Performance evaluation and carbon assessment of IGCC power plant with coal quality

Energy 188 (2019) 116063

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Performance evaluation and carbon assessment of IGCC power plant with coal quality Hyun-Taek Oh, Woo-Sung Lee, Youngsan Ju, Chang-Ha Lee* Department of Chemical and Biomolecular Engineering, Yonsei University, Republic of Korea

a r t i c l e i n f o

a b s t r a c t

Article history: Received 20 April 2019 Received in revised form 30 August 2019 Accepted 3 September 2019 Available online 5 September 2019

Techno-economic and environmental impacts of coal type were evaluated using a 500 MW-class integrated coal gasification combined cycle(IGCC), including reheat combined cycle process with threepressure level based on higher than 99.9% sulfur removal and 90% carbon capture. Efficiency and cost of electricity(COE) of four different coals in the IGCC power plant were compared: two bituminous and two sub-bituminous coals. As coal with higher heating value per unit weight was fed into a gasifier, higher cold gas efficiency of the gasifier and greater net overall plant efficiency was achieved. The highest overall plant efficiency of 31.62% could be achieved by using bituminous. Raw water consumption was also affected by the moisture content of the as-received coal. The as-received coal with the highest moisture content consumed the least amount of water. The exergy flow and destruction were presented in Grassmann diagrams to provide more detailed information on main units. However, according to the sensitivity test, the COE was mainly influenced by fuel prices and costs of CO2 transport and storage. If the price difference between bituminous coals and sub-bituminous coals is reduced, the utilization of bituminous coals becomes more competitive in terms of COE and capital cost. © 2019 Elsevier Ltd. All rights reserved.

Keywords: IGCC Performance evaluation Carbon assessment Coal type Cost of electricity Carbon capture and storage

1. Introduction Many countries have enacted environmental policies and regulations intended to reduce greenhouse gas emissions from electric power plants by decreasing the use of fossil fuels. It is estimated that coal-fired generation, which accounted for 40% of total world electricity generation in 2012, will be reduced to 29% of the total by 2040 [1]. However, coal continues to be the largest single fuel used for electricity generation worldwide because coal is the most abundant and least expensive fossil fuel [2,3]. Coal accounted for 45% of the global CO2 emissions in 2015 due to its heavy carbon content per unit of energy released [4,5]. Hence, the present challenge for electric power generation is to meet the ever-increasing demand for electricity and simultaneously mitigate greenhouse gas emissions. The integrated coal gasification combined cycle (IGCC) power plants have higher efficiencies and better environmental advantages than pulverized coal power plants. It was reported that precombustion carbon capture has the highest potential for

* Corresponding author. E-mail address: [email protected] (C.-H. Lee). https://doi.org/10.1016/j.energy.2019.116063 0360-5442/© 2019 Elsevier Ltd. All rights reserved.

application to the IGCC process because of its cost efficiency for CO2 capture [3,6,7]. CO2 captured from large-scale power plants is compressed to high pressures, transported to a storage site, and injected into a suitable geological formation. The cost of carbon capture and storage (T&S) gives a significant impact on the cost of electricity (COE) production. Overall process performance is significantly affected by coal quality, such as coal composition, moisture content, and heating value; however, the quality varies widely depending on the geographical location of the coal source [2,8]. Low-rank coals, i.e., sub-bituminous and lignite coals, retain greater fractions of moisture and volatile matter and contain less fixed carbon than highrank coals, i.e., bituminous and anthracite coals. However, since about 53% of global coal reserves and about 36% of United States (U.S.) recoverable coal reserves are of low-rank coals, i.e., subbituminous and lignite coals [9,10], the performance evaluation of IGCC power plant using low-rank coals is important. Generally, the performance evaluations of overall IGCC processes [11,12] and the co-production of electricity and H2 from the IGCC process [13,14] were conducted using single specific coal. Many studies did not consider the cost of CO2 T&S, but it has to be also evaluated by the cost analysis of IGCC performance if CO2 sequestration is practically considered [15e17].

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H.-T. Oh et al. / Energy 188 (2019) 116063

Nomenclature AGRU ASU CC CCP CO2 T&S COE DEPG DOE EIA ELECNRTL EOS GCC GT HHV

Acid gas removal unit Air separation unit Combined Cycle Carbon capture process CO2 transport and storage Cost of electricity Dimethyl ether of polyethylene glycol Department of Energy Energy information administration Electrolyte nonrandom two-liquid activity coefficient model Equation of state Grand composite curve Gas turbine Higher heating value

The performance of the gasifier is influenced by the higher heating value (HHV) of the coal source [8,18]. In addition, the moisture content of the coal also impacts on the performance of the gasifier and coal dryer [19,20]. Because the composition and amount of syngas produced from a gasifier and steam required for a coal dryer depend on the coal type, the operation of the overall IGCC process is also changed to achieve the high performance and environmental regulations. Thus, considering the mitigation of greenhouse gas and environmental regulations, the comparison study for the utilization of various types of coal in overall IGCC process is required to analyze the relative performance and cost, including the cost of CO2 capture, transport, and storage. In this study, a techno-economic analysis was performed for a 500 MW (MW)-class Shell gasifier IGCC power plant with a carbon capture process (CCP) under the conditions of higher than 99.9% sulfur removal and 90% carbon capture from flue gas. Comparative performance evaluation of four difference coals was conducted for the overall IGCC process including a reheat combined cycle (CC) process with three-pressure level and a dual-stage Selexol process with more than 90% carbon capture: two bituminous coals (Pittsburgh #8 and Illinois #6) and two sub-bituminous coals (Montana and Wyoming). The IGCC process with the Selexol process, a steam turbine (ST) and heat recovery steam generator (HRSG) was simulated in detail and the maximum net plant efficiency was estimated. Water consumption for the overall IGCC process using four different coals was also compared, and the effects of steam flow on the ST and HRSG was evaluated. The exergy analysis for one bituminous and one sub-bituminous coal was conducted to provide more detailed information on the direction of further energy improvement. The estimated COE considering CO2 T&S cost and coal price was compared among the IGCC processes using the four coals. Comparative results by using the different types of coal contribute to establishing a guideline for coal selection in an IGCC power plant and serving for planning and implementation of a carbon tax. 2. Process description The previous studies for overall IGCC process analysis can be categorized into two subjects. Firstly, research is focused on various IGCC units with a brief model of ST, gas turbine (GT), and HRSG units [13,21,22] where the efficiency of each unit and overall IGCC process was evaluated by changing the variables of a specific unit,

HP HRSG I6 IGCC IP LHV LP MR NETL P8 PC-SAFT ST TEG TOT WGSR WY

High pressure Heat recovery steam generator Coal Illinois #6 Integrated gasification combined cycle Intermediate pressure Lower heating value Low pressure Coal Montana Rosebud National Energy Technology Laboratory Coal Pittsburgh #8 Perturbed-chain statistical associating fluid theory Steam turbine Triethylene glycol Turbine outlet temperature Water-gas shift reactor Coal Wyoming

such as the CCP or gasifier. Secondly, the research is focused on the water network and the overall IGCC process configuration, which was integrated by each unit process with others [11,23e26]. However, the studies have focused on the efficiency of the entire IGCC process without considering water consumption. In addition, although the process efficiency can be affected by water/steam flow, few studies have discussed it in detail. In this study, the 500 MW-class IGCC power plant employed a coal feed, an air separation unit (ASU), Shell gasifier, water-gas shift reactor (WGSR), dual-stage Selexol process, Claus process, CO2 compressor, and CC. The configuration of the IGCC process with carbon capture was based on the National Energy Technology Laboratory (NETL) data (Case B1B, hell IGCC with CO2 Capture) and is presented in Fig. 1 [27]. Intermediate pressure (IP) steam with high purity oxygen was added to increase the efficiency of the Shell gasifier [18] and a partial water quench method was applied for the produced syngas [3]. ASU and GT were integrated because the power output of GT can be increased by maintaining optimum firing temperatures for syngas operation through increased nitrogen mass flow in the turbine [27]. From an analysis of the literature for double and triple pressure HRSGs with and without steam reheating models [28,29], the reheat HRSG with triple pressure level was used in the study. In Case B1B, the entrained-flow Shell gasifier is considered one of the most mature gasifiers for an IGCC process with a CCP [16]. The configuration of the CC using specific coal was reported in various references [30e32]. Detailed information including steam energy and effluent gas is also reported in the literature [18,19]. Since the coal quality has a significant influence on the performance of gasifier [4,5], the overall IGCC process is also affected by coal type. Comparison of the efficiency of the IGCC process in terms of the type of gasifier is widely studied, but there are few studies on the influence of coal quality (coal rank) [8]. Here, four commercial coals commonly used for power plants were utilized in the IGCC process with carbon capture: two bituminous coals (Pittsburgh # 8: Coal_P8 and Illinois # 6: Coal_I6) and two sub-bituminous coals (coal from Montana Rosebud: Coal_MR and Wyoming coal: Coal_WY). The properties of the four coals and the operating parameters for process simulation are listed in Tables 1 and 2. The specifications of each sub-unit, as well as assumptions and boundary conditions, are described in the following sub-sections.

H.-T. Oh et al. / Energy 188 (2019) 116063

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Fig. 1. Configuration of integrated gasification combined cycle (IGCC) with CCP.

2.1. Coal feed and ASU The volume of dried coal required was determined as equivalent to the HHV in the NETL Case B1B (1,590,722 kW). It was assumed that the as-received coal was dried to 5 wt% moisture for smooth flow through the dry feed system, but Coal_P8 (2.63 wt% in as-

received as shown in Table 1) was dried to a 2.5% content [18]. Coal drying was typically achieved by recycling a small volume of hot syngas or introducing superheated steam [34]. In this study, superheated steam was used for coal drying. The coal drying energy was estimated based on the documented data [20]. The cryogenic air separation unit was designed to produce

Table 1 Information of as-received coals used in the process simulation [17,18].

Reference As received HHV, MJ/kg Moisture, (ar), wt% Proximate analysis, dry wt% Volatile matter Fixed carbon Ash FC/VM ratio Ultimate analysis, dry wt% Carbon Hydrogen Nitrogen Chlorine Sulfur Oxygen Ash Mass flow rate of feed (kg/s) Coal before drying Coal after drying O2 Steam

Bituminous

Bituminous

Sub-bituminous

Sub-bituminous

Coal_P8

Coal _I6

Coal _MR

Coal _WY

[18] 30.50 2.63

[18] 27.10 11.12

[18] 19.9 25.77

[17] 19.40 30.20

36.79 53.79 9.42 1.46

39.37 49.72 10.91 1.26

40.88 48.09 11.03 1.18

41.42 50.95 7.63 1.23

75.13 5.10 1.50 0.04 2.42 6.39 9.42

71.73 5.06 1.41 0.33 2.82 7.74 10.91

67.45 4.56 0.96 0.01 0.98 15.01 11.03

69.07 4.74 1.00 0.01 0.53 17.02 7.63

52.14 52.07 41.66 5.21

58.70 54.92 41.19 5.49

79.94 62.46 44.35 0.00

82.00 60.25 43.50 0.00

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H.-T. Oh et al. / Energy 188 (2019) 116063 Table 2 Main assumptions of process simulation [33]. Unit

Parameters

Air separation unit

Oxygen purity: 95% (vol.) Oxygen delivery pressure: 8.6 bar Nitrogen delivery pressure: 3.9/12.5 bar Compressor polytropic efficiency: 80% Inter-cooling temperature: 40
C Dried coal moisture content (wt%): 5% (P8: 2.5%) Gasifier pressure: 45 bar Outlet syngas pressure/temperature: 40.3 bar/232
C Water quench type Wet scrubber pressure drop: 0.35 bar H2OeCO ratio: 1.8 Heat exchanger pressure drop: 0.35 bar Syngas inlet pressure: 35.5 bar Pumps efficiency: 75% Recycle compressor isentropic efficiency: 75% Inlet solvent temperature at absorber: 4
C H2S removal efficiency: >99.9% CO2 removal efficiency: >90% Delivery pressure: 153 bar CO2 compressor polytropic efficiency: 85% Ambient air pressure: 1.013 bar Ambient air temperature: 15
C Main air compressor polytropic efficiency: 85% Inlet syngas LHV at combustor: 4.4e4.7 MJ/m3 Turbine outlet temperature: 562
C Gas turbine isentropic efficiency: 89.8% Gas turbine pressure ratio: 15.8 bar/1.05 bar Pumps efficiency: 75% Heat exchanger gas side pressure drop: 2% Heat exchanger liquid side pressure drop: 4% Gas-Gas end temperature approach: 28
C Gas-Liquid end temperature approach: 14
C Liquid-Liquid end temperature approach: 5.5
C HP ST pressure ratio: 124 bar/40.6 bar IP ST pressure ratio: 39 bar/4.5 bar LP ST pressure ratio: 4.5 bar/0.085 bar HP/IP/LP ST isentropic efficiency: 90.3/93.5/88.2% Inlet HP/IP steam temperature: 534
C Condenser pressure: 0.085 bar Cooling water range: 11
C Evaporative losses: 0.8% Cycles of concentration: 4

Gasification (Shell Gasifier)

Water-gas shift reactor unit

Acid gas removal unit

CO2 compression Gas turbine

Heat recovery steam generator

Steam turbine

Cooling tower

95 vol% of O2 for the gasifier and Claus plant [35,36]. The cryogenic distillation unit separated air at 0.86 MPa, and the produced oxygen used in the gasifier was pressurized to 6.48 MPa via an intercooled compressor. High-purity N2 was recovered, compressed, and used as a diluent in the GT combustor. In the study, the energy input of compression work for the ASU unit was considered in the energy consumption analysis [36].

2.2. Gasifier with syngas cooling, and WGSR unit The dried, pulverized coal delivered to the Shell gasifier with IP steam and high purity oxygen at high pressure. Since cold gas efficiency indicates the conversion degree of coal energy into the thermal energy of syngas during gasification, it can be defined as the heating value of product syngas divided by the heating value of coal. The volume of steam and oxygen required for each coal was determined to maximize the cold gas efficiency of a gasifier [18]. High-temperature syngas cooling consisted of four steps. The top product syngas from the gasifier was cooled to approximately 1090
C by adding cooled recycle syngas to lower the temperature below the ash melting point. Syngas went through a syngas cooler. The gas temperature was reduced from approximately 1090
Ce900
C, and high pressure (HP) steam for the steam cycle was produced. Then, water quench followed to cool the raw syngas

from 900
C to 400
C while providing a portion of the water was required for a WGSR as shown in Fig. 1. Finally, syngas was cooled from 400
C to 235
C, and IP steam for a steam cycle was produced. Before the syngas entered a WGSR unit, the particles, NH3, and chlorides in the cooled gas were removed in a water-based scrubber. The syngas exited from the scrubber was maintained at 14
C below the dew point. To facilitate CCP and increase the efficiency of carbon capture and hydrogen production, CO was converted to CO2 through the WGSR. The steam to CO molar ratio was adjusted to approximately 1.8:1 by the addition of IP shifted steam to the syngas stream for promoting a high conversion of CO in Fig. 1. In the two-stage WGSR, high- and low-temperature shift reactors in series were used to achieve sufficient conversion. Cooling was provided between the series of reactors to control the exothermic temperature rise, and IP steam was produced for the ST. The resulting overall conversion of CO to CO2 in the WGSR process was set to 98% (molar basis).

2.3. Dual-stage selexol & claus process During the gasification process, the sulfur content in the dried coal was converted to H2S and COS. The COS in the syngas was more than 98% hydrolyzed using steam from the WGSR. A dual-stage Selexol process was used as an acid gas removal unit (AGRU) to

H.-T. Oh et al. / Energy 188 (2019) 116063

separate H2S and CO2 from the syngas. Since the Selexol solvent is a physical solvent without the reaction of CO2, the absorber column in the Selexol process was simulated as an equilibrium model rather than as a rate-based model in Aspen Plus. The simulation of the AGRU was previously reported in detail [1]. Since the solvent had a higher selectivity for H2S than CO2, the first absorber preferentially removes H2S from the syngas while the second absorber captures CO2 from the syngas [37]. To meet stringent environmental regulations and prevent turbine damage, more than 99.9% of the sulfur in the syngas entering the AGRU was separated, and the separated sulfur from the first absorber was converted to elemental sulfur solids with more than 99% efficiency through the Claus process. Several studies report that carbon capture efficiency is the most economical at 90e95% [14,17]. Therefore, in this study, the volume of Selexol solvent was determined to satisfy 90% of the carbon capture efficiency under the condition of more than 99.9% sulfur removal. 2.4. CO2 compression, transport and storage To prevent the corrosion of the CO2 transport pipeline, triethylene glycol (TEG) is used to remove water from the captured CO2 product before compressing the CO2 product. CO2 is released from the flash drums with three pressure levels for efficient absorbent regeneration. The HP CO2 stream is flashed at 2.0 MPa and compressed and recycled to the CO2 absorber. The IP CO2 stream is also flashed at 1.0 MPa. After the low pressure (LP) CO2 stream is flashed at 0.1 MPa, it is compressed to 1.0 MPa and combined with the IP CO2 stream. The combined stream is compressed from 1.0 MPa to 15.3 MPa under a temperature of 50
C, using an intercooled multi-stage compressor unit. CO2 T&S is another important factor to evaluate the cost of carbon assessment. The cost of CO2 T&S depends on the storage location due to the difference in the geological formations that make up saline aquifers [38,39]. The cost evaluation for CO2 T&S used in this study was referenced from data in the NETL report [27,38]. 2.5. Combined cycle (CC) The CC integrated the GT, HRSG, and ST. The flow direction of steam and water is indicated by arrowed lines in Fig. 1. The GT unit consisted of an air compressor to supply HP oxygen to a combustion chamber and a GT. To reduce the formation of thermal NOx, high purity N2 separated from the ASU was added to the GT as a diluent until the lower heating value (LHV) of syngas (position 3 in Fig. 1) reached 4.4e4.7 MJ/m3 [27] and the turbine outlet temperature (TOT) of GT reached 562
C through controlling the volume of high-pressure air fed to the combustor. The temperature and composition of the flue gas from the combustor were calculated by the RGibbs reactor model, which uses Gibbs free energy minimization to calculate equilibrium in Aspen Plus. The heat of the exhaust gas from a GT was recovered through a HRSG. Since the HRSG was configured by a multi-pressure steam generation unit, several studies have compared and analyzed double and triple HRSGs with and without steam reheating HRSG models [28,40e42]. Because the triple HRSG with a reheater was reported to be highly efficient and cost-effective in several studies [40,42,43], this system was selected for this study. After HP steam was expanded in a HP turbine, some of the exhaust steam was supplied to the gasifier for syngas gasification, and the other was mixed with IP steam and sent to the HRSG for reheating as presented in Fig. 1. The exhaust stream of an IP turbine was mixed with LP steam and fed directly to a LP turbine. In order to reduce energy consumption during steam

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pressurization, the exhaust steam of the LP turbine was liquefied with cooling water from a condenser. Liquefied low-pressure water was pumped to HP, IP, and LP using a pump. Some of the LP water was sent to the Claus process and the remainder was supplied to the HRSG for heating. Some of the IP water was fed to the gasifier, WGSR, and Claus process and the remainder was sent to the HRSG for heating. The generated IP steam was fed to the WGSR as a shift steam and some of the IP steam was sent to the reboiler of the dualstage Selexol process and sour-water stripper for heat supply. Some of the HP water was sent to the syngas cooler, and the remainder was supplied to the HRSG for heating. The detailed configuration of the HRSG is presented in Fig. 1. The net power of the ST was optimized by adjusting the volume of HP, IP, and LP water. Exhaust gas from the GT, which was recovered through the HRSG, was discharged at 132
C to avoid the formation of sulfuric acid in the stack [40]. Steam was generated from the HP turbine, IP turbine, and LP turbine. Since the temperature of the flue gas from the GT was 562
C, the superheated steam entered each turbine at pressure and temperature, 12.5 MPa at 534
C, 3.9 MPa at 534
C and 0.45 MPa at 240
C, respectively. The operating conditions of the CC are listed in Table 2. The cooling water was applied to several units when the temperature of the stream was not high enough to allow heat recovery by LP steam, such as in the condenser in CC and multistage compressor. The units using the cooling water are presented by a blue dash line in Fig. 1. The largest volume of raw water consumption in all the studied cases resulted from the make-up for a cooling tower. Cooling water was not consumed by the process directly because it was circulated to decrease the temperature of streams. When hot water was cooled by using air in the cooling tower, a large volume of water was evaporated into the air. In addition, water loss occurred during a blow-down period in the water-cooling apparatus to reduce scale-formation. In this study, it was assumed that only evaporation loss and blow-down loss occurred in the cooling tower. The blow-down loss was as follows: [44].

Blow  down losses ¼

Evaporative losses ðCycles of ConcentrationÞ  1

(1)

where the cycles of concentration are a measure of water quality [33]. The detailed conditions of the cooling tower are listed in Table 2. 3. Process simulation methodology The process simulation of IGCC excluding a Shell gasifier was performed using Aspen Plus (Aspen Technology Inc., USA). Effluent gas composition and conditions of each coal from the Shell gasifier were obtained separately from a dynamic simulation using a gPROMS software package (Process Systems Enterprise, UK). The detailed information of the dynamic simulation of gasifier was reported in previous studies [18,45]. Most devices and units, such as the ASU, syngas cooling unit, WGSR, GT, and flue gas in the HRSG, were modeled using the PengeRobinson EOS with the Boston Mathias alpha function (PRBM) which could accurately calculate thermodynamic and transport properties for a nonpolar or mildly polar mixture, such as CO2, H2S, and H2. Peng-Robinson EOS predict well the properties of gaseous mixture with low molecular weight components and water vapor. Therefore, it was often used to simulate a unit where the gaseous mixture with water vapor existed [8,31,46]. And PR-BM could predict reasonable results at all temperatures and pressures [21,47]. The sub-unit using PR-BM was as follows: The water scrubber was based on the API sour model to

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H.-T. Oh et al. / Energy 188 (2019) 116063

correlate the volatility data of NH3, CO2, and H2S in the aqueous system [47]. The sour water stripper was modeled using the electrolyte nonrandom two-liquid activity coefficient model (ELECNRTL) to determine the physical properties of the electrolyte liquid phase [21]. The Selexol process was modeled using the perturbed-chain statistical associating fluid theory (PC-SAFT) because it covers small to large molecule fluid systems, including normal fluids, water, polymers and copolymers, and their mixtures [47]. The steam in the turbine and HRSG was modeled using NBS/ NRC steam tables 1984 [48] because steam tables can more accurately calculate all the thermodynamic properties for systems containing pure water or steam. All EOS pure and binary parameters in this study were obtained from the databanks in the Aspen Plus V8.8 library.

4. Economic evaluation Cost evaluation is essential for a plant feasibility study. In this study, important economic indicators, such as capital cost, operating cost, CO2 T&S cost and COE, were calculated. The capital cost was considered as equipment cost, material cost and others (direct and indirect costs). The operating cost was considered as variable operating cost and fuel cost, and did not include fixed operating costs. The cost of a specific item of plant equipment is a function of size, materials of construction, design pressure and temperature, etc. Equipment cost and material cost were calculated using a power law of capacity:


SC ¼ RC 

SP RP

Exp (2)

where SC is a scaled cost, SP is a scaled parameter, RC is a reference cost, RP is a reference parameter, and Exp is an exponent. For each unit, the process parameter and exponent were different. All the parameters and reference costs in Table 3 were referenced from the Department of Energy (DOE) report [49]. As the examples listed in Table 3, the reference costs were derived from the results of Case B1B in the DOE report (Illinois # 6 with CO2 capture) and the range indicated the recommended range of scaled parameters. It is assumed that the sum of the equipment cost and material cost accounts for 30% of the capital cost [50]. The utility consumption and waste volumes were obtained from the process simulation results. Operating cost included the costs of

water usage, waste water treatment, WGSR catalyst, Claus catalyst, slag disposal, and fuel, which were based on the price list from the NETL report [27]. The cost of CO2 T&S depends on the storage location. In this study, it was assumed that CO2 was stored at different locations, depending on the location of coal production, because coal transportation also represents a cost. The transport cost was calculated using the National Energy Technology Laboratory (NETL) CO2 Transport Cost Model, assuming a 100-km dedicated pipeline for each region to connect the IGCC plant to the CO2 storage site. In other words, it was assumed that the CO2 captured from Pittsburgh # 8 coal and Illinois # 6 coal were stored in the Illinois basin, the Montana Rosebud coal in the Montana Powder River basin, and Wyoming coal in the North Dakota Williston basin. The costs calculated by the NETL CO2 Saline Storage Cost Model covered all costs including capital costs, operating costs, financing costs, and taxes, and were expressed in dollars per ton of CO2 stored [38]. The COE is the electric power generation cost by the generator per net megawatt-hour. The COE was obtained from the following equation.

COE ¼

CCF  Capital Cost þ Operating Cost MWH

(3)

where the capital charge factor (CCF) matches the capital expenditure period and MWH is the annual net megawatt-hours of power generated. In this study, the CCF was assumed as the value of 0.109 [27]. All the costs were escalated to the 2016 US dollar using the Power Capital Cost Index [51]. An exergy is defined as the maximum amount of work obtainable when a given system progresses from its initial state to equilibrium state with the reference condition [52]. The total exergy of a material stream was as follows:

  Exergy ¼ N εph þ εch

(4)

εph ¼ ðh  h0 Þ  T0 ðs  s0 Þ

(5)

εch ¼

X X xi ε0;i þ RT0 xi lnxi i

(6)

i

where N is the molar flow rate in mol/s, εph is the molar physical exergy flow rate in kJ/mol, and εch is the molar chemical exergy

Table 3 Scaling parameters and exponents in Eq. (7) for economic evaluation [49]. Item Description Fuel & Feed Coal Crushing & Drying Prepared Coal Storage & Feed Dry Coal Injection System Misc. Coal Prep & Feed Coal & Sorbent Feed Foundation Gasifier Gasifier, Syngas Cooler ASU/Oxidant Compression Flare Stack System Gas Cleanup & Piping Double Stage Selexol Elemental Sulfur Plant Fuel Gas Piping CO2 Compression CO2 Compression & Drying HRSG & Stack Heat Recovery Steam Generator Stack

Parameter

Exponent

Range

Reference Parameter

Reference Cost, $

Coal feed rate, lb/hr Coal feed rate, lb/hr Coal feed rate, lb/hr Coal feed rate, lb/hr Total Feed Flow Rate, lb/hr

0.66 0.66 0.66 0.66 0.66

18,400e1,750,000 18,400e1,750,000 18,400e1,750,000 18,400e1,750,000 18,400e1,750,000

465,264 465,264 465,264 465,264 465,264

8656 9353 8803 2249 14,094

Total Feed Flow Rate, lb/hr Total Feed Flow Rate, lb/hr Total Feed Flow Rate, lb/hr

0.66 0.7 0.5

467,000e1,750,000 285,000e1,750,000 467,000e1,750,000

435,108 353,872 435,108

349,467 251,490 3141

Gas flow to AGR, acfm Sulfur Production, lb/hr Fuel gas flow, lb/hr

0.79 0.67 0.78

6000e30,500 300-43,900 185,000e2,490,000

735,757 11,634 248,650

253,996 39,471 2278

Compressor Power, kW

0.88

28,300e43,500

30,210

81,688

HRSG duty, MMBtu/hr volumetric flow to stack, acfm

0.7 0.7

600-5000 1,010,000e2,810,000

1791 2,647,185

42,796 6769

H.-T. Oh et al. / Energy 188 (2019) 116063 Table 4 Standard molar chemical exergy [53]. Substance

State

Standard molar chemical exergy, kJ/mol

Ar CH4 CO CO2 COS H2 H2O H2O H2S N2 NH3 O2 S SO2 C2H4 C2H6

g g g g g g g l g g g g s g g g

11.7 831.2 274.7 19.5 848.8 236.1 9.5 0.9 812.0 0.7 337.9 4.0 609.6 313.4 1360.3 1495.0

flow rate in kJ/mol. h and s are the molar enthalpy and entropy at a given temperature and pressure. T0 , h0 , and s0 are the reference temperature, the molar enthalpy and entropy at reference temperature and pressure, respectively. To evaluate the molar chemical exergy, xi indicates the mole fraction and ε0;i indicates the standard molar chemical exergy of each component i. The values of the standard molar chemical exergy, ε0;i , used in this study are shown in Table 4 [53]. The standard condition (T0 ¼ 298:15 K, P0 ¼ 1:013 bar) was adopted as a reference condition in this study. All the information required to calculate exergy was obtained from performance simulations. 5. Comparative thermodynamic analysis 5.1. Plant performance Before the techno-economic analysis of the overall IGCC with the CCP, the validity of the process simulation was first confirmed. The validation of the simulation was conducted with a comparison with Case B1B of DOE report (Illinois # 6 with CO2 capture) at the same conditions. The results afforded in this study (Table 5) were near-identical to those reported by the NETL [27]. Therefore, further simulation results from this study were expected to be reliable. Since the process parameter for each unit could affect the overall plant efficiency and COE, its dependence of energy efficiency has widely been studied. In this study, all the cases were

7

simulated at the following required conditions regardless of coal type: more than 98% CO conversion efficiency at the WGSR, higher than 99.9% of sulfur removal from the syngas, and greater than 90% carbon capture efficiency at the dual-stage Selexol process. Especially, the condition of sulfur removal is prerequisite in flue gas treatment due to environmental regulations. The variable of the overall IGCC process was the coal type, and other parameters, which were assumed as constant, were listed in Table 2. The grand composite curves (GCCs) present a temperature enthalpy profile of heat supply and demand within the process. Fig. 2 presents the GCCs for Coal I6 and Coal MR as examples of bituminous and sub-bituminous coals as shown in Table 1. Although the heat capacity of the flue gas was changed with temperature (within less than 10%), the HCC line looked straight due to the wide range of enthalpy in Fig. 2. The two composite curves were located in the most adjacent positions and pinch point to achieve the maximum heat recovery. The difference between the two curves was determined by the minimum temperature approach, which was the minimum driving force for heat exchange [54,55]. The result indicated that the volumes of HP, IP, and LP water were adjusted to maximize the CC power output and minimize the difference of pinch point temperature in the process simulation. As presented in Table 1, the weight percent (wt.%) of carbon and hydrogen per unit as-received coal mass was higher in the bituminous coals than in the sub-bituminous coals. Since the HHV per unit as-received sub-bituminous coal was much lower, greater volumes of sub-bituminous coals, 79.94e82.00 kg/s, had to be supplied into the coal dryer than that of bituminous coals, 52.14e58.70 kg/s, for the same thermal input of dried coal to the gasifier. This implies that a larger volume of carbon and hydrogen sources were supplied to the gasifier when the sub-bituminous coals were used. Therefore, syngas with a different composition was produced from the gasifier as shown in Table 5, depending on coal type, and coal type could affect the overall process performance, as shown in Table 6. The higher the HHV of as-received coal fed into the gasifier, the higher the cold gas efficiency of the gasifier and net overall plant efficiency. The gross power of the GT was approximately 440 MW regardless of coal type. To reduce the formation of thermal NOx, IP steam was also used in the GT as a steam diluent until the LHV of syngas (position 3 in Fig. 1) reached 4.4e4.7 MJ/m3. In this study, however, the LHV of syngas was less than 4.7 MJ/m3 through use of N2 dilution from the ASU without steam dilution. Therefore, the IP steam, which was not used for the steam dilution, was applied for

Table 5 Syngas specifications at the outlet of gasifier (position 1 in Fig. 1) according to a coal type.

Temperature [K] Molar Flowrate [kmol/hr] Mole Fraction, % H2 CO CO2 H2O O2 CH4 N2 H2S COS NH3 Cold Gas Efficiency Steam to Coal ratio Oxygen to Coal ratio

Bituminous

Bituminous

Bituminous

Sub-bituminous

Sub-bituminous

DOE Report, COAL_I6

COAL_P8

COAL_I6

COAL_MR

COAL_WY

e 18,754

1872 17,866

1858 18,217

1830 18,253

1823 18,081

30.42 58.64 1.42 2.53 0.00 0.06 5.74 0.83 0.00 0.36 e 0.103 0.797

27.88 59.92 1.60 2.98 0.00 0.03 5.72 0.77 0.00 1.10 0.813 0.100 0.800

27.89 57.47 2.44 4.70 0.00 0.02 5.54 0.91 0.00 1.03 0.808 0.100 0.750

24.93 62.15 2.22 3.59 0.00 0.00 5.95 0.36 0.00 0.80 0.807 0.050 0.744

25.05 61.84 2.35 3.87 0.00 0.00 5.88 0.19 0.00 0.82 0.793 0.000 0.710

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Fig. 2. Grand composite curve for CC with (a) Coal I6 or (b) Coal MR.

the ST, and the power generation in the ST was greater than that reported in the DOE report [54,55]. However, because of the lack of steam dilution, the syngas flow rate was low, and the power generation in GT was lower. As a result, the sum of GT and ST power generation was similar to the result of the DOE report as shown in Table 6. Since a large volume of carbon and hydrogen were produced from the IGCC process in the case of sub-bituminous coal (with greater volumes of coal required), the molar flow rate of H2 (15,305e15,495 kmol/h) was greater than that of bituminous coals (15,189e15,315 kmol/h) as shown in Fig. 3. Therefore, the gross power of the GT in the case of sub-bituminous coals was slightly higher. However, in the case of sub-bituminous coals, the higher asreceived coal flow rate as shown in Fig. 3(b) led to a higher auxiliary load (1 MW) in coal treatment units such as coal handling and

milling. Furthermore, the ST power outputs were significantly different among the coal types as shown in Table 6. The drying energy for each coal differed from 0.2 to 54.4 MJ/s, depending on the moisture content of as-received coals. Coal drying was achieved by recycling a small volume of hot syngas or introducing superheated steam. The gasifier using sub-bituminous coals generated more heat (93.52e98.1 MW) than the gasifier using bituminous coals (84.61e86.17 MW) because the volume of coal was greater. However, since more drying energy was needed for sub-bituminous coals, the amount of heat supplied to the HRSG (39.16e53.71 MW) was less than that of bituminous coals (75.51e84.39 MW). The additional IP steam consumption in the gasifier led to decreasing the power generation of the ST. However, in the case of bituminous coals, the IP steam was added to the gasifier to increase the cold gas efficiency of the gasifier. The

Table 6 Performance summary of IGCC with various types of coal.

Power summary, kW Gas Turbine Steam Turbine Total Power Auxiliary Load summary, kW Coal handling Coal Milling Slag Handling ASU Main Air Compressor Oxygen Compressor Nitrogen Compressor CO2 Compressor Boiler Feedwater Pumps Condensate Pump Quench Water Pump Syngas Recycle Compressor Circulation Water Pump Scrubber Pump Acid Gas Removal Claus TG Recycle Compressor Total Auxiliary Net Power, kW Thermal Input, kW (a. r.) Net Plant Efficiency, %(HHV) Performance Indicator CO2 capture rate, kmol/hr Sulfur production rate, kg/hr Slag generation, kg/hr Raw water consumption rate, gpm

Bituminous

Bituminous

Bituminous

Sub-bituminous

Sub-bituminous

I6, DOE

COAL_P8

COAL_I6

COAL_MR

COAL_WY

464,000 209,400 673,400

443,205 240,111 683,316

439,504 238,036 677,540

449,100 235,057 684,158

443,201 226,995 670,196

460 2170 550 59,740 9460 32,910 30,210 3500 280 610 790 4370 360 18,650 1830 176,540 496,860 1,590,722 31.23

423 1934 430 59,622 9314 34,987 29,871 3337 257 711 1112 3712 778 30,149 2296 178,934 504,549 1,590,722 31.62

452 2172 507 59,010 9200 34,627 29,476 3296 252 725 1109 3699 804 29,878 2320 177,528 500,012 1,590,722 31.35

554 2945 582 61,747 9694 36,233 31,380 3389 252 706 1045 3724 783 30,811 2190 186,035 498,123 1,590,722 31.23

563 3020 409 60,165 9459 35,305 31,321 3339 244 696 1022 3614 775 30,612 2108 182,652 487,544 1,590,722 30.57

10,099 5277 21,137 2826

9826 4348 16,034 2680

9707 5115 19,393 2639

10,422 2008 22,594 2674

10,405 1035 15,101 2584

H.-T. Oh et al. / Energy 188 (2019) 116063

9

Fig. 3. Flowrate of component at the (a) effluent gas of syngas cooling process (position 2 in Fig. 1) and (b) treated gas (position 3 in Fig. 1) with various types of coal.

amount of heat generated in the gasifier was different depending on the coal type while the amount of heat generated in the syngas cooling zone was similar, 113.6e118.5 MW. The WGSR process also generated a lot of heat energy. The larger the flow rate of H2þCO, the more heat energy was generated. Therefore, Coal MR, which had the highest flow rate of H2þCO, generated the highest heat energy, 168.7 MW. However, a large amount of shift steam was also required to convert a large amount of CO to CO2. In other words, the amount of heat energy generated and the amount of steam needed were in a trade-off relationship. As a result, except for the WY coal case, which needed very large drying energy, coal P8, I6, and MR did not show much difference in the power generation of a ST as shown in Table 6. It was also reported that the IGCC using a dry-fed gasifier (Shell gasifier) did not show any significant change in overall plant efficiency by coal type, compared to the results of a slurry-fed gasifier [8,22]. The results confirmed that the greatest auxiliary power consumption resulted from the ASU for all the cases. Since the coal flow rates of sub-bituminous coals were higher at the same HHV input conditions, as shown in Table 1, a larger oxygen supply was required for the gasification of sub-bituminous coal. This resulted in the higher power demand (3 MW) for the ASU auxiliary load. The higher coal flow led to higher carbon molar flow, and the additional 3 MW was needed for removing acid gases and carbon. Sulfur content was much higher in the bituminous coals than in the subbituminous coals, as shown in Table 1. As a result, as shown in Fig. 3 (a), the flow rate of H2S in the outlet syngas of gasifier was much higher in the bituminous coals. Therefore, the Claus process required more auxiliary loads for the bituminous coals. On the other hand, IP and LP steams could be generated through the exothermic reaction of the Claus process. That is, the increase in auxiliary load in the Claus process did not lead to a noticeable difference in net power generation due to the increase in ST power generation. Raw water consumption is one of the important factors used to assess plant performance. As presented in Table 6, unlike the other factors, raw water consumption did not show any significant difference between bituminous and sub-bituminous coals. The case of coal WY showed the lowest raw water consumption, but it was 3.4% less than the case of coal P8 with the highest raw water consumption. In all the cases, the highest volume of water consumption was found in the cooling tower, showing more than 80% of the total raw water consumption. Cooling water was mostly required at the condenser in the ST, accounting for more than 60% of the total cooling energy. The cooling energy required for the condenser was determined by the amount of circulating water in the ST. Therefore, as presented in Fig. 4, the highest cooling energy was required for Coal P8,

1358 gallons per minute (gpm). The WY coal case required the least volume of cooling water for cooling exhaust steam because the amount of steam for the operation of the ST was small. As a result, cooling water consumption in the condenser was the lowest in the WY coal, 1272 gpm. The second largest amount of water consumed occurred at the process makeup. It was necessary to constantly resupply the consumed water into the process, such as shift water in WGSR and IP steam in the gasifier. Because the Coal P8 produced the least amount of water in the outlet syngas from the gasifier in Fig. 3 (a), and the largest amount of shift water, 435 gpm, should be supplied to promote higher conversion of CO. The intercooled compressor also required a large amount of cooling water. The multi-stage compressor was cooled in the middle to prevent the gas temperature from becoming too high through compression, and this required a lot of cooling energy. Since a large volume of dried coal was fed to the process in the subbituminous cases, this required 2e4% higher molar flow rate of air and CO2. Thus, in the sub-bituminous coal cases, the increased compressor capacity led to increased need for cooling energy in the intercooled compressor located in the ASU and the refrigerator of the dual-stage Selexol process.

5.2. Exergy analysis Fig. 5 presents the exergy flow using Grassmann diagrams for

Fig. 4. Summary of raw water consumption at various types of coal (pink: Cooling water-WGSR, green: Cooling water-Selexol, blue: Cooling water-ASU, red: Process Makeup, black: Cooling water-Condenser)

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H.-T. Oh et al. / Energy 188 (2019) 116063

Fig. 5. Exergy flow diagram for IGCC with CCS: (a) Coal P8 and (b) Coal WY.

Coal P8 and Coal WY in Table 1 as an example of bituminous and sub-bituminous coals. The total exergy destructions for Coal P8 and Coal WY were 877.1 MW and 897.0 MW. The highest exergy destruction occurred in GT regardless of the coal type, showing approximately 39.0% of the total exergy destruction. When considering the results of chemical and physical exergy

loss separately from Fig. 5, large exergy destruction occurred in the units where the chemical reaction occurred, such as gasifier, WGSR, and GT, because the chemical reaction caused a large change in chemical exergy. For example, considering only the chemical exergy loss in the Coal WY case, the gasifier, WGSR, and GT combustor showed 183.2 MW, 48.9 MW, and 255.0 MW,

H.-T. Oh et al. / Energy 188 (2019) 116063

respectively. The summation of these chemical exergy losses accounted for 54.3% of the total exergy destruction. With respect to the physical exergy loss, high exergy loss was observed in heat exchange in syngas cooling of 45.4 MW, WGSR of 48.8 MW, HRSG of 20.8 MW, and Cooling tower of 23.8 MW, which accounted for 17.5% of the total exergy destruction. And relatively small physical exergy loss was also observed in compression, expansion, and mixing. When comparing the exergy loss between two coals, the largest was found in the coal feed & gasifier unit as shown in Fig. 5. The Coal P8 (bituminous) case presented 19.2% of the total exergy loss while the Coal WY case (sub-bituminous) showed 23.0% because of the large difference in drying energy between two coals. It indicated that the improvement in the drying process and the utilization of less humid sub-bituminous coal result in possibly producing additional electric energy up to 22.7 MW. In the syngas treatment process, about 9.3e10.6% of the total exergy loss occurred. The bituminous coal with more sulfur led to a larger exergy loss while the exergy loss from the CO2 absorber, refrigerator and CO2 compressor was approximately 60.0 MW regardless of the coal type. 6. Comparative economic analysis The price per coal ton is very different among the types of coal and fluctuates depending on worldwide energy demands. CO2 T&S costs vary greatly, depending on where the IGCC plant is located. In this study, the as-received coal prices were based on the US Energy Information Administration (EIA) documents. (Coal P8: 46.4 $/ton, Coal I6: 39.2 $/ton, Coal MR: 17.9 $/ton, and Coal WY: 13.74 $/ton) [10]. An economic analysis of the IGCC power plant using each coal is presented in Table 7. When the flow rate increased, the device or unit had to also increase in size, and this led to an increase in capital cost. Gasifier and coal feeding systems accounted for more than 50% of capital cost. Therefore, the capital cost of sub-bituminous coal, which had a higher coal flow rate due to lower HHV per as-received coal at the same thermal input, was 8% higher. Especially, the coal drying system cost was 28.8% higher, which showed the most significant difference in the cost comparison among several unit processes. Coal I6 and Coal MR contained a lot of ash in the asreceived coals, which increased slag production rate, and the slag disposal costs were 34% higher as shown in Table 7. However, the fuel cost was the most important factor in determining the operating cost. The fuel price accounted for 88% of the operating cost for the bituminous coals and 79% of the operating cost for the sub-bituminous coals. Generally, fuel prices fluctuate annually. Recently, the price of bituminous coal has been falling, but

11

Fig. 6. Sensitivity response on the COE under variation of fuel price with various types of coal [10].

the price of sub-bituminous coal is increasing. For example, the price of coal P8 fell by 41.7% in 2016, from 2011 prices. On the other hand, the price of coal MR, the sub-bituminous coal, was 15.43 $/ton in 2011, and increased by 16% in 2016. The relation between fuel price and COE result for each coal is presented in Fig. 6. Since the operating cost was highly affected by the fuel price, the COE also fluctuated with the changing fuel price. From 2011 to 2016, the coal price variation was much more significant in bituminous coals than in sub-bituminous coals. Therefore, the COE of the IGCC process using a bituminous coal was greater than that using a subbituminous coal in 2011, but this situation was reversed in 2016 as shown in Fig. 6. Table 7 shows the cost evaluation result of IGCC power plant considering the coal price in 2016. In all the four coal cases, the same transport cost was used (2.24 $/CO2-ton) [27]. The storage cost depends on the storage location [27]. In the case of subbituminous coal, the saline aquifers (Montana Powder River basin and North Dakota Williston basin), which could store CO2, are located deeper than the location for the bituminous coal (Illinois basin). As such, storage costs were increased for the subbituminous cases. As shown in Table 7, excluding CO2 T&S cost, the difference in COE between bituminous and sub-bituminous was within 2.7% (favorable for bituminous coal). In addition, the bituminous coals were more favorable in terms of the initial capital investment than the sub-bituminous coals. Furthermore, when

Table 7 Cost summary of IGCC with various types of coal based on the coal price in 2016. Unit: $/net-MWh

Capital Cost Equipment cost Operating Cost Fuel Water Waste Water disposal Shift Catalyst Selexol Solution Claus Catalyst Slag Disposal Total (Excluding T&S) CO2 T&S Total (Including T&S)

Bituminous

Bituminous

Sub-bituminous

COAL_P8

COAL_I6

COAL_MR

Sub-bituminous COAL_WY

113.31 33.99

114.83 34.45

124.20 37.26

124.55 37.37

19.03 0.51 0.52 0.27 0.27 0.03 0.87 134.81 9.37 144.17

18.26 0.52 0.51 0.27 0.27 0.03 1.06 135.76 9.33 145.09

11.40 0.53 0.52 0.28 0.29 0.01 1.24 138.47 21.94 160.41

9.17 0.53 0.28 0.30 0.30 0.01 0.85 136.21 14.92 151.13

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H.-T. Oh et al. / Energy 188 (2019) 116063

considering the CO2 T&S cost, the COE difference among the coals increased to 11.3% as shown in Table 7. In the study, the COE was evaluated using the scenarios of relative influence on the location of T&S and the coal price from previous studies. Although the Coal WY case was the worst in terms of net power and net plant efficiency as shown in Table 6, the COE including CO2 T&S can be the cheapest when it is evaluated using the coal price at 2011. Furthermore, because CO2 T&S cost is highly affected by the storage location, the cost as shown in Table 7 is indicative. The comparative study for the IGCC process using the different types of coal serves as a guideline for coal selection and implementation of IGCC including CO2 T&S.

location, influences the COE estimation. In addition, although water cost was not significant in terms of the total operating cost in the study, it can be much different, depending on where the power plant is located and on the water price in each country. In this case, the utilization of steam, including the evaporated steam from coal, must be considered. Therefore, the decision on the location of the IGCC power plant can become a critical factor when considering the cost effectiveness of the plant and the reduction of greenhouse gas emissions. Conflicts of interest None.

7. Conclusions The performance analysis of the overall 500 MW-class IGCC plant with CCP was conducted through heat and power integration analysis. The net plant efficiency, raw water consumption, exergy destruction and COE were compared among the IGCC processes using four different commercial coals. Coals with higher moisture content consumed a larger volume of IP steam for drying. On the other hand, with low moisture content coals, the IP steam was consumed in the gasification to increase the cold gas efficiency of the gasifier. Thus, except for the WY coal case, which needed greater drying energy, the other coal types did not show much difference in the ST power generation. However, a larger oxygen supply was required for gasification in the asreceived coal with high moisture content because greater volumes of coal had to be fed into the gasifier under the same HHV input conditions. Therefore, sub-bituminous coals (Coal MR and WY) needed greater auxiliary power consumption than bituminous coals (Coal P8 and I6), and the net power efficiency was relatively lower. Greater volumes of circulating IP steam were consumed for coal drying of as-received coals with higher moisture content. On the other hand, raw water consumption was significantly affected by the amount of circulating water at the condenser in the ST. Additional makeup water was required for WGSR and IP steam in the gasifier. Coal P8 had the lowest moisture content and consumed more raw water than the other coal types, but the difference was not significant. In this study, some basic parameters were set as constant. Therefore, their dependence on energy efficiency is needed for further study. The large exergy destruction occurred at the units, such as gasifier, WGSR, and GT. Especially, the largest loss, 38.9e39.8% of the total exergy destruction, occurred in GT regardless of the coal type. It indicated the importance of developing chemical reaction technologies to save the chemical exergy loss. The total exergy destruction of bituminous coal case was approximately 2.4% lower than that of the sub-bituminous coal case. The highest difference in the exergy destruction between two types of coal occurred in the coal feed & gasifier process due to more drying energy and lower cold gas efficiency of sub-bituminous coals. The capital cost was highly influenced by the coal type at the same HHV input conditions. The capital cost of bituminous coal cases was approximately 10% lower than that of sub-bituminous coal cases because greater volumes of sub-bituminous coal had to be fed into the gasifier and the molar flow rate of stream became larger. The IGCC power plants using bituminous coals were advantageous in terms of operating energy efficiency. However, in terms of total operating cost, coal price was the most important factor. Based on the coal price in 2016, bituminous coals are more cost-effective in IGCC power plants. However, because the coal price fluctuates significantly annually (especially bituminous coals), stable cost estimation is very difficult when coal price is considered. The cost of CO2 T&S, which is based on the geological storage

Acknowledgment This work was supported by the Industrial Strategic Technology Development Program-Engineering Core Technology Development Project (10077467) and Korea Evaluation and Planning (20172010202070) funded by the Ministry of Trade, Industry & Energy (MOTIE, Korea). References [1] Conti J, Holtberg P, Diefenderfer J, LaRose A, Turnure JT, Westfall L. International energy outlook 2016 with projections to 2040. Washington, DC (United States): US DOE Energy Information Administration (EIA); 2016. Office of Energy Analysis. [2] Spliethoff H. Power generation from solid fuels. Springer Science & Business Media; 2010. [3] Martelli E, Kreutz T, Carbo M, Consonni S, Jansen D. Shell coal IGCCS with carbon capture: conventional gas quench vs. innovative configurations. Appl Energy 2011;88:3978e89. [4] Herron JA, Kim J, Upadhye AA, Huber GW, Maravelias CT. A general framework for the assessment of solar fuel technologies. Energy Environ Sci 2015;8: 126e57. [5] International Energy Agency. Statistics I. CO2 emissions from fuel combustionhighlights. Paris: IEA; 2017. http://www.iea.org/co2highlights/co2highlights/. [6] Strube R, Manfrida G. CO2 capture in coal-fired power plants Impact on plant performance. Int J Greenh Gas Contr 2015;5:710e26. [7] Erlach B, Schmidt M, Tsatsaronis G. Comparison of carbon capture IGCC with pre-combustion decarbonisation and with chemical-looping combustion. Energy 2011;36:3804e15. [8] Majoumerd MM, Raas H, De S, Assadi M. Estimation of performance variation of future generation IGCC with coal quality and gasification processeSimulation results of EU H2-IGCC project. Appl Energy 2014;113: 452e62. [9] British Petroleum. BP statistical review of the world energy British petroleum. 2012. [10] Energy Information Administration. Annual coal report. https://www.eia.gov/ coal/annual/. [11] Promes E, Woudstra T, Schoenmakers L, Oldenbroek V, Thattai AT, Aravind P. Thermodynamic evaluation and experimental validation of 253 MW integrated coal gasification combined cycle power plant in Buggenum, Netherlands. Appl Energy 2015;155:181e94. [12] Ryzhkov A, Bogatova T, Gordeev S. Technological solutions for an advanced IGCC plant. Fuel 2018;214:63e72. [13] Zhu X, Shi Y, Cai N. Integrated gasification combined cycle with carbon dioxide capture by elevated temperature pressure swing adsorption. Appl Energy 2016;176:196e208. [14] Lee WS, Oh HT, Lee JC, Oh M, Lee CH. Performance analysis and carbon reduction assessment of an integrated syngas purification process for the coproduction of hydrogen and power in an integrated gasification combined cycle plant. Energy 2019;171:910e27. [15] Cormos AM, Dinca C, Cormos CC. Multi-fuel multi-product operation of IGCC power plants with carbon capture and storage (CCS). Appl Therm Eng 2015;74:20e7. [16] Cormos CC. Integrated assessment of IGCC power generation technology with carbon capture and storage (CCS). Energy 2012;42:434e45. [17] Chen C, Rubin ES. CO2 control technology effects on IGCC plant performance and cost. Energy Policy 2009;37:915e24. [18] Ju YS, Lee CH. Evaluation of the energy efficiency of the shell coal gasification process by coal type. Energy Convers Manag 2017;143:123e36. [19] Ju YS, Lee HH, Oh M, Lee CH. Performance analysis on an entrained-flow gasifier by coal moisture. Chem Eng Technol 2017;40:2257e65. [20] Karthikeyan M, Zhonghua E, Mujumdar AS. Low-rank coal drying technologies Current status and new developments. Dry Technol 2009;27:403e15.

H.-T. Oh et al. / Energy 188 (2019) 116063 [21] Field RP, Brasington R. Baseline flowsheet model for IGCC with carbon capture. Ind Eng Chem Res 2011;50:11306e12. [22] Botero C, Field RP, Brasington RD, Herzog HJ, Ghoniem AF. Performance of an IGCC plant with carbon capture and coal-CO2-slurry feed: impact of coal rank, slurry loading, and syngas cooling technology. Ind Eng Chem Res 2012;51: 11778e90. [23] Kim YS, Lee JJ, Kim TS, Sohn JL. Effects of syngas type on the operation and performance of a gas turbine in integrated gasification combined cycle. Energy Convers Manag 2011;52:2262e71. [24] Lee JJ, Kim YS, Cha KS, Kim TS, Sohn JL, Joo YJ. Influence of system integration options on the performance of an integrated gasification combined cycle power plant. Appl Energy 2009;86:1788e96. [25] Gazzani M, Macchi E, Manzolini G. CO2 capture in integrated gasification combined cycle with SEWGS e Part A: thermodynamic performances. Fuel 2013;105:206e19. [26] Majoumerd MM, De S, Assadi M, Breuhaus P. An EU initiative for future generation of IGCC power plants using hydrogen-rich syngas: simulation results for the baseline configuration. Appl Energy 2012;99:280e90. [27] Fout T. Cost and performance baseline for fossil energy plants volume 1b: bituminous coal (IGCC) to electricity revision 2b e year dollar update. Washington, DC, USA: National Energy Technology Laboratory; 2015. [28] De S, Biswal S. Performance improvement of a coal gasification and combined cogeneration plant by multi-pressure steam generation. Appl Therm Eng 2004;24:449e56. [29] Mansouri MT, Ahmadi P, Kaviri AG, Jaafar MNM. Exergetic and economic evaluation of the effect of HRSG configurations on the performance of combined cycle power plants. Energy Convers Manag 2012;58:47e58. [30] Adams T, Dowell TM. Off-design point modelling of a 420 MW CCGT power plant integrated with an amine-based post-combustion CO2 capture and compression process. Appl Energy 2016;178:681e702. [31] Zhang G, Zheng J, Yang Y, Liu W. Thermodynamic performance simulation and concise formulas for triple-pressure reheat HRSG of gasesteam combined cycle under off-design condition. Energy Convers Manag 2016;122:372e85. [32] Nadir M, Ghenaiet A, Carcasci C. Thermo-economic optimization of heat recovery steam generator for a range of gas turbine exhaust temperatures. Appl Therm Eng 2016;106:811e26. [33] Chou V, Keairns D, Turner M, Woods M, Zoelle A. Quality guidelines for energy system studies: process modeling design parameters. US Department of Energy; 2014. Tech. Rep. DOE/NETL-341/051314. [34] Osman H, Jangam S, Lease J, Mujumdar AS. Drying of low-rank coal (LRC) A review of recent patents and innovations. Dry Technol 2011;29:1763e83. [35] Jones D, Bhattacharyya D, Turton R, Zitney SE. Optimal design and integration of an air separation unit (ASU) for an integrated gasification combined cycle (IGCC) power plant with CO2 capture. Fuel Process Technol 2011;92:1685e95. [36] Fu Q, Song C, Liu Y, Ishizuka M, Tsutsumi A. An advanced cryogenic air separation process based on self-heat recuperation for CO2 separation. Energy Procedia 2014;61:1673e6.

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[37] Lee WS, Lee JC, Oh HT, Baek SW, Oh M, Lee CH. Performance, economic and exergy analyses of carbon capture processes for a 300 MW class integrated gasification combined cycle power plant. Energy 2017;134:731e42. [38] McCoy ST, Rubin ES. Models of CO2 transport and storage costs and their importance in CCS cost estimates. In: Proceedings of the 4th annual conference on carbon capture and sequestration. DOE/NETL; 2005. [39] McCollum DL, Ogden JM. Techno-economic models for carbon dioxide compression, transport, and storage & correlations for estimating carbon dioxide density and viscosity. Davis: Institute of Transportation Studies, University of California; 2006. [40] Kaviri AG, Jaafar MNM, Lazim TM. Modeling and multi-objective exergy based optimization of a combined cycle power plant using a genetic algorithm. Energy Convers Manag 2012;58:94e103. [41] Bassily A. Modeling, numerical optimization, and irreversibility reduction of a dual-pressure reheat combined-cycle. Appl Energy 2005;81:127e51. [42] Bassily A. Modeling, numerical optimization, and irreversibility reduction of a triple-pressure reheat combined cycle. Energy 2007;32:778e94. [43] Carapellucci R, Giordano L. A comparison between exergetic and economic criteria for optimizing the heat recovery steam generators of gas-steam power plants. Energy 2013;58:458e72. [44] Morvay Z, Gvozdenac D. Applied industrial energy and environmental management, Part III: fundamentals for analysis and calculation of energy and environmental performance. New York: Wiley-IEEE Press; 2008. [45] Lee HH, Lee JC, Joo YJ, Oh M, Lee CH. Dynamic modeling of Shell entrained flow gasifier in an integrated gasification combined cycle process. Appl Energy 2014;131:425e40. [46] Asif M, Bak MU, Saleem MW, Kim WS. Performance evaluation of integrated gasification combined cycle (IGCC) utilizing a blended solution of ammonia and 2-amino-2-methyl-1-propanol (AMP) for CO2 capture. Fuel 2015;160: 513e24. [47] Aspen Technology Inc. Aspen physical property methods. Cambridge, MA, USA: Aspen Technology Inc; 2006. pp. 61e63. [48] Haar L. NBS/NRC steam tables. Boca Raton: CRC Press; 1984. [49] Black J. Quality guidelines for energy systems studies: capital cost scaling methodology. 2013. DOE/NETL-341/013113/. [50] Jeong YS, Jung J, Lee U, Yang C, Han C. Techno-economic analysis of mechanical vapor recompression for process integration of post-combustion CO2 capture with downstream compression. Chem Eng Res Des 2015;104:247e55. [51] IHS Markit. Engineering costs Index. 2016. www.ihsindexes.com/. [52] Sato N. Chemical energy and exergy: an introduction to chemical thermodynamics for engineers. Amsterdam: Elsevier; 2004. [53] Szargut J. Exergy method: technical and ecological applications. Southampton: WIT press; 2005. [54] Yoon SG, Lee J, Park S. Heat integration analysis for an industrial ethylbenzene plant using pinch analysis. Appl Therm Eng 2007;27:886e93. [55] Smith R. State of the art in process integration. Appl Therm Eng 2000;20: 1337e45.