
CO2
Case 1
Case 2
0.71
–
32.37
67.0%
137.77
846.88
484.71
0.61
47
40.53
73.7%
148.65
863.96
446.14
H2 O
O2 flowrate
220
0.6
Units
Case 1 (MM
€)
Case 2 (MM
€)
Reformer
Solid Handling Facility
Gasification Island
Syngas processing unit
Acid Gas Removal Unit
Air Separation Unit
Offsite Unit and Utilities
MMCUFTD
tonne of coal/hr
tonne of coal/hr
tonne of coal/hr
tonne of CO2/hr
tonne of O2/hr
Equipment Cost
(25%)
–
53.25
213.85
53.00
118.51
125.98
141.15
5.13
53.25
213.85
53.00
124.77
125.98
143.99
Total Installed Cost
Contingency
MM €
Installed Cost
(15%)
Installed Cost (5%)
705.73
105.86
719.96
107.99
35.29
36.00
MM €
MM €/tonne H2
846.88
26.17
863.96
21.32
Land Cost
Total Investment Cost
Total Investment Cost per
tonne H2
Gasifier mole fraction
Plant sub-system
0.5
180
0.4
0.3
160
0.2
O2 flowrate (Tonne/hr)
200
Table 7
Estimation and comparison of capital cost expenditures.
140
0.1
0.0
120
10
40
70
100
130
160
190
Water flowrate (Tonne/hr)
Fig. 13. Influence of water flowrate on gasifier product.
the gasification process unlike case 1. The typical pressure to carry out
SMR process is around 32 bar. The second most influential parameter is
the reformer’s temperature, which effects the syngas composition, CH4
conversion, and H2/CO (HCR) ratio. As depicted in Fig. 10, an increase
in reformer’s temperature increases H2 and CO contents while decrease
the CO2 and H2O. As SMR reactions are endothermic, the most important parameter to evaluate the reformer performance is the CH4
conversion to syngas which usually increases with an increase in the
reformer temperature. For instance, at 700℃, the syngas has 12.6%
methane that decreases gradually to 0.16% at 1100℃. Moreover,
carbon monoxide mole fraction increases from 2.4% to 13.2% when
temperature increases from 700℃ to 1100℃. Similarly, H2 mole fraction in product stream increases from 30.4% to 52.9% with increase in
temperature from 700℃ to 1100℃. The reason is that reforming
decreases. The results also showed that the increase in slurry’s temperature minimizes the conversion of CO and H2 to CO2 and H2O, respectively. Conclusively, pre-heating of coal slurry water not only increases the CGE of the gasification process but also decreases the
oxygen requirement which ultimately reduces the load on air separation unit.
3.3. Effect of Reformer’s operational parameters on NG derived syngas
In case 2, steam methane reforming (SMR) process is coupled with
8
Fuel 266 (2020) 117111
U. Hamid, et al.
560
H 2 production
CO 2 emission
41.0
550
40.5
540
40.0
39.5
530
syngas at temperature around 912℃. Fine-tuning in the reformer’s
temperature and recycle blocks are used to ensure maximum conversion of methane.
CO2 emission (Tonne/hr)
H2 production (Tonne/hr)
41.5
3.4. Comparison of steam demand requirements for case 1 and case 2
Water gas shift process in the promising technology to convert
carbon monoxide into CO2 in the presence of steam to maximize the H2
production. The CGE of the syngas at the inlet of WGS reactor affects
the overall H2 production and the process steam requirements. It can be
seen form the process flow diagram of case 2 that the coal and NG
derived syngas is mixed at the inlet of WGS reactors which significantly
increase the CGE of the case 2. For instance, the CGE of case 1 and case
2 is calculated as 67.0% and 73.7%, respectively, resulting 10% increase in CGE. The syngas composition analysis also showed that H2/
CO for the case 2 is 37% higher as compared to case 1 before WGS
because of an additional NG reforming. Similarly, the ratio of steam to
CO ratio plays an important role in the conversion of carbon monoxide
to CO2 which also effect the H2 mole fraction in the product. The ratio
selected for this comparison ranges from 0.5 to 2.0 with step size of 0.5.
Fig. 11 shows that with an increase in steam to CO ratio, the conversion
of CO to CO2 and H2 increases. To ensure maximum conversion of CO,
the steam/CO ratio of 2.0 is maintained to produce maximum H2. The
comparative analysis at the outlet of WGS reactor showed that case 2
increases the H2 mole fraction in the syngas up to 7% whereas reduce
the CO2 and H2O content.
Lastly, comparison of case 1 and case 2 at the WGS outlet has been
presented in the Fig. 12. The results indicated that case 2 provide enhanced production of H2 and 7% higher H2 contents as compared to
case 1. Most importantly, the overall CO2 content for case 2 is 6.8% less
than the case 1 which is beneficial for acid gas removal section because
it requires less solvent and energy for hydrogen purification. Also, cold
gas efficiency (CGE) of both cases reveal that case 2 present higher CGE
which hint the enhance overall efficiency of the process.
39.0
38.5
0.60
0.65
0.70
0.75
0.80
0.85
0.90
520
0.95
CH4 for Reforming (MMCUFTD)
Fig. 14. Comparison for CO2 emissions and H2 production.
102.0
890
CAPEX
OPEX
101.8
885
101.6
880
101.2
101.0
870
100.8
865
100.6
OPEX (MM €/year)
CAPEX (MM €)
101.4
875
860
100.4
855
100.2
100.0
850
20%
25%
30%
35%
40%
45%
3.5. CO2 specific emissions analysis for H2 generation
Water Content
Fig. 15. CAPEX and OPEX comparison against slurry water content.
O2 requirement
OPEX/H2 production
2.65
210
O2 requirement (Tonne/hr)
2.60
200
2.55
190
180
2.50
170
2.45
160
150
OPEX / H2 production (MM €/Tonne H 2)
220
Considering the clean and sustainable process for H2 production a
comparison against conventional and integrated model is presented.
Though, H2 is the main constituent of the syngas and water gas shift
product but CO2 is also inevitable during H2 production. In addition,
CO2 is a greenhouse gas that results in global warming and many other
environmental problems. Therefore, processes which produces less
amount of CO2 upon production of H2 gaining priority. Considering the
CO2 emissions per unit H2 production parameter, the results of both
cases have been presented. According to Table 6, case 2 showed higher
potential of H2 production and less CO2 specific emissions compared to
the conventional gasification process. For instance, CO2:H2 ratio for
case 1 is 0.71 while for case 2 the ratio is 0.61 which indicates that case
1 causes more CO2 emission as compared to case 2. Considering the
energy and environment protection indicators, case 2 has been not only
evaluated as a better alternative in terms of H2 production and energy
utilization but it also provides promising results towards the environmental aspects. The model developed in case 1, hydrogen purity after
acid gas removal is 91.33% but case 2 it is 94.8%. Afterward, acid gas
removal section is used to separate CO2 from the H2 upon using the
Selexol as solvent. Precisely, comparison against various parameters
indicated remarkable improvements for case 2 as compared to case 1.
For example, 25% increase in hydrogen production, 14% enhanced
lower heating value (LHV), 10% improved cold gas efficiency (CGE),
and 13% enhanced hydrogen to carbon ratio with 13% reduced carbon
dioxide emissions make case 2 significant improved process.
2.40
20%
25%
30%
35%
40%
45%
Water Content
Fig. 16. Comparison for O2 requirement and OPEX/H2 production cost.
reactions are endothermic and require large amount of heat i.e.
206.12 MJ/kmol. According to Le Chatelier’s principle, increase in
temperature promote forward rate of reaction which confirm that more
and more methane conversion is possible at high temperature. Importantly, more than 85% of methane is already being converted into
9
Fuel 266 (2020) 117111
U. Hamid, et al.
Fig. 17. Effect of methane flow rate on CO2 emission and economic analysis.
4. Economic analysis
while case 1 offer 32.37 tonne/hr of H2. The results showed that case 2
provide better and cost-effective method to produce H2 which resulted
in 25% increase in hydrogen generation with 18% reduction in total
investment cost per tonne of H2 generation.
4.1. Estimation of capital expenditure (CAPEX)
As described in the process description, hydrogen is a clean fuel
which could be produced from gasification of coal and reforming of
methane in a parallel configuration. Hence, multiple processes i.e. solid
handling, gasification island, syngas processing, acid gas removal, air
separation, and reformer are harmoniously integrated to maximize the
hydrogen production while utilizing the heat within the overall process.
Therefore, economic analysis of integrated process is essential to evaluate the capital cost of conventional gasification process (case 1) and
heat integrated gasification and steam methane reforming process (case
2) that contain additional reactors. The influencing factors for cost estimation are plant capacity, raw material utilization, size of plant, plant
efficiency, and some inevitable uncertainties associated with the process plant. So, Eq. (2) provides reasonable cost estimation for integrated
gasification process that also known as power law of capacity where CE
indicate equipment’s cost, CB represents cost of base case, Q denotes
reference capacity, M is a constant, IB is the cost index of base case, and
IE is the cost index of calculating year [25,44]. Through literature, the
capital and variable costs for both cases were determined upon specifying the capacity of each section [45–47].
CE = CB ×
Q
QB
M
×
IE
IB
TIC per tonne of H2 =
Total investment cost
Hydrogen generation
(3)
4.2. Estimation of operational and maintenance (O&M) expenditures
Fixed and variable expenditures are represented by the operational
and maintenance (O&M) expenditures. The key dependence of O&M
expenditure is on fuel prices, Selexol and catalyst requirement, maintenance cost, and operational hours. The O&M expenditures for both
cases are reported in Table 8 with some operational indicators extracted
from the literature [45–47]. The results showed that case 2 present
higher total fixed and variable cost compared to case 1 because of additional natural gas and steam methane reforming catalyst. However,
total fixed and variable cost to produce per tonne of H2 indicated that
case 2 is economically benign because of low cost and high H2 generation. Similarly, lifetime cost of the process is also being computed
with the assumption that process construction time is three year with
life span of 25 year having operational hours per year equals to 7500 h.
Cost comparison indicated that case 2 has the tendency to produce H2
offer 8% life-time cost reduction as compared to case 1. Therefore,
operational and maintenance expenditures revealed that case 2 is present more benefits in term of cost to produce H2.
(2)
Some economic analysis-based assumptions have been made that
are offsite and utilities are 25% of equipment cost, contingency is 15%
of installed cost, land cost is 5% of installed cost. Finally, calculations
are performed to estimate the cost of each section using and ultimately
total investment cost, which is sum of total installed cost and total cost.
The capital cost expenditure of case 1 and case 2 is presented in Table 7.
Case 1 and case 2 possess same cost for solid handling, gasification
island, syngas processing unit, and air separation unit because the base
remain same. As, case 2 presents higher syngas flowrates due to which
acid gas removal section cost increases. Based on the equipment cost of
each section, installed cost, contingency, land cost, and utilities’ cost
are being computed. Afterward, comparison of total investment cost
(TIC) associated with case 1 and case 2 is reported based on cost per
tonne of H2, as represented in Eq. (3). The results indicated that case 2
presents higher flowrate of hydrogen production i.e. 40.53 tonne/hr
4.3. Compilation and comparison of key results
Some of the key results are summarized in Table 9 to provide quick
comparison between case 1 and case 2. For instance, case 2 has lower
CO2 emission per unit H2 production as compared to case 1 while offering energy saving of 47 MWth and increased overall hydrogen production. Additionally, higher heating value (HHV) for case 2 is also
higher than case 1 design. Besides technical assessment, economic
analysis indicates that hydrogen cost can be reduced from 484.7 €/
tonne to 446.14 €/tonne due to an improved process efficiency, energy
savings, hydrogen production and product purity regardless of increased capital investment.
10
Fuel 266 (2020) 117111
U. Hamid, et al.
5. Selecting optimum process variables
product gas. Therefore, benefits of both processes collectively come up
with more innovative option to be implemented pragmatically.
Ultimately, a new heat integrated model has been developed which
efficiently utilize the heat recovered from the syngas to preheat the
steam methane reforming reactants and proceed the endothermic reforming reactions while maintain reformer temperature. Case 1 is a
conventional coal-based gasification model, whereas, case 2 represents
the strategy to utilize the gasification heat for NG reforming. Both case
studies are compared based on H2/CO (HCR) ratio, cold gas efficiency
(CGE), higher heating value (HHV), syngas compositions, and CO2
specific emissions. Comparative analysis showed that case 1 presents
HCR ratio of 0.766 while case 2 offers 1.05, which indicates the higher
H2 content as compared to carbon monoxide for case 2. Similarly, based
on cold gas efficiency of process, case 2 presents 73.7% CGE, whereas,
case 1 represents only 67.0% CGE. The results of HHV for case 2 and
case 1 has been evaluated as 149.56 MJ/kmol and 137.75 MJ/kmol,
respectively. Similarly, CO2 specific emissions analysis also showed that
case 2 offers less CO2 emission compared to case 1. Besides these,
economic analysis indicated that case 2 reduces both capital cost expenditures and operational and maintenance expenditure as compared
to case 1. So. Case 2 design has been evaluated as the best case in terms
of cold gas efficiency, H2 to carbon monoxide ratio, higher heating
value, improved H2 production, and less greenhouse gas emissions as
compared to case 1. Conclusively, case 2 is not only technically a valid
schematic but also presents with cost-effective solution to maximize the
hydrogen generation by 25%.
5.1. Optimal parameters for CO2 emission
The optimum value of some important parameters is obtained upon
performing the sensitivity analysis. The effect of water flowrate on
gasifier product composition and oxygen requirement for gasification
process is reported in the Fig. 13. The results showed that carbon dioxide content in product syngas increases with an increase in water
flowrate for gasification process. Additionally, carbon monoxide and
hydrogen which are essential components of syngas keep on decreasing
while water content in the syngas also started to increase steadily.
Besides these, oxygen requirement to maintain the gasifier temperature
of 1370°C also indicates an increasing trend starting from nearly
142 tonne/hr to 208 tonne/hr. Furthermore, methane flowrate for reforming process is also being analyzed with respect to CO2 emissions
and H2 production as presented in Fig. 14. The results indicate that the
overall CO2 emission started to increase as methane flowrate increases
while hydrogen production started to decrease. Therefore, the sensitivity analysis of these parameters indicated that 30% water in coal
slurry is enough to maintain the flow of slurry and insure the feasible
water content of 18% in the gasifier product syngas. Similarly, the
optimum methane flowrate considered to be 0.715 MMCUFTD which
result efficient heat integration to maintain reformer temperature
around 912°C, minimum CO2 emission, and H2 production.
5.2. Optimal process parameters for economic analysis
CRediT authorship contribution statement
To obtain the optimal process values considering both capital expenditures (CAPEX) and operational expenditure (OPEX) are represented besides CO2 emissions. According to Fig. 15, slurry water
contents have been evaluated against CAPEX and OPEX, which reveal
that both the costs increases with an increase in water content in the
slurry. Furthermore, the cost associated with hydrogen production for
the overall process has been evaluated which showed that the H2 production cost increases with an increase in the slurry’s water. Moreover,
the load on air separation unit (ASU) also increases to meet the process
requirement which itself is an energy intensive process. The results in
Fig. 16 showed that cost associated with hydrogen production can be
minimized by reducing the water content in the coal-slurry while
meeting the process constraints.
Usman Hamid: Simulation, Validation, Formal analysis,
Investigation, Writing - original draft. Ali Rauf: Writing - review &
editing, Supervision, Project administration. Usama Ahmed:
Conceptualization, Methodology, Supervision. Md. Selim Arif Sher
Shah: Visualization. Nabeel Ahmad: Formal analysis.
Declaration of Competing Interest
The authors declare that they have no known competing financial
interests or personal relationships that could have appeared to influence the work reported in this paper.
Acknowledgements
5.3. Effect of process parameters on emissions and economics
The authors would like to acknowledge the support provided by the
Lahore University of Management Sciences (LUMS), Pakistan, and King
Fahd University of Petroleum & Minerals (KFUPM), Saudi Arabia, for
conducting this research.
Coal and methane flow rates are the two components that influence
the overall CO2 emissions by contributing carbon. For both cases, coal
flowrate remains constant due to which the only variable that can be
considered for simultaneous process economics and CO2 emissions is
the methane flowrate. It has been seen from the results that CO2
emissions become a function of methane flow rate and increases with
an increase in natural gas used for reforming. According to Fig. 17,
CAPEX, OPEX and CO2 emissions increases with an increase in the
methane flowrate to the reforming unit. The huge volumes of reformer
derived synthesis gas increases the size of downstream process units
that increases the CAPEX of the overall process. Similarly, the results
showed that the OPEX is also increased with an increase in methane
flowrate.
References
[1] BP. BP Energy Outlook 2019. 2019.
[2] Zeng S, Chen J. Forecasting the Allocation Ratio of Carbon Emission Allowance
Currency for 2020 and 2030 in China. Sustainability. 2016;8.
[3] Enerdata. Global Energy Statistical Yearbook 2018. Enerdata 2018.
[4] Jackson RB, Le Quéré C, Andrew RM, Canadell JG, Peters GP, Roy J, et al. Warning
signs for stabilizing global CO2 emissions. Environ Res Lett 2017;12.
[5] IPCC. IPCC Fourth Assessment Report: Climate Change 2007. Interngovernmental
Panel on Climate Change (IPCC); 2007.
[6] Ali U, Font-Palma C, Akram M, Agbonghae EO, Ingham DB, Pourkashanian M.
Comparative potential of natural gas, coal and biomass fired power plant with post combustion CO2 capture and compression. Int J Greenhouse Gas Control
2017;63:184–93.
[7] (EIA) EIA. Annual Energy Outlook 2019 with Projections to 2050. 2019.
[8] (EIA) EIA. Annual Energy Outlook 2018 with projections to 2050. 2018.
[9] Sánchez-Hervás José María, Moya Gregorio Molina, Ortiz-González Isabel. Current
status of coal gasification. New Trends in Coal Conversion. Elsevier; 2019. p.
175–202.
[10] Kalamaras CM, Efstathiou AM. Hydrogen production technologies: current state and
future developments. Conference Papers in Energy 2013;2013:1–9.
[11] Council WE. World Energy Resources. 2013.
6. Conclusions
The gasification and steam methane reforming are among the most
cost effective and environmentally feasible technologies to produce
syngas and H2. Steam methane reforming process usually offers higher
H2/CO (HCR) ratios but less resources of NG compared to coal limit its
wide utilization. In contrast, gasification operation is being carried out
at high temperature followed by the significant heat recovery from
11
Fuel 266 (2020) 117111
U. Hamid, et al.
[12] (EIA) EIA. International Energy Outlook 2016. 2016.
[13] Tremel A, Gaderer M, Spliethoff H. Small-scale production of synthetic natural gas
by allothermal biomass gasification. Int J Energy Res 2013;37:1318–30.
[14] Cormos A-M, Dinca C, Cormos C-C. Multi-fuel multi-product operation of IGCC
power plants with carbon capture and storage (CCS). Appl Therm Eng
2015;74:20–7.
[15] Pettinau A, Ferrara F, Amorino C. Combustion vs. gasification for a demonstration
CCS (carbon capture and storage) project in Italy: a techno-economic analysis.
Energy 2013;50:160–9.
[16] Simpson AP, Lutz AE. Exergy analysis of hydrogen production via steam methane
reforming. Int J Hydrogen Energy 2007;32:4811–20.
[17] Onozaki M, Watanabe K, Hashimoto T, Saegusa H, Katayama Y. Hydrogen production by the partial oxidation and steam reforming of tar from hot coke oven gas.
Fuel 2006;85:143–9.
[18] Basile AB, Tong J, Millet P. Inorganic membrane reactors for hydrogen production:
an overview with particular emphasis on dense metallic membrane materials.
Handb Membr React 2013:42–148.
[19] Sengodan S, Lan R, Humphreys J, Du D, Xu W, Wang H, et al. Advances in reforming and partial oxidation of hydrocarbons for hydrogen production and fuel
cell applications. Renew Sustain Energy Rev 2018;82:761–80.
[20] Aasberg-Petersen K, Bak Hansen JH, Christensen TS, Dybkjaer I, Christensen PS,
Stub Nielsen C, et al. Technologies for large-scale gas conversion. Appl Catal A
2001;221:379–87.
[21] Yilmaz F, Ozturk M, Selbas R. Design and thermodynamic analysis of coal-gasification assisted multigeneration system with hydrogen production and liquefaction.
Energy Convers Manage 2019;186:229–40.
[22] Ahmed U, Zahid U, Jeong YS, Lee C-J, Han C. IGCC process intensification for simultaneous power generation and CO2 capture. Chem Eng Process Process Intensif
2016;101:72–86.
[23] Hasan A, Dincer I. Comparative assessment of various gasification fuels with waste
tires for hydrogen production. Int J Hydrogen Energy 2019.
[24] Al-Zareer M, Dincer I, Rosen MA. Effects of various gasification parameters and
operating conditions on syngas and hydrogen production. Chem Eng Res Des
2016;115:1–18.
[25] Cormos C-C. Integrated assessment of IGCC power generation technology with
carbon capture and storage (CCS). Energy 2012;42:434–45.
[26] Ng KS, Zhang N, Sadhukhan J. Techno-economic analysis of polygeneration systems
with carbon capture and storage and CO2 reuse. Chem Eng J 2013;219:96–108.
[27] Gasification Higman C. Combustion engineering issues for solid fuel systems.
Elsevier; 2008. p. 423–68.
[28] Wang Y, Wang J, Luo X, Guo S, Lv J, Gao Q. Dynamic modelling and simulation of
IGCC process with Texaco gasifier using different coal. Syst Sci Control Eng
2015;3:198–210.
[29] Kapetaki Z, Ahn H, Brandani S. Detailed process simulation of pre-combustion IGCC
plants using coal-slurry and dry coal gasifiers. Energy Procedia 2013;37:2196–203.
[30] Cormos A-M, Szima S, Fogarasi S, Cormos C-C. Economic assessments of hydrogen
production processes based on natural gas reforming with carbon capture. Chem
Eng Trans 2018;70:1231–6.
[31] Szima S, Cormos A-M, Cormos C-C. Flexible Hydrogen and Power Co - generation
[32]
[33]
[34]
[35]
[36]
[37]
[38]
[39]
[40]
[41]
[42]
[43]
[44]
[45]
[46]
[47]
12
based on Dry Methane Reforming with Carbon Capture. 28th European Symposium
on Computer Aided Process Engineering2018. pp. 1281–1286.
Qian Y, Man Y, Peng L, Zhou H. Integrated process of coke-oven gas tri-reforming
and coal gasification to methanol with high carbon utilization and energy efficiency. Ind Eng Chem Res 2015;54:2519–25.
Yi Q, Feng J, Wu Y, Li W. 3E (energy, environmental, and economy) evaluation and
assessment to an innovative dual-gas polygeneration system. Energy
2014;66:285–94.
Adams TA, Barton PI. Combining coal gasification and natural gas reforming for
efficient polygeneration. Fuel Process Technol 2011;92:639–55.
Cormos A-M, Cormos C-C. Techno-economic assessment of combined hydrogen &
power co-generation with carbon capture: the case of coal gasification. Appl Therm
Eng 2019;147:29–39.
Ahmed U, Kim C, Zahid U, Lee C-J, Han C. Integration of IGCC and methane reforming process for power generation with CO2 capture. Chem Eng Process Process
Intensif 2017;111:14–24.
Ahmed U, Zahid U, Lee Y. Process simulation and integration of IGCC systems for
H2/syngas/electricity generation with control on CO2 emissions. Int J Hydrogen
Energy 2019;44:7137–48.
Adams TA, Barton PI. High-efficiency power production from coal with carbon
capture. AIChE J 2010;56:3120–36.
He C, Feng X, Chu KH, Li A, Liu Y. Industrial-scale fixed-bed coal gasification:
modeling, simulation and thermodynamic analysis. Chin J Chem Eng
2014;22:522–30.
Motta IL, Miranda NT, Maciel Filho R, Wolf Maciel MR. Sugarcane bagasse gasification: simulation and analysis of different operating parameters, fluidizing media,
and gasifier types. Biomass Bioenergy 2019;122:433–45.
Fout T, Zoelle A, Keairns D, Turner M, Woods M, Kuehn N, et al. Cost and
Performance Baseline for Fossil Energy Plants Volume 1b: Bituminous Coal (IGCC)
to Electricity Revision 2b—Year Dollar Update. Washington, DC, USA: United States
Departmeant of Energy; 2015.
Woods MC, Capicotto P, Haslbeck JL, Kuehn NJ, Matuszewski M, Pinkerton LL,
et al. Cost and performance baseline for fossil energy plants. National Energy
Technology Laboratory; 2007.
Field RP, Brasington R. Baseline flowsheet model for IGCC with carbon capture. Ind
Eng Chem Res 2011;50:11306–12.
Seider WD, Seader JD, Lewin DR. Product & process design principles: synthesis
analysis and evaluation. John Wiley & Sons; 2009.
Majoumerd MM, Assadi M, Breuhaus P. Techno-Economic Evaluation of an IGCC
Power Plant With Carbon Capture. ASME Turbo Expo 2013: Turbine Technical
Conference and Exposition: American Society of Mechanical Engineers; 2013. p.
V002T03A16-VT03A16.
Gandrik A, Wood R, Patterson M, Mills P. HTGR-Integrated Hydrogen Production
via Steam Methane Reforming (SMR) Economic Analysis. Document ID: TEV-954
INL. 2010.
Padurean A, Cormos C-C, Agachi P-S. Pre-combustion carbon dioxide capture by
gas–liquid absorption for Integrated Gasification Combined Cycle power plants. Int
J Greenhouse Gas Control 2012;7:1–11.