Reducing the Cost of CO2 Capture from Flue Gases Using Aqueous

Article pubs.acs.org/IECR

Reducing the Cost of CO2 Capture from Flue Gases Using Aqueous Chemical Absorption Anggit Raksajati,†,‡ Minh T. Ho,†,‡ and Dianne E. Wiley*,†,‡ †

School of Chemical Engineering, The University of New South Wales, Sydney 2052, Australia The Australian Cooperative Research Centre for Greenhouse Gas Technologies (CO2CRC), Barton, Australian Capital Territory, 260, Australia

‡

S Supporting Information *

ABSTRACT: Chemical absorption is widely regarded as the most promising technology for CO2 capture from large industrial sources in the short term. The cost of CO2 capture from postcombustion power plants using monoethanolamine (MEA), the benchmark for chemical absorption, is currently over US$70 per metric ton of CO2 avoided. This high cost is considered as the major obstacle to current large-scale implementation of carbon capture and storage (CCS). Thus, there has been significant focus on the development of new solvents with the aim to reduce costs. This paper provides insights into the impact of solvent properties on the cost of capture to assist in the development of new solvents based on a 500 MW supercritical black coal power plant as the emission source. The effect of solvent properties, specifically solvent loading, heat of reaction, solvent loss, and solvent concentration is examined. The effect of improvements in process design, specifically high pressure stripper operation, advanced structured packing, use of concrete for the process vessels, and advanced heat exchangers, is also evaluated. Sensitivity analysis and Monte Carlo simulation are performed to provide an estimate of the capture cost variability. The results show that the development of aqueous chemical absorption technology for CO2 capture should focus on new solvents with good stability toward SOx and NOx, high solvent concentration (above 50 wt %), and high working capacity (above 0.35 mol of CO2/mol of solvent). These three parameters have the most significant impact on the capture cost. Based on Monte Carlo simulation, within a 95% confidence level, the capture cost with improved solvent properties and process design is estimated at US$62−80 per metric ton of CO2, with the most likely cost of US$71 per metric ton of CO2 avoided. This number reduces to US$44−59 per metric ton of CO2, with the most likely cost of US$52 per metric ton of CO2 avoided, if the flue gas desulfurization (FGD) and selective catalytic reduction (SCR) units can be eliminated.

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INTRODUCTION

cost of capture can be minimized independently from the rest of CCS chain.4 Various technologies are available to capture CO2 from flue gases; however, chemical absorption is regarded as the most promising technology in the short term. Monoethanolamine (MEA) is the most widely studied solvent and is considered to be a “first generation solvent”. Previous research related to capture of CO2 using MEA has covered the thermodynamic and kinetic model,5,6 absorption model,7 degradation study,8,9 inhibitor assessment,10 and solvent regeneration optimization.11−16 A recent study by Ho et al.17 shows that the cost of CO2 capture (in 2008 terms) from a 500 MW subcritical lignite power plant is over US$70 per metric ton of CO2 using commercially available unpromoted 30 wt % MEA solvent. When advanced heat integration with the power plant is applied, the capture cost decreases to US$55 per metric ton of CO2 avoided. However, this cost needs to decrease even further in order to facilitate large-scale implementation of CCS. As a comparison, in Australia the carbon tax in 2012 was A$23 per metric ton of CO218 and the European Union Emissions

There is a growing interest to address climate change, as currently it is considered one of the greatest environmental, social, and economic challenges. Empirical evidence of climate change has increased over the past century as shown by rises in air and sea temperatures, increasing sea level, melting glaciers, reduction of snow cover, and trends in extreme weather patterns.1 It is widely believed that such changes have occurred due to an increase in greenhouse gas concentration in the atmosphere. The concentration of carbon dioxide, the most abundant greenhouse gas, has risen from 280 to 400 ppm since the industrial revolution.2 Carbon dioxide is generated and emitted by human activities, most notably from the combustion of fossil fuels, such as oil, coal, and gas. The reduction of carbon dioxide emissions can be achieved through multiple options such as increased use of renewable energy and improved energy efficiency. However, the complete decarbonization of the power sector and industries is currently too costly. Carbon capture and storage (CCS) is considered to be one of the key technologies that will enable large-scale fossil fuel powered power stations and industries to continue to be economically viable while reducing their CO2 emissions in a carbon-constrained world.3 The cost of CCS is dominated by the cost of capture, and it has been reported that, in order to minimize the avoidance cost of a CCS project, the avoidance © 2013 American Chemical Society

Received: Revised: Accepted: Published: 16887

July 9, 2013 October 25, 2013 October 28, 2013 October 28, 2013 dx.doi.org/10.1021/ie402185h | Ind. Eng. Chem. Res. 2013, 52, 16887−16901

Industrial & Engineering Chemistry Research

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Figure 1. Simplified process flow diagram for the chemical absorption unit CO2 capture process: (1) pretreatment (FGD, SCR, DCC, and blower), (2) absorber, (3) solvent pump, (4) cross heat exchanger, (5) stripper, (6) reboiler, (7) lean cooler, (8) condenser, (9) posttreatment (dehydration and CO2 transport compressor).

Trading System (ETS) carbon price was €3.5 per metric ton of CO2, which means that the current cost of implementing CCS is higher than both the Australian carbon tax and the EU carbon credit price. One of the reasons for the high capture cost using MEA solvent absorption is due to the large amount of energy needed to operate the capture plant. This energy requirement is dominated by the energy required to regenerate the solvent in the stripper column. The high energy requirement is considered to be the main drawback of using MEA-based solvent capture. The other major limitations are high solvent loss and corrosion at high MEA concentration. Various groups have developed alternative solvents to MEA in an effort to improve the CO2 separation performance and reduce the cost of capture. For example, KS-1, a proprietary hindered amine based solvent, has been developed by Mitsubishi Heavy Industries (MHI). KS-1 has a larger solvent loading capacity (the differential between rich and lean loading) than MEA by up to 40%. Thus, the circulation rate of KS-1 can be reduced by up to 40% compared to the MEA-based process, resulting in a reduction in the regeneration energy of up to 32%.19 KS-1 exhibits low corrosion, which enables the use of carbon steel material for most of the capture plant, which reduces the cost of the facility. It also has a much better stability than MEA: solvent loss is only 10% of that for MEA.20 The KS1 solvent can be considered as one of the “second generation” advanced amine-based solvents. KS-1 is currently at the demonstration stage and is likely to be commercially implemented within 5 years.20 Development of a potassium carbonate based solvent has attracted increasing interest in recent times. This solvent has several advantages compared to MEA, such as lower solvent cost, less corrosion, less solvent degradation, and less solvent slip. This solvent also has a lower heat of reaction, which potentially reduces the heat of regeneration for the low pressure stripper operation.21 However, the absorption of CO2 into a potassium carbonate solution is limited by slow absorption rates.22 Several researchers have shown that the use of promoters can increase the rate of CO2 absorption. One of the promoters that has been extensively studied is piperazine, which can improve the absorption rate to be comparable with that of MEA.23,24 Other studies have shown that injection of boric acid, which is more environmentally benign than

piperazine, doubled the overall rate of CO2 absorption.22,25,26 This solvent also potentially offers an advantage of involving a phase-change step (precipitation) during the CO2 absorption.27 Based on Le Chatelier’s principle, the precipitation process during the CO2 absorption will shift the reaction toward a higher solvent capacity, which potentially reduces the regeneration energy. The other examples of phase-change solvents, which can be considered as “third generation” solvents, are amino acids28−30 and chilled ammonia.31 However, this solvent group is still in the early stage of research and development. Thus, this solvent generation is not included in this paper. Only a few studies have evaluated the potential cost reduction due to solvent development. Unlike most papers that focus on reducing the energy consumptions of solvent systems, this paper focuses on examining the opportunities to reduce the capture cost. This is important because the cost determines the viability of CO2 capture projects. The cost can be reduced in a number of ways including reducing the energy consumption, reducing the capture plant capital cost by using a solvent with a faster absorption rate, using a contactor with a better mass transfer efficiency, changing operating conditions, and/or using a cheaper material for the process vessels. In this paper, the effect of the solvent properties and the process design on the solvent absorption performance (solvent flow rate, regeneration energy, and equipment size) is examined. The subsequent impacts of variations in these parameters on the capital costs, operating costs, and the cost of capture are quantified. The sensitivity of capture cost to the key parameters is also analyzed. By identifying and ranking the solvent properties and their impact on the capture cost, the extent of the effects of improvements in solvent working capacity, solvent concentration, heat of reaction, mass transfer rate, and NOx and SOx stability on the capture cost reduction are examined.

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METHODOLOGY Flue Gas. This paper uses flue gas from a newly built 500 MW supercritical black coal power plant in Australia as the emission source. The flue gas is assumed to contain 13 mol % CO2, 75 mol % N2, 5 mol % O2, 7 mol % H2O, 400 ppm SOx, and 450 ppm NOx. The CO2 emission intensity of the power 16888

dx.doi.org/10.1021/ie402185h | Ind. Eng. Chem. Res. 2013, 52, 16887−16901


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plant is 0.808 metric ton/MWh electrical. Due to the higher quality of the coals and less stringent regulation of SOx and NOx emission, most of the existing power plants in Australia do not have flue gas desulfurization (FGD) and selective catalytic reduction (SCR) units. Therefore, these units are included as a part of the capture plant pretreatment process in this paper. This approach is consistent with our previous studies.32,33 For regions where FGD and SCR are already part of the power plant, most of the insights of this paper pertaining to the cost reductions opportunities through solvent development are still applicable. CO2 Capture. A simplified schematic of the CO2 capture process using chemical absorption is shown in Figure 1. This is based on commercially available technology using unpromoted MEA solvent. The pretreatment unit involves a flue gas desulfurization (FGD) unit to reduce the SOx content in the flue gas to 10 ppm, a selective catalytic reduction (SCR) unit to remove the NOx content in the flue gas, a direct contact cooling tower (DCC) if the flue gas temperature is higher than the absorber temperature of 40 °C, and a blower to reach the absorber pressure (slightly above atmospheric) and counter the pressure drop across the piping and the absorber column. With increasing absorption temperature, the solubility of CO2 in the aqueous solvent decreases but the reaction rate of dissolved CO2 with the amine solvent increases. Hence to strike a balance between solubility and reaction kinetics, the flue gas temperature at the absorber inlet is usually kept at around 40−45 °C in the case of 30 wt % unpromoted MEA solvent. The low absorber temperature also has the benefit of preventing excessive solvent loss due to vaporization in the absorber.34 Flue gas is fed to the bottom section of the absorber, while the lean solvent is fed to the top section. The flue gas flows upward through the packing, which increases the surface area for gas liquid contact. The reaction between CO2 and MEA results in CO2-rich solvent that exits the column from the bottom. Meanwhile, the lean flue gas, with much lower CO2 concentration than the feed gas, is emitted from the top. The pressure of the rich solvent is increased from atmospheric to above 2.5 bar to reach the optimal stripper pressure (typically 2 bar for unpromoted MEA), by taking into account the pressure drop across the piping and rich/lean heat exchanger as well as the geodetic pressure increase at the stripper. It is desirable to operate the stripper at a high pressure to reduce the overall equivalent energy requirement of the capture system by lowering the solvent regeneration energy and reducing the energy required to compress CO2 for transport. However, there is an upper limit to the pressure (e.g., 2 bar for unpromoted MEA) beyond which solvent degradation becomes excessive.21 After leaving the absorber, the rich solvent stream passes to the rich/lean heat exchanger, where the rich solvent temperature is increased by using heat from the hot lean solvent that comes from the stripper. The heated rich solvent goes to the stripper column, where the CO2 is removed by heating and the solvent is regenerated. Heat supplied by the reboiler is used to increase the temperature at the bottom of the stripper column to approximately 120 °C and to reverse the reaction between CO2 and MEA. The stripped CO2 exits the top of the stripper column, while the lean solvent goes to the rich/lean exchanger from the bottom of stripper. There the lean solvent stream is cooled to 40 °C before returning to the absorber to complete the closed-loop process. The typical processing conditions for the MEA absorption process are summarized in Table 1.

Table 1. Processing Conditions and Solvent Properties Used for the Unpromoted 30 wt % MEA-Based Chemical Absorption Process parameter

value67−69

absorption press. (bar) absorption temp (K) stripper press. (bar) stripper temp (K) solvent lean loading (mol/mol) rich loading (mol/mol) heat of reaction (MJ/kmol) solvent heat of vaporization (kJ/kg) solvent loss (kg/metric ton of CO2 captured) density (kg/m3) sp heat at 40 °C (KJ/kmol·K) CO2 solubility profile

1.2 313 2 393 30 wt % MEA 0.22967 0.45567 8268 82668 1.668 1020 84.4 Hilliard et al.69

The analysis in this paper assumes that 90% of the CO2 contained in the flue gas is recovered by the capture plant. In the posttreatment process, the separated enriched CO2 stream is dehydrated and then compressed to 100 bar ready for transport. The solvent used for the baseline case in this paper is unpromoted 30 wt % MEA. The results for KS-1 and Econamine FG are also presented in the section Baseline Economic Results. Furthermore, hypothetical generic solvents, with better solvent properties compared to the baseline, are evaluated in the case studies. Technical Calculations. The calculations are performed using an in-house techno-economic model developed by the University of New South Wales for the CO2CRC. Apart from straightforward mass and energy balance equations, the following absorption specific process and design equations are used. The absorber is designed based on a shortcut method as described by Strigle35 and Kohl,36 while the stripper is designed based on a shortcut method as described by Strigle35 and Ludwig.37 The column dimensions are calculated using flue gas characteristics, solvent properties, operating conditions, and packing characteristics. Packing characteristics are obtained from Kister.38 The heat exchanger area for the rich/lean heat exchanger, amine cooler, and reboiler are calculated using a shortcut method as described in Section 21 of the Gas Process Suppliers Association (GPSA) handbook.39 The mass transfer rate for the baseline case, using unpromoted MEA solvent, is calculated using the following equation:36 K Ga = F(L /μ)2/3 [1 + 5.7(Ce − Cs)M exp(0.12T − 3.35 × 10−5p − 3.08)]

(1)

where KGa is the overall mass transfer coefficient in lb·mol·h−1· ft−3·atm−1, F is the packing correction factor, L is the liquid flow rate (lb·h−1·ft−2), μ is the solvent viscosity (cP), Ce is the equilibrium CO2 loading in solution (mol of CO2/mol of solvent), Cs is the CO2 loading in solution (mol of CO2/mol of solvent), M is the molar concentration of solvent (mol/L), T is the absorption temperature (K), and p is the partial pressure of CO2 flowing into the absorber (Pa) The regeneration energy is calculated using a mass and energy balance for the stripper. The regeneration energy consists of the sensible heat required to increase the solvent temperature to the stripper temperature (as shown in eq 2), the 16889



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Ttop is the stripper top temperature (K), and Tref is the reference temperature (K). The equivalent electrical energy requirement is the sum of the regeneration energy and power for compression, pumping, FGD, SCR, and flue gas blower, which is calculated using the following equation:

heat required to reverse the CO2 reaction with the solvent (eq 3), and the heat of vaporization of water (eq 4). Q sensible =

∫T

Treg

(msolvent + mCO2)cp dT

(2)

st,in

⎛ ° +R Q reaction = nCO2⎜ΔHabs ⎝ H vap

∫H

Q vapor = n H2O,v

∫T

Treg

0

⎞ dT ⎟ ⎠ R

cp

Weq = Wreboiler + Wpump + Wblower + Wcomp + Wfgd + Wscr

(3)

(8)

where Weq is the equivalent electrical energy requirement (MWe) and Wreboiler, Wpump, Wblower, Wcomp, Wfgd, and Wscr are the electrical power requirements for the reboiler, pumps, blower, compressors, FGD, and SCR, respectively (MWe). The conversion of thermal energy to electrical power for the reboiler is calculated using the equations from the study by Oexmann:40

dH (4)

liq

Using Dalton’s law, the term nH2O,v can be expressed as pH O n H2O,v = nCO2 2 pCO 2

(5)

⎛ Tcond ⎞ ⎟ Wreboiler = ηeff Q reg ⎜⎜1 − Treg + 10 ⎟⎠ ⎝

where Qsensible is the sensible heat (MW), Treg is the solvent regeneration temperature (K), Tst, in is the stripper inlet temperature (K), msolvent is the solvent mass flow rate (kg/s), mCO2 is the captured CO2 mass flow rate (kg/s), cp is the overall specific heat capacity (MJ/kg·K), Qsensible is the heat of reaction (MW), nCO2 is the captured CO2 molar flow rate (mol/s), ΔHabs ° is the heat of reaction at standard condition (MJ/mol of CO2 captured), R is the universal gas constant (MJ/mol·K), T0 is the temperature of standard condition (K), Qvapor is the heat of water vaporization (MW), nH2O,v is the vaporized water molar flow rate at the stripper top (mol/s), Hliq is the enthalpy of water saturated liquid at the stripper top temperature (MJ/ mol), Hvap is the enthalpy of water saturated vapor at the stripper top temperature (MJ/mol), and pH2O and pCO2 are the partial pressures of water vapor and CO2 in the gas phase at the stripper top outlet (Pa), respectively. Neglecting the flashing of the solution at the stripper inlet, the total regeneration energy (Qreg) can be expressed as sensible heat

←⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯→

ηeff = 0.78555 + 1.485e−7pext

heat of reaction

←⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯→

heat of vaporization

←⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯⎯→

pH O 2

pCO

nCO2

2

(6)

where Qreg is the total regeneration energy (MW), ΔHabs is the heat of absorption of solvent with CO2 at the stripper temperature (MJ/mol of CO2 captured), and ΔHvap,H2O is the heat of vaporization of water at the stripper top temperature (MJ/mol). The partial pressure ratio of water to CO2 in the gas phase at the stripper top outlet is estimated using eq 7 (based on Oexmann et al.21), which is derived from the Gibbs−Helmholtz equation and the Clausius−Clapeyron equation for water pH O 2

pCO

2

=

pH O,ref 2

pCO ,ref 2

⎛⎡ Ttop − Tref ⎤ exp⎜⎜⎢ ⎥[|ΔH vap,H2O| ⎝⎣ RTTref ⎦

⎞ − |ΔHabs|]⎟⎟ ⎠

(10)

where ηeff is the effective efficiency for conversion of thermal energy to electrical energy, Tcond is the condenser outlet temperature (K), and pext is the pressure at which the steam is extracted from the power plant to meet the reboiler energy (Pa). The electrical power required for compression, for pumping, and for the blower is calculated using mass and energy balances, assuming an equipment efficiency of 85%. The energy requirement for the SCR unit consists of the electricity for the ammonia injection system and the steam required for ammonia vaporization and injection. The energy required by the FGD unit is used to pump the slurry into the FGD unit. The energy requirements for the FGD and SCR units are based on Berkenpas et al.41 In the baseline case for this paper, the total energy required (electrical and lost generation due to steam usage) for the capture plant is nearly 28.9%, as shown in Table S1 in the Supporting Information. Economic Asssumptions. Capital and operating costs are estimated for precapture treatment (FGD, SCR, and blower), CO2 capture, and CO2 compression (postcapture). The total capital cost includes all process equipment shown in Figure 1, plus a general facilities cost. The operating cost includes fixed general maintenance costs comprising labor, nonincome government taxes that may be payable, and general insurance cost. The variable operating costs include costs for flue gas desulfurization, cooling water, and solvent makeup. Further details about the methodology employed in the economics are provided elsewhere.42 The economic feasibility of a CO2 capture project is measured in terms of the specific cost of CO2 avoided. This is the cost of establishing a CO2 capture plant for an industrial plant or power plant, and is calculated using the following equation:

Q reg = (msolvent + mCO2)cp(Treg − Tst,in) + ΔHabsnCO2 + ΔH vap,H2O

(9)

(7)

C=

where pH2O,ref and pCO2,ref are the partial pressures of water vapor and CO2 in the gas phase at the reference temperature (Pa),

total costcap − total cost ref CO2,avoided

(11)

The terms in eq 11 are given by 16890



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Figure 2. Capital cost, energy requirement, and operating cost breakdown for the chemical absorption process using unpromoted 30 wt % MEA. *The reboiler energy requirement represents the total equivalent electrical loss of the power plant (including the lost generation capacity from the steam requirement). COEref =

Total Costcap = (Capex PP,cap + OpexiPP,cap)

(Capex PP,cap + Opex PP,cap) + (CapexCP,cap + OpexCP,cap + Decom)

+ (CapexCP,cap + OpexCP,cap + Decom)

Ecaphcap

(12)

Total Cost ref = Capex PP,ref + Opex PP,ref

(17) COEref =

(13)

(Capex PP,cap + Opex PP,cap) + (CapexCP,cap + OpexCP,cap + Decom)

CO2,avoided = CO2,eref − CO2,ecap

Ecaphcap

(14)

(18)

where C is the cost of CO2 avoided ($ per metric ton of CO2), Total Costcap and Total Costref are the present value (PV) of the sum of all expenditures for the case with CO2 capture and for the case without CO2 capture respectively ($), CO2,avoided is the PV of the total CO2 avoided (metric ton), CapexPP,cap and CapexCP,cap are the PVs of the power plant capital expenditure and capture plant capital expenditure for the case with CO2 capture ($), CapexPP,ref is the PV of the power plant capital expenditure for the case without CO2 capture ($), OpexPP,cap and OpexCP,cap are the PVs of total power plant operating cost and total capture plant operating cost for the case with CO2 capture respectively ($), OpexPP,ref is the PV of power plant annual operating cost for the case without CO2 capture ($), Decom is the PV of decommissioning cost ($), and CO2,eref and CO2,ecap are the PVs of CO2 emitted for the reference plant without CO2 capture and for the plant with CO2 capture (metric ton). For power plants, the cost of CO2 avoided can also be calculated using following equation: C=

where COEcap and COEref are the costs of electricity generation for the case with and without CO2 capture respectively ($/MWh), EIref and EI2,cap are the rates of CO2 emissions for the reference power plant and the power plant with CO2 capture (metric ton/MWh), Eref and Ecap are the net power plant outputs for the reference power plant and the power plant with CO2 capture (MW), and href and hcap are the total operating hours for the reference power plant and for the power plant with CO2 capture (h). The cost year of the analysis is 2011, with cost reported in U.S. dollars (US$). The costs are evaluated on a pretax basis using a real discount rate of 7% assuming the project life is 25 years. The load factor of the power plant and the capture plant is 85%. The cost for coal is taken as US$1.5/GJ resulting in a baseline cost of electricity of US$58.7/MWh. The solvent cost for unpromoted −MEA is assumed to be US$1.5/kg.

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RESULTS AND DISCUSSION Baseline Economic Results. The capture cost for the baseline process using chemical absorption with MEA solvent is US$88 per metric ton of CO2 avoided. The capital investment for the capture plant is approximately US$1,415/kW. This high cost is related to the high energy requirement to operate the capture plant, which is approximately 203.5 MW for the base case (equal to 28.9%). The baseline value for the thermal regeneration energy requirement in this paper is consistent with the numbers presented in the literature for current commercial

COEcap − COEref EI ref − EIcap

(15)

where COEref =

Capex PP,ref + Opex PP,ref Eref href

(16) 16891



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Figure 3. Capture costs using various commercial solvents.

et al. In the studies by Oexmann et al. and Abu-Zahra et al., the FGD and SCR are considered to be part of the base power plant whereas these processes are part of the capture pretreatment in this paper. Moreover, there is also an additional cost for dehydration (as a part of posttreatment) in this paper that was not accounted for in the other two studies. This results in a 24% lower capital cost reportedly by Oexmann et al. compared to this paper. In addition, Abu-Zahra et al. did not include the CO2 transport compressor, resulting in a 49% lower capital cost than this paper. The capture plant operating cost for this paper (US$140/ kW/year) is 10% higher than the operating cost of Oexmann et al. (US$126/kW/year), while the operating cost of Abu-Zahra et al. (US$98/kW/year) is 24% lower than this paper because of the difference in the capture plant process boundary. As shown in Table S1 in the Supporting Information, the energy requirement for this paper of about 28.9% is within the range of the results reported by Abu-Zahra et al.50 and Oexmann et al.51 (30.6 and 23.4%, respectively). The discrepancy is due to the difference in the power plant CO2 emission intensity, the difference in the CO2 concentration in the flue gas, and the difference in the thermal regeneration energy requirement. The baseline value for the thermal regeneration energy requirement in this paper is in the range of other studies for unpromoted MEA of approximately 4.2− 4.4 MJ/kg CO2.43−46 This value is approximately 12% higher than the value that is reported by Abu-Zahra et al. and is approximately 38% higher than the value reported by Oexmann et al. The difference is most likely because this paper evaluates unpromoted MEA while Abu-Zahra et al. and Oexmann et al. consider a state-of-the-art MEA. This is supported by the fact that both Abu-Zahra et al. and Oexmann et al. used a higher solvent rich loading (0.484 mol of CO2/mol of MEA and 0.55 mol of CO2/mol of MEA, respectively). Despite the lower thermal regeneration energy, the total energy requirement in Abu-Zahra et al. is higher than in this paper. This is because of the difference in the formula used to convert thermal energy to electrical energy. The capture costs from Oexmann et al. and Abu-Zahra et al. are approximately 20% lower than that in this paper because of the differences in capital cost, operating cost, and energy requirement. Using the same economic assumptions and adjusting for differences in the costing year, their costs equate to US$68 and US$67 per metric ton of CO2 avoided, respectively. These numbers are close to the capture cost of this paper if the FGD and SCR are excluded from the

unpromoted 30 wt % MEA at approximately 4.2−4.4 MJ/kg CO2.43−46 Detailed results for the baseline case are presented in Table S1 in the Supporting Information, while the breakdown of the capital cost (Capex), operating cost (Opex), and energy requirement for this case is shown in Figure 2. The capture cost using this solvent decreases to US$81 per metric ton of CO2 avoided if heat integration is employed where the energy used is sourced from the crossover pipe between the intermediate pressure (IP) and low pressure (LP) turbine cylinders (with extraction pressure set as close as possible to the required steam temperature); in this way the heat extraction has a lower impact on the net electricity production from the power plant. More details of the heat integration opportunities of the CO2 capture plant with the base power plant are presented elsewhere.47−49 The capture costs for some solvents in addition to unpromoted MEA are shown in Figure 3. The capture cost using an advanced amine such as Econamine FG, which is a MEA-based solvent formulation developed by Fluor, is US$82 per metric ton of CO2 avoided, while the cost for Econamine FG Plus, which includes an optimized regeneration design with heat integration, is approximately US$79 per metric ton of CO2 avoided. The capture cost using a hindered advanced amine solvent such as the KS-1 is US$74 per metric ton of CO2 avoided. This value would be reduced to US$70 per metric ton of CO2 avoided when heat integration is employed. According to MHI, these two solvents have a “much higher” solvent capacity than MEA, thus lowering the required solvent circulation rate in the system and reducing the cost.43 Comparison with Published Studies. Techno-economic analyses of CO2 capture from black coal power plants using MEA-based chemical absorption have also been carried out by Abu-Zahra et al.50 and Oexmann et al.51 Comparisons of capital costs, operating costs, fuel costs, capture costs, and energy requirements between these studies and this paper are shown in Table S1 in the Supporting Information. As shown in Table S1 in the Supporting Information, the capture plant capital cost used in this paper (US$1,415/kW) is higher than that of Abu-Zahra et al.50 and Oexmann et al.51 Adjusting for the difference in the costing year and currency, the capital cost of Abu-Zahra et al. is approximately US$722/ kW, while the capital cost of Oexmann et al. is US$1,072/kW. There is a difference in the reported costs because of the differences in the process and cost boundaries between this paper and those conducted by Oexmann et al. and Abu-Zahra 16892



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Figure 4 shows the impact of changes in the solvent working capacity due to changes in the rich loading on the capture cost

pretreatment processes (US$65 per metric ton of CO2 avoided). Comparison with Other Capture Technologies. Table S2 in the Supporting Information shows the capture cost comparison of different capture technologies. A technoeconomic analysis of CO2 capture from the emission sources with different feed gas CO2 compositions using unpromoted MEA-based chemical absorption, membrane separation, pressure swing adsorption (PSA), and vacuum swing adsorption (VSA) was carried out by Hasan et al.52,53 The results from Ho et al.32,33 that examined the capture cost using membrane separation, PSA, and VSA are also shown in Table S2 in the Supporting Information. The baseline MEA capture cost from this paper (US$67−88 per metric ton of CO2 avoided), which is equal to US$47−59 per metric ton of CO2 captured, is slightly higher than the value reported by Hasan et al. for absorption of US$40 per metric ton of CO2 captured. Hasan et al. assumed that the flue gas does not contain SOx and NOx, which is comparable to a value of US$47 per metric ton of CO2 captured in this paper when the flue gas is assumed to not contain SOx and NOx. The small difference in the cost may be due to the difference in the cost year. These values are lower than the capture cost using membrane, PSA, and VSA technologies, which are approximately US$45−70 per metric ton of CO2 captured as reported by Hasan et al.52,53 and Ho et al.32,33 Both research groups have suggested that adsorption and membrane based separations may be more competitive for capturing CO2 from flue gases that have higher CO2 concentrations such as those from cement or iron and steel production.54 Improving Solvent Properties. The Australian government recently introduced a carbon tax of A$23 per metric ton of CO2 emitted in 2012. Even though the tax is likely to be changed, any value currently under consideration is substantially lower than the current cost of CO2 capture using solvents such as MEA or advanced amines. Thus, improvements in solvent technology are essential to make CCS competitive as a greenhouse gas mitigation option. There are several options for reducing the capture cost through improvements of solvent properties. These include improving solvent loading, reducing the heat of reaction, improving solvent concentration, reducing solvent loss, and improving the solvent stability to impurities. The following sections evaluate the effect on cost reductions as solvent properties are improved. The baseline properties of the hypothetical solvents are assumed to be similar to unpromoted MEA, as shown in Table 1. In the following analysis, one parameter of the hypothetical solvent is varied at a time, with the other parameters held constant. The process parameters (absorber temperature, absorber pressure, stripper pressure, and stripper temperature) are assumed to be constant as each solvent property is varied. In this analysis, the capture plant is designed without including advanced heat integration with the power plant. Solvent Loading. Solvent loading is an important parameter because it determines the required solvent circulation rate and the mass transfer rate, thus affecting the regeneration energy requirement and the absorber size. The lean solvent loading is the ratio of the amount of CO2 to the amount of solvent as it enters the absorber. The rich solvent loading is the ratio of the amount of CO2 to the amount of solvent leaving the absorber. The difference between rich solvent loading and lean solvent loading is the solvent working capacity.

Figure 4. Effect of solvent working capacity on capture cost and thermal regeneration energy when the rich loading is varied and other properties are held constant.

and thermal regeneration energy while all of other parameters including the lean loading and the corresponding partial pressure of CO2 in the solution at the top and the bottom of the absorber are held constant. Thus, the rich loading is varied from the baseline value of 0.455 mol of CO2/mol of solvent assuming that the corresponding partial pressure of CO2 in the solution at the bottom of the absorber remains unchanged at 0.8 kPa. The changes in cost of capture and thermal regeneration energy due to changes in the lean loading (and hence the working capacity) are shown in Figure 5. Other parameters are

Figure 5. Effect of solvent working capacity on capture cost and thermal regeneration energy when the lean loading is varied and other properties are held constant.

held constant including the rich loading, the corresponding partial pressure of CO2 in the solution at the top and the bottom of the absorber, and the partial pressure ratio of water to CO2 at the top of the stripper. It is important to note that the thermodynamic profile of the hypothetical solvent is assumed to be such that the various partial pressures can remain unchanged even though the lean loading is varied. As shown in Figure 4, the capture cost and regeneration energy decrease when the solvent working capacity is increased by increasing the rich solvent loading. The required solvent circulation rate decreases when the solvent working capacity increases because less solvent is needed to absorb the same 16893



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amount of CO2. Thus, the regeneration energy in the stripper also decreases. For a higher rich loading, a lower solvent circulation rate results in lower sensible heat, while the heat of reaction and heat of vaporization are constant (eq 6). Because the regeneration energy (Figure 2) is the largest contributor to the capture plant energy requirement (approximately 56%), the capture cost tends to decrease as the regeneration energy decreases. Moreover, as shown in Figure 2, the largest capital cost contributor in the capture plant is the absorption unit, which consists of the absorber column, stripper column, reboiler, and heat exchangers. Improving solvent working capacity decreases the solvent flow rate and increases the mass transfer (because of higher “Ce − Cs” in eq 1), thus reducing the equipment size. At a solvent capacity of 0.5 mol of CO2/mol of solvent, which is more than double the baseline value, the capture cost decreases by about 19% from US$88 per metric ton of CO2 avoided to about US$71 per metric ton of CO2 avoided. As shown in Figure 4, the capture cost decreases most rapidly between solvent working capacities of 0.1 and 0.4 mol/mol, where the decrease in regeneration energy is highest. Similarly, the capture cost and regeneration energy decrease if the solvent working capacity is increased by decreasing the lean solvent loading, as shown in Figure 5. The higher solvent working capacity again leads to a lower solvent circulation rate, which leads to a lower sensible heat (eq 2). Thus, the cost of capture can be reduced by using solvents with a large working capacity. This can be achieved by developing a solvent with low lean loading or high rich loading. The rich loading is limited by the equilibrium solubility of CO2 in the solvent, which in turn is a property of a solvent. The rich loading should be set as high as possible, but it must also be below the partial pressure of CO2 in the flue gas in order to provide sufficient driving force for CO2 absorption. Moreover, a higher rich loading means a higher CO2 concentration in the rich solution, which potentially affects the corrosivity of the rich solvent stream. This needs to be taken into account when selecting the material of construction for the absorber system. On the other hand, the lean loading should be set as low as possible by adjusting the partial pressure ratio of water to CO2 at the top of the stripper. This is important to prevent excessive heat of regeneration due to vaporization in the stripper (for example, the lean loading for MEA is usually limited at about 0.2−0.3 mol of CO2/mol of solvent). Heat of Reaction. Another opportunity to reduce the regeneration energy is by reducing the energy required to reverse the CO2 reaction with the solvent. Note that the results reported here are for a constant stripper pressure which is different from the results reported by Oexmann et al.,21 who examined operating solvents with a high heat of reaction at higher pressure to reduce the regeneration energy. Assuming that a reduction in heat of reaction does not adversely affect the overall rate of reaction and that the partial pressure ratio of water to CO2 in the stripper top is constant, the effect of heat of reaction on capture cost and regeneration energy is shown in Figure 6. The capture cost decreases from US$88 to US$79 per metric ton of CO2 avoided (a 10% reduction) if the heat of reaction halves. In practice, the benefit of having a solvent with a low heat of reaction may require adjustment of the solvent loading. This is important to prevent excessive water vaporization caused by the change in the partial pressure ratio of water to CO2 at the top

Figure 6. Effect of heat of reaction on capture cost and thermal regeneration energy with other properties held constant.

of the stripper. This trade-off has been applied for solvents such as potassium carbonate and ammonia. Solvent Concentration. In the aqueous chemical absorption process, solvent concentration is normally limited to prevent corrosion. The solvent is dissolved in another liquid, commonly water. For example, the currently available unpromoted MEA solvent uses 30 wt % MEA in water with a corrosion inhibitor. Improvement in this solvent property offers another opportunity to reduce the regeneration duty and the absorption unit size. Figure 7 shows that both the capture cost and regeneration energy decrease with increasing solvent concentration, which

Figure 7. Effect of solvent concentration on capture cost and thermal regeneration energy with other properties held constant.

reduces the required solvent circulation rate because a smaller amount of solvent solution is required to absorb the same amount of CO2. The lower solvent circulation rate leads to a lower sensible heat that needs to be supplied in the stripper (eq 4). As shown by eq 1, the required absorption unit size will also reduce with increasing solvent concentration because the solvent flow rate decreases and the mass transfer increases (because of higher M in eq 1). As shown in Figure 7, the capture cost decreases most rapidly between concentrations of 10 and 50 wt %. If the solvent concentration increases from 30 to 50 wt %, the capture cost decreases by about 16% from US$88 to US$72 per metric ton of CO2 avoided. In aqueous chemical absorption systems, high solvent concentrations up to 50 wt % have already been used with methyldiethanolamine (MDEA).55 Furthermore, GE 16894



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Global Research in collaboration with GE Energy and the University of Pittsburgh has recently developed new solvents for postcombustion CO2 capture with high solvent concentration.56 That study suggests that aminosiloxanes solvents have higher CO2 loading, higher heat of reaction, and smaller solvent heat capacity than MEA. Mass Transfer Rate. The mass transfer rate is one of the key solvent properties in the absorption process. This property determines the required absorber surface area, thus affecting the capture plant capital cost. In fact, MEA is used commercially because of its fast mass transfer rate despite its other drawbacks. Improvements in mass transfer rate can be achieved by adding promoters (such as piperazine). The relationship between mass transfer rate and capture cost is presented in Figure 8 with other properties held constant.

Figure 9. Effect of solvent loss and stability to SOx and NOx on capture cost at various solvent prices with other properties held constant.

replacement only accounts for a small fraction of the operating cost at this price (Figure 2). For a more expensive solvent costing US$7.5/kg, the same reduction in solvent loss results in a 2% reduction in capture cost from US$88 to US$86 per metric ton of CO2 avoided. However, a more substantial reduction in capture cost can be achieved by developing solvents that tolerate high levels of SOx and NOx in the flue gas and eliminate the need for the FGD and the SCR units. To achieve this, the solvent system would need to simultaneously produce byproducts with the SOx and NOx, while not affecting the CO2 capture efficiency or causing solvent degradation; e.g., in potassium carbonate solvent, SOx and NOx form potassium sulfate and potassium nitrate that can be separated as byproducts. As shown in Figure 2, the FGD and SCR pretreatment cost is the second largest contributor to the capture plant capital cost. At a solvent cost of US$1.5/kg and a solvent loss of 0.2 kg/ metric ton of CO2, removing the FGD and SCR facilities from the pretreatment step and handling the SOx and NOx elsewhere may reduce the capture cost by 28% to US$63 per metric ton of CO2 avoided. Improving Process Design. There are several options for reducing the capture cost by improving the process and equipment design. These include operating the stripper at a higher pressure, improving the efficiency of the column packing, developing advanced heat exchanger designs to reduce the temperature difference across the stripper, and using alternative materials for the process vessel equipment. Stripper Pressure. Operating the stripper at a high pressure is potentially one method to reduce the energy requirement of absorption-based CO2 capture. The effect of stripper pressure on the capture cost and electrical energy requirement for a hypothetical solvent (with heat of reaction of 82 kJ/mol CO2) is shown in Figure 10. The estimated stripper temperature at different stripper pressures is shown in Figure 11. The stripper temperature at elevated pressure is estimated using the equation reported by Oyenekan et al. for the generic solvent.57 The hypothetical solvent is assumed to not degrade even at very high stripper temperatures. In practice, the pressure of the stripper is limited by the maximum temperature permitted to prevent solvent degradation. As reported by Rochelle et al.,58 typical linear alkanolamines and diamines degrade by polymerization and urea formation at 100−130 °C, tertiary amines degrade by arm switching and elimination at 120−140 °C, and piperazine and related cyclic amines degrade by ring opening at 150−165 °C.

Figure 8. Effect of mass transfer rate on capture cost and absorber height with other properties held constant.

The cost is observed to decrease from US$88 to US$85 per metric ton of CO2 avoided when the mass transfer rate is up to 50% faster than MEA. On the other hand, the capture cost increases substantially as the mass transfer rate decreases, especially if the solvent mass transfer rate is 50% slower than MEA. This means that significantly improving the mass transfer rate above that of MEA will not have a significant impact on cost. Moreover, Figure 8 also indicates that the mass transfer rate should not be reduced significantly in an effort to improve other solvent properties, such as solvent loading and solvent concentration. Solvent Stability. One of the main components of operating costs for the absorption plant is the replacement cost of solvent due to solvent loss. This loss occurs mainly due to evaporation in the absorber and stripper columns. Moreover, around 15− 25% of amine-based solvent loss occurs due to solvent degradation, particularly because of irreversible reactions between the solvent and the acidic components of the flue gas, such as SOx and NOx, that form stable salts. The formation of stable salts has other adverse effects on the performance of the absorption process; e.g., the degradation products are thought to increase corrosion.42 The relationship between solvent losses and capture cost is presented in Figure 9. At a cost of US$1.5/kg, solvent loss only has a marginal effect on the capture cost. For example, an 8-fold reduction in the solvent loss from the baseline value of 1.6 to 0.2 kg/metric ton of CO2 (a value similar to KS-1) only results in a reduction in capture cost of approximately 5% from US$88 to US$84 per metric ton of CO2 avoided. This is because the cost of solvent 16895



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to the minimum capture cost at this pressure. This agrees with results reported by Lucquiaud et al.,48 who reported that solvent regeneration at 150 °C has the highest thermal efficiency for heat recovery. Column Packing. Packing is required inside the absorber and stripper columns to increase the gas−liquid contact area. Currently, structured packing is used instead of random packing due to its higher specific surface area.60−62 The packing efficiency affects the mass transfer rate and pressure drop in the column, thus affecting the column dimensions.61 The pressure drop inside the column also affects the size and the energy consumption of the flue gas blower.47 A number of researchers have developed new types of packing with the potential to increase the mass transfer rate.63−65 Figure 12 shows the effect of mass transfer increase (as a result of packing improvement) on the capture cost, assuming Figure 10. Effect of stripper pressure on capture cost and thermal regeneration energy with other properties held constant.

Figure 12. Effect of mass transfer rate and packing efficiency on capture cost with other properties held constant.

that any change in the resistance of the packing to fluid flows is not affected by the mass transfer rate. The capture cost only decreases from US$88 to US$84 per metric ton of CO2 avoided (a 4% reduction) if the mass transfer rate doubles. This is due to an approximately US$20 million capital cost reduction as a result of smaller absorber and stripper columns. Further mass transfer rate improvement does not affect the capture cost reduction substantially; for example, a 5-fold increase in the mass transfer rate only reduces the capture cost to US$82 per metric ton of CO2 avoided (a 7% reduction). Therefore, the development of new types of structured packing, even with superior mass transfer properties, is unlikely to achieve a reduction in capture cost as substantial as that achieved by improvement in solvent properties. Heat Exchanger Design. Another opportunity to reduce the capture cost is by reducing the energy requirement in the stripper, which is affected by the temperature of the incoming stream. As shown in eq 2, the regeneration energy requirement can be reduced by lowering the temperature difference across the stripper, which will reduce the sensible heat inside the stripper. This might be achieved by lowering the temperature difference in the cross heat exchanger or by using novel stripper designs where the stripper and heat exchanger are integrated.43 Currently, the most advanced absorption process has a temperature difference between the stripper inlet and outlet streams (ΔTinlet/outlet stripper) of approximately 15 °C. Figure 13 shows the effect of the temperature difference between the stripper inlet and outlet streams on the regeneration energy and the capture cost. If the temperature difference is reduced from

Figure 11. Effect of stripper pressure on stripper temperature and electrical energy requirement with other properties held constant.

As the stripper pressure increases from 2 to 10 bar, Figure 10 shows that the capture cost decreases from the baseline US$88 per metric ton of CO2 avoided to US$82 per metric ton of CO2 avoided (6% reduction). Despite the slight increase in sensible heat (because of higher specific heat at higher temperature) and heat of reaction (as reported by Gupta et al.59), the total thermal regeneration energy at 10 bar (i.e., about 147 °C) decreases by approximately 18% from the baseline value (as shown in Figure 11). Because the term |ΔHvap,H2O| − |ΔHabs| in eq 7 is negative, the partial pressure ratio of water to CO2 decreases, so there is less water vaporization and the heat of vaporization decreases from 1.45 to 0.61 MJ/kg CO2. Above 10 bar, the capture cost increases with stripper pressure and reaches US$86 per metric ton of CO2 avoided at 30 bar, despite the decrease in total thermal regeneration energy (3.39 MJ/kg CO2). The capture cost increases because of the increasing total electrical energy requirement (as shown in Figure 10). The reboiler always remains at boiling conditions; thus if the stripper is operated at higher pressure and temperature, the steam that is drawn from the LP/IP crossover must also be at higher pressure and temperature. According to eqs 9 and 10, this leads to an increase in power loss from the power plant. The minimum total electrical energy requirement is observed at the stripper pressure of about 10 bar and stripper temperature of about 147 °C, which corresponds 16896



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Table 2. Hypothetical Solvent Baseline Case Properties and Sensitivity Analysis Basis parameter working capacity (mol/mol) solvent concn (wt %) heat of reaction (MJ/kmol CO2) mass transfer rate (kmol/m3·h· kPa) solvent loss (kg/ metric ton CO2) solvent stability

Figure 13. Effect of stripper temperature difference between the stripper inlet and outlet streams on capture cost and regeneration energy with other properties held constant.

15 to 5 °C, the capture cost decreases by about 10% to US$79 per metric ton of CO2 avoided. This is mainly related to a 22% reduction in the regeneration energy required by the stripper. By reducing the temperature difference across the stripper from 15 to 5 °C, the sensible heat decreases from 1.07 to 0.36 MJ/kg CO2. Because the stripper pressure and temperature at the bottom of the column is held constant, the stripper top temperature increases when the temperature across the stripper decreases. This reduces the heat of vaporization (from 1.45 to 1.19 MJ/kg CO2) because of the lower partial pressure ratio of water to CO2 at constant solvent loading (eq 7). Process Vessel Material. Because the absorption unit (absorber, stripper, reboiler, and heat exchangers) accounts for about 35% of total capital costs (Figure 2), using lower cost construction materials instead of stainless steel might lower capture costs.66 Figure 14 shows the difference in the capture

−50% baseline

baseline

stripper press. (bar) packing ΔT inlet/outlet stripper (°C) process vessel material

+50% baseline

process improvement

0.226

0.113

0.339

−

30

15

45

−

82

41

123

−

1 × MEA

0.5 × MEA

1.5 × MEA

−

1.6

0.8

2.4

−

require FGD and SCR 2

−

−

no FGD and SCR

1

3

−

Mellapak 250Y 15

−

−

7.5

22.5

hypothetical packing −

SS-304

−

−

concrete

Table 3. Monte Carlo Simulation Conditions parameter

minimum likeliest

heat of reaction (kJ/mol CO2) solvent loss (kg/metric ton CO2) ΔT inlet/outlet stripper (°C) stripper pressure (bar)

50

70

85

0.2

0.4

1.6

5 1.5

parameter working capacity (mol of CO2/mol of solvent) concn (wt %) mass transfer rate ratio to MEA parameter solvent stability to SOx and NOx

maximum

10 15 2 10 95% uncertainty mean range 0.35

±0.15

40 0.75 case 1

±20 ±0.25

FGD and SCR

distribution

triangular

distribution normal

case 2 no FGD and SCR

The results of the sensitivity analysis in Figure 15 show that variation in the baseline value for the heat of reaction, the temperature difference across the stripper, the solvent concentration, and the solvent working capacity have a large impact on cost. The stripper pressure, mass transfer rate, and

Figure 14. Effect of process vessel material on capture cost and column cost with other properties held constant.

cost and the capital cost of the absorber and stripper vessels when they are constructed using either 304 stainless steel or geopolymer concrete; the difference is about 10%. Parameter Ranking. In order to rank the effectiveness of solvent and process design improvements on capture cost, both a sensitivity analysis and a Monte Carlo simulation were completed. The Monte Carlo simulation was conducted to explore the cost of capture using the hypothetical solvents and to estimate what the cost of absorption-based CO2 capture could be if new solvents with more desirable properties are developed. Tables 2 and 3 show the range of values used for each of those analyses.

Figure 15. Tornado chart for effect of solvent properties and process design on capture cost. 16897



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stripper temperature difference are based on data from Peeters et al.43 In varying the mass transfer rate in the Monte Carlo simulation, the data set is assumed to be distributed within the range where the rate is equivalent to that of MEA and half of MEA. The upper limit is assumed to be equivalent to MEA because MEA is considered to be one of the fastest solvents for CO2 absorption, and thus new solvents are most likely to have mass transfer rates up to that of MEA. Meanwhile, the lower limit of the confidence level is set as equivalent to half of that of MEA as mass transfer rates below this value are uneconomical (Figure 8). The likeliest value for the stripper pressure is assumed to be that of unpromoted MEA solvent system (2 bar), while the minimum value is assumed to be similar to the low end for typical amine solvents. The maximum value for the stripper pressure is assumed to be 10 bar, as this value represents the optimum stripper operating pressure in order to achieve the minimum capture cost, as shown in Figure 10 (assuming the stripper pressure is not limited by the solvent degradation). The capture cost distribution obtained from the Monte Carlo simulation is shown in Figure 16. If the capture plant requires

type of packing have a modest impact. Meanwhile, the solvent loss has the least impact on cost. The solvent stability to SOx and NOx and the use of geopolymer concrete as vessel materials are binary values, which means that the capture plant either has them or does not have them. If the solvent is stable to SOx and NOx, the cost reduction is significant, even higher than the most sensitive parameter, solvent concentration. In the Monte Carlo simulation, four parameters are varied using a triangular distribution (heat of reaction, solvent loss, temperature difference across the stripper, and pressure of the stripper) and three parameters are varied using a normal distribution (solvent concentration, solvent working capacity, and mass transfer ratio to MEA). The data distribution shapes for each parameter are selected based on an evaluation of the likeliest variation of each solvent property from various solvents. The standard deviation for the normal distribution is set so that 95% of the data falls within the maximum and minimum range that is shown in Table 3. Two series of Monte Carlo simulations were completed: one where the solvent is stable to SOx and NOx and does not require FGD and SCR, and another series where pretreatment with FGD and SCR is required. The vessel material and structured packing are omitted from the Monte Carlo simulation. Given that the variability of the vessel material is an on/off function between 304 stainless steel and geopolymer concrete, this parameter is omitted. The type of packing is also omitted due to its relatively small impact on capture cost. The mean and standard deviation for solvent working capacity are selected so that the high end of the data is in the range of the values for methyldiethanolamine (MDEA) and diethanolamine (DEA) solvents. This also represents the working capacity at which reductions in the cost start to become minimal, as shown in Figure 4. The working capacity of MEA solvent is near the low end of the distribution. The mean value is assumed to be the median between these two values, which is similar to the values of 2-amino-2-methyl-1isopropanol (AMP) and diglycolamine (DGA) solvents. The maximum value of heat of reaction assumed puts the baseline MEA value at the high end of the distribution. The minimum value for the heat of reaction is assumed to be that of potassium carbonate solvent, while the likeliest value is assumed to be similar to that for most amine solvents, such as piperazine, MDEA, methylpiperazine (MPZ), and AMP. The mean and standard deviation for solvent concentration are selected so that the high end of the data distribution is similar to the values of MDEA, DGA, and Cansolv.52 This value represents the highest solvent concentration currently used in CO2 removal aqueous solvent systems. As shown in Figure 6, the cost reduction appears to diminish at this point. Meanwhile the low end of the data range is similar to the solvent concentration in a generic MEA application (without corrosion inhibitor). The mean value is assumed to be the median between these two values, which is similar to the solvent concentrations of AMP, piperazine activated potassium carbonate, and DEA. For the solvent loss, it is assumed that the maximum value for this parameter is equivalent to that typical for an unpromoted MEA process. The minimum value is based on the value for KS-1, while the likeliest value is assumed to be the median between the maximum value and the minimum value, which is in a similar range to solvent loss for DGA and DEA. The likeliest value, maximum value, and minimum value for the

Figure 16. Monte Carlo simulation result for capture cost, within a 95% confidence level (CL).

FGD and SCR, the cost is estimated to be between US$62 and US$80 per metric ton of CO2 avoided within a 95% confidence level, with the most likely cost of US$71 per metric ton of CO2 avoided. On the other hand, if the capture plant does not require FGD and SCR, the capture cost is estimated to be between US$44 and US$59 per metric ton of CO2 avoided, with the most likely cost of US$52 per metric ton of CO2 avoided. The capture cost distribution range within a 95% confidence level for the capture plant with FGD and SCR is wider than the distribution range for the capture plant without FGD and SCR (US$28 versus US$25 per metric ton of CO2 avoided). This is because when FGD and SCR facilities are utilized, the energy requirement increases and hence the CO2 avoided decreases as shown in eq 8. The decrease in the amount of CO2 avoided coupled with the increase in capital cost widens the range of capture costs. For the flue gas evaluated in this paper, developing new solvents with good stability to SOx and NOx is an important parameter. As shown in Figure 16, the high end of costs for a capture plant that does not require FGD and SCR is US$63− 65 per metric ton of CO2 avoided. The corresponding solvent properties for this range are such that the solvent has good stability to SOx and NOx, a working capacity of lower than 0.3 16898



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plant does not require FGD and SCR, capture cost is estimated to be between US$44 and US$59 per metric ton of CO2 avoided, with the most likely cost of US$52 per metric ton of CO2 avoided. Significant cost reductions can be achieved if new solvents have good stability to SOx and NOx, can be operated under high concentration (above 50 wt %), and have a high working capacity (above 0.35 mol of CO2/mol of solvent) while maintaining a mass transfer rate comparable to MEA. In the Monte Carlo simulation, the individual parameters are assumed to be completely independent of each other (not cross-correlated). In reality, there might be a cross-correlation between some parameters; for example, a solvent with a high heat of reaction tends to have a high mass transfer rate. Including cross-correlation effects could change the shape of the resulting probability density distribution. However, this is not expected to change the conclusions of this paper because the suggested key solvent properties (solvent stability to SOx and NOx, solvent concentration, and solvent working capacity) do not have a cross-correlation with other properties. The results presented in this paper suggest that the direction for future solvent improvement is in accordance with the current research direction in solvent development, which focuses on the development of solvents that involve a phase change during CO2 absorption. Some examples of these types of solvents are amino acid solvents,28−30 precipitating potassium carbonate solvent,27 and the chilled ammonia process.31 The main advantages of using these types of solvents are the higher solvent capacity because of the precipitation process based on Le Chatelier’s principle. These solvents also have a better stability to SOx and NOx and can remove SOx and NOx as well as CO2 from the flue gas and so do not require FGD and SCR facilities. Amino acid solvents also have better ecotoxicity and better corrosion rates compared to MEA, thus enabling a higher solvent concentration. Therefore, future work on this project will focus on modeling this new generation of solvents, in order to gain a better understanding of their cost drivers and their disadvantages compared to aqueous solvents.

mol of CO2/mol of solvent, a solvent concentration of 20% or lower, and a mass transfer rate equivalent to 50−80% of MEA. For the capture plant that does require FGD and SCR, the lower end of the cost range is also within this bracket of US $63−65 per metric ton of CO2 avoided. The solvent properties that correspond to this range of costs are a solvent that does not have good stability to SOx and NOx, has a working capacity higher than 0.35 mol of CO2/mol of solvent, has a solvent concentration higher than 50%, and has a mass transfer rate that is equivalent to MEA. This means that the benefit of having a solvent with good stability to SOx and NOx, even when other solvent properties are moderately good, is approximately equal to having a solvent with superior properties that does not tolerate high levels of impurities. In order to achieve the lowest cost estimate of US$44−46 per metric ton of CO2 avoided in the results for the Monte Carlo simulation, the corresponding solvent properties are such that the solvent must have good stability to SOx and NOx, a working capacity of 0.35 mol of CO2/mol of solvent or higher, a solvent concentration of 50 wt % or higher, and a mass transfer rate that is equivalent to MEA. Model Limitations. The economic analysis of the chemical absorption process presented in this paper is a preliminary assessment and, as such, is only indicative of the possible cost benefits that can be achieved. Simplifying assumptions have been made based on standard rules of thumb. Variations in the assumptions or different economic conditions will change the results presented in this paper.

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CONCLUSION The results presented in this paper show that solvent stability to SOx and NOx, heat of reaction, solvent concentration, and solvent working capacity have the largest influence on CO2 capture cost (among the solvent properties). Considering its relatively small impact on cost, the solvent loss can be compromised if all of the solvent properties cannot be improved concurrently, unless the cost of new solvent is very high. In practice, a solvent property cannot be individually improved without taking into account the effect on other solvent properties. For example, the benefit of having a solvent with a low heat of reaction may require changes in the solvent loading to limit water vaporization. The operation of the stripper at a higher pressure, the use of concrete for process vessels, reducing the temperature difference across the stripper, and improvements in column packing also have cost benefits. Among these process design improvements, the highest potential cost benefit can be achieved by reducing the temperature difference across the stripper. The relative effects of the other process design improvements on capture cost reduction are smaller than the effect from improving solvent properties. Due to the high quality coal and the less stringent regulation of SOx and NOx emission, none of the existing power plants in Australia has FGD and SCR units. Therefore, these processes must be included as part of the pretreatment process of the capture plant. However, for regions where the FGD and SCR are already part of the power plant, the other insights of this paper are still applicable. Based on the Monte Carlo simulation results, capture cost is estimated to be between US$62 and US$80 per metric ton of CO2 avoided within a 95% confidence level if the capture plant requires FGD and SCR, with the most likely cost of US$71 per metric ton of CO2 avoided. On the other hand, if the capture

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ASSOCIATED CONTENT

S Supporting Information *

Tables showing detailed baseline results and cost comparisons to other capture technologies from other literature. This material is available free of charge via the Internet at http:// pubs.acs.org.

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AUTHOR INFORMATION

Corresponding Author

*Tel.: +61 2 9385 4755. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors would like to acknowledge the funding provided by the Australian Government through its CRC program to support this CO2CRC research project. A.R. would also like to acknowledge the University of New South Wales and the Faculty of Engineering for scholarship funding.

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REFERENCES

(1) Sierra, K.; Kyte, R.; Bond, J. Development and Climate ChangeA Strategic Framework for the World Bank Group; The International Bank for Reconstruction and Development/World Bank: Washington DC, 2008.

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