Comparison of carbon capture IGCC with chemical-looping combustion and with calcium-looping process driven by coal for power generation

chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 110–124

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Chemical Engineering Research and Design journal homepage: www.elsevier.com/locate/cherd

Comparison of carbon capture IGCC with chemical-looping combustion and with calcium-looping process driven by coal for power generation Lin Zhu, Peng Jiang, Junming Fan ∗ Key Laboratory of Gas Process Engineering, School of Chemistry and Chemical Engineering, Southwest Petroleum University, Chengdu 610500, PR China

a r t i c l e

i n f o

Article history:

a b s t r a c t Three types of coal based integrated gasification combined cycle (IGCC) systems with CO2

Received 25 January 2015

capture using physical absorption, chemical looping combustion (CLC), and calcium loop-

Received in revised form 7 July 2015

ing process (CLP) for power generation are modeled using Aspen Plus. The effects of key

Accepted 24 July 2015

variables on the thermodynamic performance, such as the energy efficiency and the exergy

Available online 31 July 2015

efficiency, are investigated separately. The process variables examined are mainly in the gasification unit, they are steam to coal mass ratio (S/C) in the range of 0–20% and oxygen to

Keywords:

coal mass ratio (O/C) in the range of 70–105%. The performances of the above three capture

Technical comparison

technologies are compared with respect to their energy and exergy efficiencies. An IGCC

IGCC

plant without carbon capture is also considered as a benchmark to quantify the efficiencies.

Chemical looping combustion

The results show that the CLC technology has an energy efficiency of 39.78%, which is 2.06%

Calcium looping process

and 3.57% higher than the CLP and physical absorption-based technologies, respectively.

Power generation

Additionally, the CLC technology has an exergy efficiency of 35.67%, in contrast to 33.86%

Carbon dioxide capture

for CLP and 32.74% for physical absorption technology. Furthermore, the economic evaluations are performed for estimation of capital cost, net present value (NPV), internal rate of return (IRR) and other economic values. © 2015 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved.

1.

Introduction

Coal plays an important role in the world energy consumption structure. Statistically, it accounts for 28.8% of the world energy supply and generates 41.3% of the world’s electricity (IEA, 2013).An estimation is that coal, as a primary source of energy, is irreplaceable for at least the next 5–6 decades (Armor, 2014). However, coal is the most carbon-intensive fossil fuel and contributes the largest share of carbon dioxide (CO2 ) emission from fuel burning (42.9%) (Erlach et al., 2011; IEA, 2013). Carbon dioxide capture and storage (CCS) is a promising way for coal-fired plants to achieve significant CO2

∗

emission reductions. Currently, the CO2 -cutting technologies in coal-based power plants include three approaches, namely post-combustion systems, oxy-combustion systems, and precombustion systems (Fan et al., 2012). A post-combustion system directly separates CO2 from the flue gas after combustion. In an oxy-combustion system, fuel combustion takes place using pure oxygen instead of air, generating a sequestrable CO2 stream (Adanez et al., 2012). In a pre-combustion decarburization process, CO2 is captured prior to fuel combustion. The chemical looping combustion (CLC) process and calcium sorbent-based looping process (CLP) have been developed in recent years as a variety of post-combustion,

Corresponding author. Tel.: +86 28 8303 7304. E-mail address: [email protected] (J. Fan). http://dx.doi.org/10.1016/j.cherd.2015.07.027 0263-8762/© 2015 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved.

chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 110–124

Nomenclature AGR Al2 O3 ASPEN ASU CaCO3 CaO CCS CH4 CLC CLP CO CO2 COS Cu Excoal Fe GT h H2 H2 O H2 S HP HRSG I IRR IGCC HHVcoal LP mcoal MDEA Mn MP N2 Ni NiO NPV O/C O2 O&M Qin R S/C ST WGS

acid gas removal aluminum oxide advanced system for process engineering air separation unit calcium carbonate calcium oxide capture and storage methane chemical looping combustion calcium looping process carbon monoxide carbon dioxide carbon oxysulfide copper the chemical exergy of coal (kJ/kg) iron gas turbine enthalpy (kJ/kmol) hydrogen water hydrogen sulfide high pressure (bar) heat recovery steam generator the irreversibility generated upon the destruction or dissipation of exergy (MW) internal rate of return integrated gasification combined cycle higher heating value of the coal (MJ/kg) low pressure (bar) mass flow of coal feeding to the gasifier (kg/s) methyl di-ethanol amine manganese medium pressure (bar) nitrogen nickel nickel oxide net present value oxygen to coal mass ratio oxygen operational and maintenance total input heat of plant(MW) represents the universal gas constant, R = 8.314 kJ/(kmol K) steam to coal mass ratio steam turbine water gas shift

Greek letters exergy efficiency of system ex net net energy efficiency of system CO2 capture efficiency of system CO2

oxy-combustion, and pre-combustion with CO2 capture to improve the thermodynamic performance of the power plant (Anheden and Svedberg, 1998; Dean et al., 2011). CLC is an indirect fuel combustion strategy by which the CO2 is inherently separated from the other flue gas components. CLC can be classified as a pre-combustion capture technology as the CO2 in the syngas has been captured prior to combustion (Boot-Handford et al., 2014; Fan, 2011). CLP is a

111

chemical-looping-based process, using CaO as a regenerable sorbent of CO2 , and is viewed as pre-combustion capture technology for CO2 . The CO2 is captured before combustion, when it is combined with a water gas shift (WGS) reaction via the sorbent-enhanced reforming process (Connell et al., 2013; Fan, 2011). The reaction scheme for CLC generally involves two main reactions which are carried out in two separate reactors (Bhavsar et al., 2014; Fan and Zhu, 2015). The schematic process diagram of CLC is illustrated in Fig. 1. In the fuel reactor (FR), the fuel syngas is converted completely into CO2 and H2 O, while the oxygen carrier particle (here NiO) is reduced to its elementary state. As shown in Fig. 1, an almost pure CO2 stream from the FR is easily obtained when H2 O(g) is condensed. Meanwhile, the reduced oxygen carrier Ni is then transferred to the AR where it is regenerated to its original state (NiO) through contact with air. Notably, the heat derived from the two reactors is the same as it would be for a normal combustion reaction where the air directly contacts the syngas. Lewis and Gilliland first proposed the concept of CLC in the 1950s, when they put forward the idea of using the reaction of copper oxide with syngas to produce carbon dioxide (Lewis and Gilliland, 1954). Since then, research on CLC has primarily focused on three aspects: process design, reactor design, and oxygen carrier investigation (Ishida and Jin, 1994; Knoche and Richter, 1968). Recently, studies on CLC have mainly concentrated on the development of cyclic chemical intermediates and on their applications in CLC processes with respect to gaseous fuels, while solid fuels (such as coal) also have been studied for their use in the CLC system (Moghtaderi, 2011). For instance, Mattisson et al. investigated 300 different oxygen carriers based on oxides of metals and selected three carriers based on Ni, Fe and Mn to develop using syngas in a CLC system (Mattisson et al., 2007). Leion et al. experimented with solid fuel using a chemical-looping combustion system in a laboratory fluidized bed reactor, as well as in a small pilot plant (Leion et al., 2008; Lyngfelt, 2014). Guío-Pérez conducted a research on the solids residence time distribution in the fuel reactor of a dual circulating fluidized bed system with ringtype internals (Guío-Pérez et al., 2014). These studies confirm the feasibility of using the chemical looping combustion for power production. Similarly, for pre-combustion capture configuration, the calcium looping process consists of two reactors: the carbonator and the calciner (Fig. 2). In the carbonator, CaCO3 is formed from the reaction of CaO with CO2 , thereby removing CO2 products and also enhancing the water gas shift reaction for production of H2 according to Le Chatelier’s principle (Cormos and Cormos, 2013; Zhu and Fan, 2014). In the calciner, CaCO3 is decomposed into CaO, resulting in the production of concentrated CO2 . An overview is given by Dean et al. (Dean et al., 2011). An important feature of the calcium looping process is that calcium compounds are inexpensive and that after a fixed number of cycles, the waste from the reactors can become a raw material for cement. Another benefit of this cycle is that it can be used for CO2 absorption at a large industrial (Cormos and Cormos, 2013). Integrated gasification combined cycle (IGCC) plants are one of the present research focuses in the energy industry, as they enable the usage of coal-to-electricity conversion with CO2 removal for high efficiency power production in a combined cycle unit (Skorek-Osikowska et al., 2012). CLC and CLP processes can be deftly coupled with IGCC for power generation while also capturing CO2 . Rezvani et al.

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Depleted Air

CO2,H2O

Condenser

CO2

H2O Air reactor (AR) Ni+1/2O2

NiO

Fuel reactor (FR)

NiO

NiO+CO NiO+H2

Ni

Ni+CO2 Ni+H2O

Syngas

Air

Fig. 1 – The schematic process diagram of chemical looping combustion (here NiO as metal oxide). techno-economically compare coal-based IGCC systems with CO2 capture using physical absorption, membrane reactors and chemical looping (Rezvani et al., 2009). Erlach et al. compared carbon capture IGCC with pre-combustion decarburization by physical absorption and with chemical looping combustion, and concluded that CLC-IGCC offers the advantages of higher plant efficiency(37.7%) and more complete carbon capture than conventional pre-combustion with 34.9% plant efficiency of physical absorption (Erlach et al., 2011). Further work by Cormos presented a detailed methodology to access the performance of an iron-based CLC system using critical design factors such as gasifier selection and overall energy efficiency. This work suggested that a CLC-based IGCC system has a low energy penalty and has an overall efficiency up to 64.1%, which is higher than an entrained flow gasifier in an IGCC, with an overall efficiency of 46.6% (Cormos, 2012a). Kunze et al. conducted an exergy analysis of an IGCC plant and showed that 60% of the exergy was lost in the overall process, with 17.1% in gas treating and 11.3% in the gasifier (Kunze et al., 2011). Chen et al. developed a CaO sorption-enhanced looping process that was incorporated into a coal-driven IGCC for power production with CO2 separation. They found that an IGCC with CLP process had satisfactory system performance and a net electricity efficiency of approximately 30–33%. The system had a CO2 capture efficiency up to 97%. Moreover, compared with the conventional physical absorption IGCC process, the schematic diagram of a CLP-based was simplified (Chen et al., 2012). Connell et al. performed a preliminary technoeconomic analysis and indicated that using CLP-based IGCC technology contributed a 9–12% cost reduction in electricity production when compared to the conventional physical absorption CO2 capture technology (Connell et al., 2013). Although the aforementioned studies have evaluated a series of power generation plants with CCS, their assumptions and feedstock, etc., are scarcely consistent, rendering comparisons infeasible. The purpose of this work is to develop a methodology to evaluate and compare three different CO2 capture technologies from both thermodynamic

Concentrated CO2

H2-rich gas

Carbonator CO+H2O CO2+H2 (WGS) CaCO3 CaO+CO2 Syngas

Steam

CaCO3 CaO

Calciner CaO+CO2 CaCO3

Heat

Fig. 2 – The schematic process of the proposed CO2 capture process based on carbonation/calcination loop.

and economic perspectives. The processes described here include the following: (i) IGCC with physical absorption carbon capture, (ii) IGCC with CLC carbon capture, and (iii) IGCC with CLP pre-combustion carbon capture. This article also adopted IGCC without carbon capture as benchmark option. This work firstly describes and simulates the physical absorption-based, CLC-based and CLP-based IGCC technologies. The net energy and exergy efficiencies are separately defined and presented in Section 2. In Section 3, sensitivity analysis of the investigated system are carried out to examine the effect of steam coal mass ratio (S/C) and oxygen carbon mass ratio (O/C) of the gasifier on the overall process thermodynamics performance with respect of energy and exergy efficiencies. Based on that, Grassmann diagram is employed to indicate exergy flow. Section 4 carried out economic analysis of the researched systems in order to determine its feasibility. Section 5 ends up with a conclusion.

2.

Methodology

2.1.

Process description

Three different configurations of coal gasification-based power plants have been demonstrated. They are as follows: IGCC with pre-combustion using a physical absorption-based CO2 capture technology, IGCC coupled with a CLC carbon capture process, and IGCC integrating CLP process. The chemical process simulations are conducted in Aspen Plus.

2.1.1.

IGCC with the physical absorption-based process

An IGCC system with the physical absorption-based process is demonstrated in Fig. 3. The plant includes four sections: the air separation unit (ASU), gasification island, power island, and CO2 capture and sequestration. The ASU is used to produce high purity oxygen through cryogenic separation for gasification process. The coal is crushed and fed under pressure to the Shell gasifier, where coal is pyrolyzed and then reacts with oxygen and steam. The gas obtained at the outlet of the gasifier is quenched by recycling cool syngas. The quenched gas is routed to a heat exchanger for cooling, raising intermediate pressure steam. After the temperature of syngas drops, the syngas is divided into two parts: one part is recompressed and recycled to the gasifier for the cooling of raw gas, the other moves down to the next water gas shift unit. In the WGS reactor, steam is injected into the reactor to promote the production of H2 (CO + H2 O = CO2 + H2 ). The H2 -rich gas products are then cooled for the acid gas removal unit (AGR). Sulfur components and CO2 are captured using a Selexol solvent, which is more for energy-efficient than other physical and chemical solvents such as MDEA and

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Coal

Gas Quench and Cooling

Coal Preparation

Acid Gas To Clause plant Removal

Syngas Cooling

Water Gas Shift

CO2

Gasifier

O2

Compression

Steam

Steam

Slag

Steam

N2

CO2

Syngas Reheat Steam

ASU

Steam

Air

CO2 to storage

Flue Gas

Air Compression

Exhaust

Heat Recovery Steam Generator

G

Gas Turbine Air

LP

HP

MP

G

Steam Turbine

Fig. 3 – IGCC with pre-combustion using physical absorption-based CO2 capture technology.

chilled methanol (Padurean et al., 2012). The H2 S concentrated stream is converted into elemental sulfur as a byproduct using a Clause plant. The CO2 -rich gas is compressed to 110 bar via a three-stage intercooling compressor and delivered to a pipeline for transportation. The H2 -rich clean syngas is reheated and diluted (to reduce the formation of NOx ) with nitrogen supplied by the ASU, and then fed to the gas turbine (GT). The exhaust gas from the GT is transferred to the HRSG for the purpose of recovering heat with the production of different pressure levels of steam to drive a steam turbine.

2.1.2.

IGCC coupled with CLC

The IGCC coupled with a CLC cycle is depicted in Fig. 4. From Fig. 4, it is evident that the gasification, gas quench and cooling are the same as in the physical absorptionbased technology. The ash-free gas is sent to a desulfurization unit to strip the H2 S using a Selexol solvent. The low sulfur

Coal

Sulfur Removal

Steam CO2 Turbine Steam

Slag

Steam

O2

Steam

ASU

To Clause plant

Syngas Cooling

Gas Quench and Cooling

Coal Preparation Gasifier

Air

syngas, mainly consisting of H2 and CO, is fed to CLC system. Notably, oxygen carrier selection for the CLC is important. As alluded to earlier, an oxygen carrier particle can be made of metal or metal oxides which have excellent mechanical strength and high temperature stability in addition to having good oxygen-carrying capacity (Yahom et al., 2014). The oxygen carrier considered in this process is a nickel based material consisting of NiO (60%) and Al2 O3 (40%) for its superior thermal stability and reactivity (Dueso et al., 2012). In a CLC system, the syngas is converted completely into H2 O and CO2 upon contacting of NiO, and the reduced Ni particles leaving the FR are then fed to the AR to undergo an oxidization process in the presence of air. The depleted air produced from the AR runs through the air turbine for power production. The air turbine exit gas with high temperature is sent to the HRSG for heat recovery and to raise steam before it is released to the environment. Simultaneously, the CO2 -rich stream from the FR passes through the CO2 turbine for its

CLC

Air Turbine

Power transmission

Air Compression

Flue Gas

Air Heat Recovery Steam Generator CO2 Condenser LP

MP

HP

G Water

Steam Turbine

Fig. 4 – IGCC integrated CLC CO2 capture technology.

CO2 Compression

CO2 to storage

G

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Table 1 – Physical and chemical properties of Illinois #6 coal. Proximate analysis Moisture Ash Fixed carbon Volatiles HHV (MJ/kg)

Wt% (dry)

Ultimate

Wt% (dry)

Sulfate analysis

Wt% (dry)

0 10.91 49.72 39.37 30.53

C H O N S Cl

71.72 5.06 7.75 1.41 2.82 0.33

Pyritic Sulfate Organic

1.70 0.02 1.10

relatively high temperature, then goes to the HRSG and last to the CO2 condenser. After that, the separated CO2 is compressed via a three-stage gear compressor and sent to the pipeline for transportation.

2.3.

Plant thermodynamic performance evaluation

The plant energy efficiency, net (%), is defined as Eq. (1). Wnet × 100% mcoal · LHVcoal

net =

2.1.3.

IGCC coupled with CLP

Fig. 5 presents a schematic diagram of the IGCC with the CLP pre-combustion carbon capture process. An expander is added after the gas quench and cooling unit. The depressurized gases, along with high-pressure steam, are directly injected into the calcium looping process. In the CLP system, CaO, acting as the sorbent from the calciner, simultaneously reacts with CO2 to form CaCO3 , resulting in a high purity hydrogen stream via the enhancement of the water shift reaction. CaO can also be used to remove both H2 S and COS. The H2 -rich stream from the CLP system is sent to a gas turbine for electricity generation. To lower the combustion temperature (avoiding NOx formation by reacting the N2 in gasification gas with oxygen), and to increase the gas flow into the GT, a stream of N2 from the ASU unit is conveyed to the GT. The exhaust gas with high temperature from the GT is forced through the HRSG to recover heat for power generation in the steam system. Heat required to decompose CaCO3 is supplied by the combustion of extra coal and O2 . The concentrated CO2 from the calciner runs through the HRSG to release heat before it is compressed to 110 bar.

2.2.

Modeling assumption and simulation

Illinois #6 coal is selected as the feed stock, and its physical and chemical properties are given in Table 1. The PR-BM equation is selected as the property calculation method for the globe system. STEAM-TA is employed to calculate the thermodynamic properties of steam stream (Kang et al., 2014). The components such as H2 , N2 , CH4 , O2 , CO, CO2 , etc., in the simulation are considered to be of conventional type. Coal and ash are set as nonconventional components and their enthalpy and density are calculated by HCOALGEN and DCHARIGT correlations, respectively. To simplify and run these models, the process parameters and basic assumptions are shown in Table 2 (Chen et al., 2012; Cormos, 2010; Guo et al., 2012; Kunze and Spliethoff, 2012). The RGibbs reactor block, mainly dealing with those reactions with unavailable kinetic equations, is adopted for the calculation of reactions involving solids, including the gasifier, the fuel reactor, the air reactor, the carbonation reactor, and the calciner reactor (Zheng et al., 2013). The Ryield block is used to decompose the selected coal into its conventional components by means of a FORTRAN calculator according to the coal’s proximate and ultimate analysis. The water shift reactor is modeled using a REquil block. MCOMPR models the three-stage CO2 compression process, and the HRSG is modeled with a MHeatX block. All of the performed simulations are carried out in the steady state.

(1)

where Wnet is the plant net power production (MW), mcoal represents the coal mass flow that injected to the gasifier (kg/s), and LHVcoal denotes the lower heating value of the coal (MJ/kg). The plant exergy efficiency, ex (%), is defined as Eq. (2). In this work, the streams such as flue gas, slag, etc. do not have the ability to work for other processes, their exergy are not included in output exergy value (Li et al., 2014). Szargut’s environment condition model is chosen as the basic reference for the component’s chemical exergy calculation (Szargut, 1980). The detailed exergy analyses see the following references (Kawabata et al., 2013; Li et al., 2014). ex =

Wnet × 100% Excoal

(2)

where Wnet is the plant net power production (MW), Excoal is the exergy of the input coal (MW), which is obtained by the correlation conducted by Song et al. (2012). In order to obtain the exergy destruction or loss of each unit, according to the second law of thermodynamics, the exergy balance for a unit is expressed as follows in Eq. (3).



Exin +

 where

 

ExQ =



Exout + I

(3)



Exin , ExQ are the total input exergy (MW). Exout is the output stream exergy (MW). I represents the irreversibility generated upon the destruction or dissipation (MW). To access the CO2 capture efficiency of each system, CO2 (%), is defined as Eq. (4). CO2 =

mCO2 × 100% mCO2 ,total

(4)

where mco2 is the CO2 flow for transportation (kg/s), and mCO2,total denotes the total CO2 flow of the process (kg/s).

3.

Thermodynamics analysis

In this part, the process variables S/C and O/C are defined, and the sensitivity analyses of those variables on the system thermodynamic performance are carried out. Then, the detail of each studied plant from the point of energy and exergy are investigated. This section finishes with Grassmann diagrams of CLC- and CLP-based technologies.

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O2 Coal

CLP

Gas Quench and Cooling

Coal Preparation

Steam

O2

Slag

Expander

Steam

N2

Steam

ASU

Steam

Air

CO2

Gasifier

H2-rich gas

Coal

Air

G

Compression

Flue gas

Heat Recovery Steam Generator

Gas Turbine Air

LP

HP

MP

CO2

G

Compression

CO2 to storage

Steam Turbine

Fig. 5 – IGCC with CLP pre-combustion carbon capture process.

Parametric variables

In the coal gasification section, the gasification temperature, oxygen to coal mass ratio, and steam to coal mass ratio are the most important factors affecting raw gas composition and system efficiency. In this paper, the gasification temperature of each studied object is assumed to be fixed, and therefore, the other two parameters are the investigated variables. On the one hand, the mass ratio of oxygen to coal directly affects the gasification temperature and furthers the output flow of steam; on the other hand, the addition of steam to the gasifier has a significant effect on the output of electricity. Both of the two factors can be investigated by defining the mass ratio of oxygen or steam added in the gasifier to coal as a reactant for gasification, which are expressed as follows. O/C =

moxygen × 100% mcoal

(5)

S/C =

msteam × 100% mcoal

(6)

where moxygen is the mass flow rate of oxygen added to the gasifier (kg/s). Where msteam denotes the mass flow rate of medium pressure steam added to the gasifier (kg/s).

3.2.

Parameter analysis

3.2.1. Sensitivity analysis for the physical absorption-based process The S/C ratio is one of the vital parameters that influence the total heat output, syngas compositions and other important indicators of gasifier quantities. In this work, the S/C ratio is varied from 0 to 20% by adjusting the flow rate of steam and the ratio of O/C is kept at a conventional value of 86.7%. The effect of the S/C ratio on the properties of the gasifier product syngas is displayed in Fig. 5. With the increase in S/C ratio, the concentration of H2 increases from 28.4% to 29.1%, while the CO concentration decreases and the total syngas (CO + H2 ) content also declines. The increase in S/C means that more medium pressure steam from the medium pressure turbine will be injected into the gasifier, and this will directly cause a

decrease in electricity output from the MP turbine. Meanwhile, more steam is fed to the gasifier, suggesting that the heat output from the gasifier will gradually decrease when the input flow rate of oxygen is fixed, and this will also impact the flow rate of steam in the HRSG. When the amount of steam exceeds a certain limit, the gasifier cannot maintain its heat balance and needs an external heat supply, and the total syngas concentration will be reduced. This reduction is not conducive to the gasification process. The effect of the S/C ratio on the plant energy efficiency (net ) and exergy efficiency (ex ) is shown in Fig. 6. With the addition of S/C, the net power efficiency and exergy efficiency decrease separately from 37.37% to 36.20% and from 33.66% to 32.61%, whereas CO2 capture efficiency increases from 85.67% to 88.59% within the 0 to 20% range. As S/C rises, steam employed for power generation in ST from HRSG diminishes, which causes the decrease of power production. As the CO2 capture efficiency increases, more CO2 is captured, which results in the addition of energy consumption in the AGR system for CO2 separation, drying and compression. The energy that the ASU and oxygen compression needed is a

100 Composition of the raw gas/(vol,%)

3.1.

80

60

CO CO2

40

H2 CO+H2

20

0

0

5

10 S/C ratio /(wt,%)

15

20

Fig. 5 – The impact of S/C ratio on the composition of raw gas.

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chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 110–124

40

ηex 38

87

36

84

34

81

32

78

30

0

5

10 15 S/C ratio /(wt,%)

ηnet

90

20

90

38 88

36 86

34

84

32

30

75

ηCO2

ηex

82

70

75

80

85 90 95 O/C ratio /(wt,%)

100

105

CO2 capture efficiency/%

ηCO2

Energy and exergy efficiency/%

ηnet

CO2 capture efficiency/%

Energy and exergy efficiency/%

40

80

Fig. 6 – Effect of S/C on the efficiencies of physical absorption-based process.

Fig. 8 – Effect of O/C on the efficiencies of physical absorption-based process.

constant for the oxygen flow rate injected into the gasifier. The power consumption for gasification island coal preparation mainly includes the recycle gas compression power for the gas quench and cooling units, which has a slow increase when S/C rises. For the power island section, the energy consumption is constant. As a result, based on the above analysis and combined with Eq. (1), it can be easily inferred that the plant net power production will increase. With regards to the exergy efficiency, the calculated exergy value of coal is 739.51 MW, which is greater than the coal thermal energy based on its lower heating value (i.e., 663.58 MW), and from the exergy definition expression (Eq. (2)) with the above analysis, the efficiency of exergy is less than net power efficiency and has the same trend as the former. The mass ratio of oxygen to coal is another core operation condition, which will also affect the composition of gasification raw gas and the heat output of the gasifier. In this section, we assume that the flow rate of medium steam is fixed as a constant, at S/C = 12%. Adjusting the flow rate of oxygen enables the O/C ratio to change from 70% to 105%. The effect of the O/C ratio on the properties of raw syngas is depicted in Fig. 7. It should be noted, the trace gas, such as CH4 and, H2 S, is not shown in this figure. From Fig. 7, the raw gas concentration is sensitive to the O/C ratio. With the increase in O/C, more and more CO and H2 react with oxygen to make H2 O and CO2 ,

leading to an increase in CO2 and the opposite trend in CO and H2 O concentrations. Meanwhile, the combustion of CO and H2 will release more heat, which is beneficial for the generation of HP steam in the HRSG. Fig. 8 shows the variation of efficiencies and CO2 capture efficiency with O/C in the conventional Selexol-based process. According to Fig. 8, by increasing O/C from 70% to 105%, the energy efficiency increases from 35.20% to 37.25% and the exergy efficiency increases from 31.71% to 33.56%, while the CO2 capture efficiency slightly drops from 87.79% to 86.32%. As reasonably expected, with the increase in O2 , the flow rate of CO and H2 will decrease, and the conversion ratio of CO in the WGS reactor is assumed to be 95% according to Table 2, so the flow rate of H2 to the combustor will simultaneously decrease, eventually causing the power generated in the GT to also decline. In other cases, however, the increase in the output heat of the gasifier leads to an addition of HP steam flow rate and ultimately influences the power output of the steam turbine. The energy consumption of the ASU + O2 compression and AGR, CO2 drying and compression unit increases slightly, however, the total ancillary power consumption incremental ratio is less than the power generation rate, resulting in an increase in energy and exergy efficiencies.

CO CO2

80

Energy and exergy efficiency/%

Composition of the raw gas/(vol,%)

100

H2 CO+H2

60

40

20

0

70

75

80

85 90 95 O/C ratio/(wt,%)

100

105

Fig. 7 – The impact of O/C ratio on the composition of raw gas.

42

ηnet

41

ηex

40 39 38 37 36 35 34

0

5

10 15 S/C ratio /(wt,%)

Fig. 9 – Effect of S/C on the efficiencies of CLC-based technology.

20

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chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 110–124

Common processes Air separation unit Coal preparation Gasification process

Syngas quench cooling option

CO2 compression

Heat recovery steam generator

Physical absorptionbased process Water gas shift reaction Acid gas removal

Turbine of GT

CLC-based process Sulfur removal Fuel reactor Air reactor Air turbine

CO2 turbine

CLP-based technology

Oxygen purity: 95% (vol.) Power consumption: 225 kWh/ton O2 O2 delivery pressure: 35 bar Specific work: 0.022 kWh/kg coal Shell gasifier with gas quench configuration Gasification pressure: 30 bar Gasification temperature:1400 ◦ C Fixed carbon conversion: 99.9% Heat losses in the gasifier: 0.5% of input HHV O2 pressure to gasifier: 35 bar Steam pressure to gasifier: 32 bar Oxygen to coal ratio: 0.7–1.1 kg/kg Steam to coal ratio: 0–0.2 kg/kg Quenched gas temperature:925 ◦ C Cold recycled gas temperature:500 ◦ C Recycle gas compressor isentropic efficiency: 83% Recycle gas compressor mechanical efficiency: 99% CO2 storage pressure: 110 bar compressor isentropic efficiency: 83% Compressor mechanical efficiency: 99% Pinch point: 10 K HP Steam: 124 bar MP steam: 32 bar LP steam: 3 bar Condenser pressure: 0.046 bar Steam turbine isentropic efficiency: 87% Steam turbine mechanical efficiency: 99%

Pressure: 30 bar Temperature: 230 ◦ C Steam/CO molar ratio: 2:1 Two-stage Selexol configuration for CO2 and H2 S removal Syngas temperature at absorption tower inlet: 35 ◦ C CO2 removal yield: 90% H2 S removal yield: 99.9% Isentropic efficiency: 88% Mechanical efficiency: 99% Inlet temperature: 1416 ◦ C Discharge pressure: 1.03 bar

Selexol based physical absorption H2 S removal yield: 99.9% Pressure: 15 bar Heat duty: 0 kW Pressure: 15 bar Temperature: 1200 ◦ C Isentropic efficiency: 87% Mechanical efficiency: 99% TIT: 1200 ◦ C Discharge pressure: 1.03 bar Isentropic efficiency: 87% Mechanical efficiency: 99% Discharge pressure: 1.03 bar

Table 2 – (Continued) Syngas expander Carbonator

Outlet pressure: 2 bar Carbonation reactor temperature: 650 ◦ C Steam/CO molar ratio: 2 Calcium/CO molar ratio: 1.32 Heat loss: 0.5% Temperature: 950 ◦ C Reaction temperature: 1.5 bar Coal feedstock: meet all the CaCO3 decomposition O2 feed (weight basis): 2.28mcoal

Calciner

3.2.2.

Sensitivity analysis for CLC-based process

Fig. 9 indicates the variation of energy and exergy efficiency with the S/C ratio of the CLC-based process. The simulation finds that the carbon capture is greater than 99.9%, which means that the CO2 generated in this process is nearly completely captured, so the carbon capture efficiency is not presented in Figs. 9 and 10. As observed in Fig. 9, when the O/C ratio is fixed at 86.7%, increasing S/C from 0 to 20% slightly suppresses the energy efficiency from 39.94% to 39.49% and the exergy efficiency from 35.89% to 35.67%. The Ni-based looping combustion system leads to the absence of the AGR unit and results in CO2 generated in the fuel reactor that is captured and compressed for storage, and therefore, the power consumption for the AGR system (QAGR ) only includes desulfurization and will not, within the studied scope, change H2 S removal yield: 99.9%. Thus, the value of efficiencies for the CLC-based process are larger (approximately 3%) than the physical absorption-based technology. As explained above, as the S/C increases, the flow rate of syngas (CO and H2 ) increases, which leads to the output power of the GT exhibiting the same trend as the S/C. However, there is the flow rate of MP steam derived from the MP steam turbine, resulting in the decline of output power in the ST. Synthesizing the two factors between the GT and ST, the total output power declines. The power consumption of the CO2 drying and compression, ASU and oxygen compression units is almost constant. This analysis and results indicate a slight drop in the plant efficiencies. Fig. 10 illustrates the effect of O/C on the efficiencies in the case of a CLC-based process. From Fig. 11, it can be obtained that both energy and exergy efficiencies slightly decrease from 40.85% to 38.64% and from 36.20% to 34.82%, respectively.

44 Energy and exergy efficiency/%

Table 2 – The general process parameters and basic assumptions.

ηnet ηex

42 40 38 36 34 70

75

80

85 90 95 O/C ratio /(wt,%)

100

Fig. 10 – Effect of O/C on the efficiencies of CLC-based technology.

105

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Energy and exergy efficiency/%

40

ηnet ηex

38

36

34

32

30

0

5

10 15 S/C ratio /(wt,%)

20

Fig. 11 – Effect of S/C on the efficiencies of CLP-based technology. When O/C increases from 70% to 105% the flow rate of CO2 increases while the flow rate of CO and H2 drops remarkably. The heat released from the fuel reactor through the reaction of CO and H2 with NiO also decreases, which causes the power output in the CO2 turbine to decline. Consequently, because of the reduction of the reduced Ni, the heat released in air reactor also declines. In the HRSG system, the flow rate of HP steam increases, which is favorable for power generation. However, the ASU + O2 consumption, and AGR + CO2 drying and compression unit power consumption also exhibit the same tendency as the O/C, and its consumption rate is faster than the output power increase rate, and subsequently, there is a slow decrease in net and ex , according to Eqs. (1) and (3).

increase in S/C, more steam is split from the medium pressure steam turbine, leading to a decrease in plant power output, but its impact is limited. The flow rate of H2 at the exit of the gasifier increases, resulting in additional power output in the gas turbine. Furthermore, adding more steam to the gasifier means promoting the WGS reaction, increasing the concentration of CO2 (Fig. 8), which indicates that the consumption of CaO in the carbonator increases and more coal is required to meet the heating needs for decomposition of CaCO3 in the calciner, instead of generating more in the ASU system. As a whole, from the above analysis, the aggregate performance of Wnet increases slightly and the energy efficiency goes up minimally as inferred in Eq. (1). The exergy of coal is larger than the heat of coal input, and other exergy is destroyed in some reactions and separation units, and thus, the exergy efficiency is smaller than energy efficiency, but the exergy trend is consistent with the energy efficiency. Fig. 12 highlights the effect of the O/C on the efficiencies in the CLP-based process. With the increase in O/C from 70% to 105%, net increases from 35.48% to 38.97%, and ex rises from 32.95% to 35.39%. To control the temperature of the combustor, all nitrogen from the ASU is directly injected into the combustor (seeing Fig. 3). With the addition of oxygen, a large amount of nitrogen is also sent to the combustor, which results in an increase in power output in the GT. In the meantime, the continual increase in heat derived from the gasifier requires the addition of recycle water for its removal, and this will also promote output power. Nevertheless, the power consumption of ASU and O2 compression, AGR + CO2 drying and compression also increases at a slower rate than the energy output. In sum, the energy and exergy efficiencies are consistent with the increase in the O/C ratio.

3.3. 3.2.3.

The effect of S/C on the efficiencies of energy and exergy fot the CLP-based process is clearly shown in Fig. 11. Because of the CO2 capture efficiency is nearly 99.9% throughout simulation, and the carbon capture rate is not shown in Figs. 11 and 12, which are the same as the CLC-based process From Fig. 11, it is found that the S/C has a very limited impact on the efficiencies. The results indicate that increasing the S/C from 0 to 20% causes the energy and exergy efficiencies to rise from 37.38% to 37.75% and 33.71% to 33.93%, respectively. The steam to the gasifier is bled off from the outlet of the MP turbine. With the

Energy and exergy efficiency/%

40

ηnet ηex

38

36

34

32

30

Power balance of the studied system

Sensitivity analysis for CLP-based process

70

75

80

85 90 95 O/C ratio /(wt,%)

100

105

Fig. 12 – Effect of O/C on the efficiencies of CLP-based plant.

This section focuses on the energy and exergy balance of the three analyzed systems. To compare them to the IGCC without CCS system, the benchmark process is also modeled, simulated and listed in Table 3. Notably, the comparisons of these processes from the point of thermodynamics are built in the same feedstock conditions, S/C and O/C, etc. As illustrated in Table 3, the benchmark process (IGCC without CCS) exhibits the technical indicators of the current industrial plants using the Shell gasifier (the net power efficiency and exergy efficiency are 44.10% and 39.54%, respectively) (Cormos and Cormos, 2013). The simulation result of the CLCbased process gives a net power efficiency of 39.78%, which is in good agreement with the efficiencies in previous studies (Mukherjee et al., 2014; Rezvani et al., 2009). The other two carbon capture cases (physical absorption-based and CLP-based processes) generate approximately 240–350 MW net power, with net power efficiencies in the range of 36% to 38% and carbon capture rates of 86.97% and 99.9%, respectively. The CLC-based process realizes an exergy efficiency of 35.67%, in contrast to 33.86% for CLP-based and 32.74% for physical absorption-based technologies. As for the penalty expressed in the net power efficiency related to CO2 capture, for the physical-based process, the energy penalty is approximately 8% net electricity, which corresponds to the range of 6–10 percentage points seen in the study of Carbo et al. (2007). This value is higher than the CLC-based penalty of 4.32% and the CLP-based penalty of 6.38%. From this penalty point of view, the CLC-based and CLP-based processes are very promising, and further improvements can be conducted through detailed

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Table 3 – Overall plant performance indicators. Description

Units

Benchmark process

S/C O/C Coal feed Feedstock thermal energy, LHV Gas turbine output Steam turbine output Gross electric power output Gasification island and coal preparation ASU consumption + O2 compression AGR + CO2 drying and compression Power island power consumption Total ancillary power consumption Net power output Net power efficiency, net Exergy efficiency, ex CO2 capture efficiency, CO2

kg/kg kg/kg kg/h MWth MWel MWel MWel MWel MWel MWel MWel MWel MWel % % %

0.12 0.87 88,360 662.58 181.75 146.42 328.16 3.41 28.53 – 8.01 39.95 292.21 44.10 39.54 0

heat integration and different schemes for chemical looping characteristics. However, the commercial for large-scale technological applications of those chemical looping integrated processes still require time to investigate the different system parts, such as the design of the loop fluid-bed reactor and the gas turbine (Cormos and Cormos, 2013). From Table 3, one important characteristic has to be noticed for the in the CLP-based process In a comparison between the physical absorption-based and CLC-based processes, the power generated by steam turbine corresponds to its gas turbine generation, while the others show that the gas turbine output power is larger than its steam turbine output. In CLP-based process, especially in the calcium looping section, both the carbonation reaction (CaO + CO2 = CaCO3 ) and the water gas shift reaction (CO + H2 O = CO2 + H2 ) are exothermic, which is beneficial to the steam generation in HRSG, and leads to an increase in the energy output. The power consumption of AGR and CO2 compression is higher in physical absorption-based than in a CLC-based process due to external electricity required for the absorption of CO2 . The CLP-based process consumes significantly more coal and oxygen than the physical absorption-based and CLC-based plant. Notably, the additional fuel and O2 consumptions are used to meet the energy demand for CaCO3 -decomposition, substituting the HRSG to recover high-quality heat of CO2 from the calciner, which contributes significant power with regard to overall plant. Compared with the physical absorption-based

Physical absorption-based

CLC-based

0.12 0.87 88,360 662.58 181.42 132.31 313.73 3.41 28.53 24.21 14.75 70.91 242.82 36.64 32.74 86.97

0.12 0.87 88,360 662.58 165.61 151.34 317.95 3.41 28.53 19.12 3.31 54.37 263.58 39.78 35.67 99.9

CLP-based 0.12 0.87 122,544 918.91 219.21 220.57 439.78 4.16 45.09 25.71 18.15 93.11 346.67 37.72 33.86 99.9

and CLC-based technology, the CLP-based has a greater net energy output and does not require acid gas removal units. Under the above analysis, the CLC-based process is selected as the illustrative one to exhibit the key points of the plant. The pressure, temperature and composition of the process main stream nodes are presented in Table 4.

3.4.

Exergy balance of CLC and CLP-based processes

In this section, the detailed exergy analysis of CLC and CLPbased processes by means of exergy balance calculation are presented in the form of a Grassmann diagram, where the exergy flows with different streams in each unit are presented as arrows with width proportional to the size of exergy flow. For a steady-state process, the exergy destruction is caused by internal irreversibility which can be calculated by Eq. (3). The exergy loss is the exergy leaving the process contained in streams that cannot be used in further processes, such as the exhausted gas and the depleted air from HRSG (Anheden and Svedberg, 1998). In this paper, to express the flow direction and value of exergy conveniently, the exergy destruction and loss are assembly viewed and expressed as exergy loss. Figs. 13 and 14 show the Grassmann diagram of the CLC and CLP-based technologies, respectively. From Fig. 13, it can be observed that most exergy losses in CLC-based plant are caused by chemical looping combustion and the gasifier, which combined, account for 50% of the total exergy loss. On

Table 4 – The main streams parameters of the CLC-based process. Stream

Coal (gasifier)

Oxygen (gasifier)

Steam (gasifier)

Raw syngas (Ex.gasifier)

Pressure (bar) Temperature (◦ C) Mass flow (kg/h) Molar flow (kmol/h) Composition (mol%) O2 N2 H2 CO CO2 H2 S H2 O Other

1.01 25.00 88 344

35.00 133.40 76 696.03 2 411.33

30.40 600.00 10 601.22 588.45

30.10 1 400.00 166 407.43 8 111.03 0.00 2.00 28.89 58.10 3.70 0.90 6.40 0.01

95 5

100

Syngas (Ex.sulfur removal)

FR outlet gas

AR outlet gas

Captured CO2

28.45 36.40 154 355.26 7 522.49

15.00 1160.20 264 647.30 7 524.71

15.00 1200.00 363 431.49 12 973.25

111.46 40.00 214 110.76 4868.74

0.00 2.21 31.13 62.41 4.03 0.00 0.2 0.02

0.00 2.22 0.30 1.72 64.69 0.00 31.02 0.05

97 ppm 99.98 0.00 0.00 0.00 0.00 0.00 0.01

0.00 0.59 0.08 0.10 99.12 0.00 0.00 0.11

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Fig. 13 – The Grassmann diagram for CLC-based process. the one hand, the exergy destruction are mainly caused by the irreversible chemical conversion of the syngas in CLC system, on the other hand, the heat that is released from the air reactor is limited, leading to the limitation of temperature for the depleted air from the AR and directly affecting the power generation for the air turbine. In the gasifier, all of the solid fuels are transferred into crude synthesis gas, which has unavoidable exergy destruction. Significantly, the exergy loss in the CCS includes CO2 compression and cooling. For the CLP-based process, the most exergy losses rests with the combustion chamber, the calcium looping process, the CO2 capture and the gasifier section. For the combustion chamber, the irreversible chemical combustion reaction contributes to the exergy loss, and the combustion temperature determines the power generation in the GT. For the CLP unit, the exergy losses mainly lie in the combustion of coal which supplies energy for the calcination of CaCO3 .

4.

Economic feasibility

The thermodynamics analysis of these power plants is useful to realize the feasibility of process from technology, while knowing those costs is also significant to design and evaluate power plants. In this part, the economic analyses of the researched processes are conducted. It should be noted that the economic evaluation of this paper is just simply to reveal possible bottlenecks and compare different processes. Subsequently, the results of this evaluation may not reflect the final cost of these plants accurately. The economic indicators of each process are determined by computing the capital estimations, operational and maintenance costs, payback period, net present value (NPV), and internal rate of return (IRR). For the calculation of capital

estimation, each sub-system or equipment is estimated using the following equation: CE = CB × (Q/QB )

M

(7)

where CE represents equipment cost with capacity Q; CB represents the base cost with capacity QB , which can be derived from published cost data (Black, 2010; Cormos, 2012b, 2014; Olaleye and Wang, 2014); M represents the constant varying on equipment type, and M typically assumed in the range between 0.6 and 0.7 (Skorek-Osikowska et al., 2014). Mathematically, the internal rate of return (IRR) is defined as NPV = 0 =

N t=1

It /(1 + IRR)

t

(8)

where It is net income in t years; N is total plant life. Table 5 presents the basic parameters assumed for calculating economic indicators in all design cases (Cormos, 2012b, 2014; Olaleye and Wang, 2014; Salkuyeh, 2015). Based on these parameters and assumptions, the total investment cost of each sub system and equipment for the researched systems are summarized and listed in Table 6. From Table 6, the total investment costs are the highest in the CLP-based process, followed by the physical absorption-based, CLC-based and benchmark IGCC without CCS processes. With regard to the total investment cost per kWe of those cases, the benchmark process is lowest (1872.37 D /kW) when compared with the residual processes (2547.5D /kW for physical absorptionbased, 2245.62 D /kW for CLC-based and 2425 D /kW for CLP-based), and the main reasons for the differences among

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Expander

604.9

596.56

H2-rich gas

O2 8.51

Air in 7.85 71.36

N2 20.25

GT

Air compressor loss 23.79 MP Steam 40.23 212.11

GT loss 18.79

211.97 Air in 4.09

CLP loss

55.42

AC

105.76

495.44

126.69

102.34 138.89 Combustion loss 135.16

ST

Power in 46.45

Coal 280.1 CLP

ASU

318.61

Power 8.34

Combustion chamber

Flue gas 36.59 HRSG

Loss

622.61

To HRSG 32.36 HRSG loss 71.93

CCS 116.53

O2 9.55

649.57

Loss 1.81

Scrubber

737.12

Loss 13.71 Gas quench cooling

Loss 87.53 Gasification

Coal 739.04

Preparation

Loss 1.94

Flue gas Power

Condenser loss 17.81 ST loss 16.53

174.59

Other:4.35 Unit:MW

Fig. 14 – The Grassmann diagram for CLP-based process.

Table 5 – The basic economic assumptions. Coal price Limestone Nickel oxide Electricity BFW and process water price Cooling water price Selexol solvent price BFW and process water treatment Slag disposal cost Direct labor Average annual direct labor cost Administrative, support and overhead cost Annual maintenance costs Working capital Discount rate Carbon tax Construction time Plant life Depreciation residual ratio Taxes Capacity factor Each process utilities and offsite units costs Owner’s cost and contingency cost Land, permitting, surveying. etc cost

2.2D /GJ 20D /t 14,000D /t 0.1D /kWh 0.10D /t 0.001D /t 6500D /t 90,000D /month 10.0D /t 90 persons 50,000D /person 30% from direct labor cost 3.5% from capital expenditure 30 days’ supply 8% None 3 years 25 years 6% 20% 85% 25% of the main plant subsystems cost 15% of total installed cost 15% of total installed cost

these values lie in the cost penalty for carbon capture and its related CO2 drying. The calculation of operational and maintenance (O&M) cost can be divided into two parts: one is fixed O&M costs, the other is variable O&M costs. All of the costs are estimated according to published papers and books (Cormos, 2014; Peters et al., 2003; Rezvani et al., 2009). After that, the main economic indicators of the overall plants are presented in Table 7. As shown in Table 7, the dynamic payback year represents the periods that the net cumulative return earning recovers the investment, and the smaller the payback period, the more promising the project is. With the introduction of different carbon capture units, the dynamic payback period shows an increase approximately 7.15 years for physical absorptionbased process, 3.35 years for CLC-based and 3.11 years for CLP-based processes. In addition, the internal rate of return measures the exponential nature of growth in the capability to do electrical work (Warr and Ayres, 2012). The IRR for the physical absorption-based case is 10.5%, followed by 12.5% for the CLC-based and 12.8% for the CLP-based. The process simulation and economic estimation in this paper indicate that the CLC-based and CLP-based processes seem to be promising carbon capture options integrated with an IGCC plant, with significantly higher thermodynamic performances and lower energy and cost penalties for CO2 capture. Both the CLC-based and CLP-based carbon capture units are conducted on a circulated fluidized bed (CFB) reactor platform. However, large scale commercial CFB technology has been difficult to develop. Among the technical challenges

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Table 6 – estimation of process sub-systems cost. Plant sub-systems

Units

Solids handing facilities ASU Gasification island Syngas processing unit Acid gas removal unit Chemical looping combustion Calcium looping unit CO2 processing and drying Sulphur removal unit power island Utilities and offsite units Total install costs Owner’s cost and contingency Land, permitting, surveying, Etc. Total investment costs Net power output Total investment cost per kWe

MD MD MD MD MD MD MD MD MD MD MD MD MD MD MD MWe D /kW

Benchmark process

Physical absorption-based

CLC based

CLP based

18.79 71.78 105.24 27.32 0.00 0.00 0.00 0.00 11.44 130.18 91.19 455.94 68.39 22.80 547.13 292.21 1872.37

18.79 71.78 105.24 27.32 40.86 0.00 0.00 13.17 12.31 122.92 103.10 515.49 77.32 25.77 618.59 242.82 2547.50

18.79 71.78 105.24 27.32 0.00 15.86 0.00 15.02 12.32 128.27 98.65 493.25 73.99 24.66 591.90 263.58 2245.62

29.90 114.47 109.09 29.34 0.00 0.00 66.72 16.53 12.32 182.08 140.11 700.56 105.08 35.03 840.68 346.67 2425.00

CLC-based

CLP-based

Table 7 – the main economic indicators for the overall plants. Parameters

Units

Total capital investment Total fixed and variable O&M costs Dynamic payback period Net present value Internal rate of return

MD MD /y years MD %

Benchmark process

Physical absorption-based

547.12 61.71 10.1 502.05 15.7

618.58 64.86 17.25 162.21 10.5

related to CFB are the reactor cooling, looping seal design, and high steady recycle sorbent or oxygen carriers particles, etc (Linderholm et al., 2014; Ma et al., 2015; Romano et al., 2013).

5.

Conclusion

In this work, the technical aspects of IGCC power plant with CCS in three distinct configurations was simulated and compared from both the thermodynamic and economic view. They are IGCC with pre-combustion CO2 capture by physical absorption, IGCC with CLC process, and IGCC with CLP process. The effects of S/C and O/C on the performances of these three processes with respect to energy and exergy efficiencies were investigated. Grassmann diagram was adopted to show the detailed exergy analysis of CLC and CLP-based technologies. Sensitivity analysis demonstrated that with the increase of S/C from 0 to 20%, the energy and exergy efficiencies of the physical absorption-based system decreased and these were differed from the CLP-based process which exhibited opposite tendency, while the energy and exergy efficiencies of CLCbased technologies dropped slowly. As the O/C ratio jumped from 70% to 105%, the energy and exergy efficiencies of the physical absorption-based and CLP-based processes showed the growth trend. As for CLC-based process, its energy and exergy efficiencies declined with the reduction of O/C. Under the same feed conditions, the physical absorption and CLP-based technologies displayed a net power efficiency of 36.21% and 37.72%, respectively, which is lower than CLCbased technology of 39.78%. The CLC technology realized an exergy efficiency of 35.67% in contrast to 33.86% for CLP and 32.74% for physical absorption-based technologies. The overall CO2 capture efficiencies of the CLC and CLP-based processes were 99.9% while the physical absorption-based process was 86.97%. Exergy analysis of CLC and CLP with a Grassmann diagram demonstrated that the most exergylossing parts of the CLC-based process were in chemical

591.9 65.85 13.45 290.75 12.5

840.67 86.62 13.21 389.05 12.8

looping combustion, followed by the gasifier. The sum of exergy losses for these two parts accounted for 50% of the total exergy loss. For the CLP-based process, the most exergy losses were in the combustion chamber, the calcium looping process, the CO2 capture, and the gasifier section, in that order. The economic evaluation performed showed that the CLC and CLP-based processes have a payback period of 13.45 and 13.21 years, which is lower than the physical absorption-based technology with 17.25 years. The IRR evaluations suggest that the physical absorption-based case was 10.5%, followed by 12.5% for CLC-based and 12.8% for CLP-based. In addition to the evaluation of those three systems conducted in current work, other potential improvements will be investigated in the future, such as using two-stage CLC system, replacing ASU with membrane separation and applying pressure swing adsorption in CO2 removal units. Furthermore, the scale-up issues towards commercial size of CLC and CLPbased circulated fluidized bed technology need to be tenderly researched.

Acknowledgments This work was financially supported by Graduate Innovation Foundation of Southwest Petroleum University (no. CXJJ2015011). The authors also would like to thank the Center of Oil and Gas Processing at the Southwest Petroleum University for providing the support of Aspen Plus.

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