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Energy and the Environment

Systems Analysis of Physical Absorption of CO2 in Ionic Liquids for Pre-Combustion Carbon Capture Haibo Zhai, and Edward S. Rubin Environ. Sci. Technol., Just Accepted Manuscript • DOI: 10.1021/acs.est.8b00411 • Publication Date (Web): 28 Mar 2018 Downloaded from http://pubs.acs.org on March 29, 2018

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Environmental Science & Technology

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Systems Analysis of Physical Absorption of CO2 in Ionic Liquids for Pre-Combustion Carbon Capture

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Haibo Zhai* and Edward S. Rubin

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Department of Engineering and Public Policy

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Carnegie Mellon University, Pittsburgh, PA15213

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*Corresponding Author Email: [email protected]; Phone: (412) 268-1088

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Revised for publication in Environmental Science and Technology

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ACS Paragon Plus Environment


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ABSTRACT

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This study develops an integrated technical and economic modeling framework to investigate the

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feasibility of ionic liquids (ILs) for pre-combustion carbon capture. The IL 1-hexyl-3-

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methylimidazolium bis(trifluoromethylsulfonyl)imide is modeled as a potential physical solvent

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for CO2 capture at integrated gasification combined cycle (IGCC) power plants. The analysis

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reveals that the energy penalty of the IL-based capture system comes mainly from the process

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and product streams compression and solvent pumping, while the major capital cost components

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are the compressors and absorbers. Based on the plant-level analysis, the cost of CO2 avoided by

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the IL-based capture and storage system is estimated to be $63 per tonne of CO2. Technical and

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economic comparisons between IL- and Selexol-based capture systems at the plant level show

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that an IL-based system could be a feasible option for CO2 capture. Improving the CO2 solubility

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of ILs can simplify the capture process configuration and lower the process energy and cost

31

penalties to further enhance the viability of this technology.

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TABLE OF CONTENTS ART

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INTRODUCTION

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Ionic liquids (ILs) are under active development as one of emerging technologies for carbon

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dioxide (CO2) capture because of their favorable properties, such as relative nonvolatility, high

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CO2 solubility and selectivity, and endless tenability.1-3 ILs also can be used for CO2 capture

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without water dilution, which reduces the parasitic load of capture systems.1 However, their high

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viscosity poses a challenge for large-scale applications.

42

Research on ILs has been emphasized on laboratory experiments, molecular and phase

43

equilibrium simulations, and materials synthesis.2,4-7 More novel ILs with high capacity and low

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viscosity are being synthesized, especially for post-combustion CO2 capture.8-10 While research

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activities on ILs focus mainly on the application to post-combustion CO2 capture,11-14 there have

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been relatively few studies of ILs for pre-combustion CO2 capture at integrated gasification

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combined cycle (IGCC) power plants. The shifted gas stream consists mainly of CO2 (~38%)

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and hydrogen (H2) (~56%), with CO2 partial pressures that can be several tens of bars at IGCC

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plants. So, ILs also appear attractive as a potential physical solvent for pre-combustion CO2

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capture.15 Shiflett and Yokozeki measured and predicted the solubility of CO2 in 1-hexyl-3-

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methylimidazolium bis(trifluoromethylsulfonyl)imide ([hmim][Tf2N]).16 Recent computational

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studies demonstrated the technical feasibility of absorption processes using [hmim][Tf2N] and

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two TEGO ILs as a physical solvent for capturing CO2 at an IGCC plant, but lacked cost

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evaluation.17-18 García-Gutiérrez et al. also demonstrated the feasibility of using ILs as a physical

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solvent for capturing CO2 from biogas streams.19

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A variety of thermodynamic models have been developed to predict the phase behavior of

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CO2 and other gases in various ILs, such as classical cubic equations, activity coefficient and

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group contribution methods, and quantum chemistry calculations,3,20-23 and data mining



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algorithms.24-25 However, much less attention has been paid to the simulation and design of IL-

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based capture systems,4 and the cost of IL-based capture systems has rarely been evaluated. The

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major objectives of this study, therefore, are to evaluate the techno-economic feasibility of ILs

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for pre-combustion CO2 capture at IGCC plants and to explore the measures needed to improve

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the viability of this technology.

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To achieve these objectives, this study develops an integrated technical and economic

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modeling framework for assessing pre-combustion capture processes, with a focus on the

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separation of CO2 and H2. Given the technical feasibility for CO2 capture,16-17 the ionic liquid

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[hmim][Tf2N] is selected for the assessment and serves as a proxy solvent for “back engineering”

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to explore desired IL properties and their potential role in improving overall process viability.

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Quantifying potential benefits from improved solvent properties is helpful to guide the

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development of new materials for carbon capture.

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MATERIALS AND METHODS

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This section of the paper presents details of the performance and cost models developed to

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analyze the IL-based option for IGCC plants. Subsequent sections present plant-level results and

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comparisons with current commercial alternatives.

76 77 78

Performance Modeling of CO2 Capture System The performance modeling framework for pre-combustion CO2 capture mainly includes the

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phase equilibrium model, plus the mass and energy balance models for an IL-based capture

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system.

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Physical Properties of Solvent and Gases Solvent properties correlate with temperature and/or pressure. Basha et al. developed

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correlations for density, surface tension, viscosity, and heat capacity for [hmim][Tf2N] based on

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data from the literature.17 The vapor pressure of [hmim][Tf2N] is extremely low and negligible. 17

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The correlations of density, viscosity, and heat capacity with temperature and pressure are

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formulated for CO2 and H2 based on data available from the National Institute of Standards and

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Technology's properties database for fluid systems.26 For CO2, the gas-phase diffusivity is

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predicted using the correlation equation developed by Fuller et al.,27 whereas the liquid-phase

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diffusion coefficient is estimated as a function of the solvent viscosity.28 Solvent and gas

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properties are summarized in Tables S1, S2, and S3 of the Supporting Information (SI).

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Phase Equilibrium based on Equation of State

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The generic Redlich-Kwong Equation of State (RK EOS) relates temperature, pressure, and

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volume of gases, can be applied to predict the phase behavior of gases in solvents, such as ILs.16

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The modified RK EOS is expressed as:16,29 P =

RT a(T) V-b VV+b

(1)

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For N-component mixtures, mixing rules are applied to estimate the parameters (a, b) in the RK

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EOS as: 16,29 a= ai aj fij (T)1-kij xi xj N

(2)

i,j=1

2 R2 Tc,i ai =0.427480 α T Pc,i i

(3)



αi T=

βk Tc ⁄T - T⁄Tc , ≤3

k

k=0 β0 +β1 exp21- T⁄Tc -1 , τij fi,j T=1+ T lij lji xi +xj kij = lji xi +lij xj

T⁄Tc ≤1

T⁄Tc >1

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(4)

(5) (6)

1 b= (bi +bj ) 1-kij 1-mij xi xj 2 N

(7)

i,j=1

bi =0.08664

RTc,i Pc,i

(8)

99 100

where Tc,i is the critical temperature of the ith species (oK), Pc,i is the critical pressure of the ith

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species (kPa), R is the universal gas constant and xi is the mole fraction of the ith species. There

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are four binary-interaction parameters for each pair of components in the system: , , , and

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. However, two or three parameters are sufficient for most equilibrium applications.29 The

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fugacity coefficient of the ith component derived from the RK EOS is described as:29 ln∅i = ln

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1 RT a a aí bi V + - +1 ln +bi V-b RTb(V+b) RTb a b P(V-b) V+b

in which, á i =2 ai aj fij xj 1-kij N

j=1

lij lji lij -lji xi xj lji xi +lij xj

b i = (bi +bj )1-mij xj 1-kij N

j=1

106 107

(9)

2

-a

lij lji lij -lji xi xj lji xi +lij xj

2

-b

For a component (), the K-value of vapor-liquid equilibrium (VLE) is calculated as the ratio of liquid versus vapor fugacity coefficients (∅Li , ∅Vi ) from Equation (9):

∅iL Ki = V ∅i

(10)


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The critical component parameters and constants in the VLE model are summarized in Table

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S4 of the SI. The binary-interaction parameters can be estimated by correlating the experimental

110

solubility data with the RK EOS. The binary parameters of interactions between CO2 and

111

[hmim][Tf2N] and between CO2 and H2 come from the estimates by Shiflett and Yokozeki.16, 29

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The binary interaction parameters for H2 and [hmim][Tf2N] are determined by fitting the

113

solubility data available from the literature.30 The binary-interaction parameters used for each

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pair are summarized in Table S5 of the SI. The phase equilibrium model of CO2 and H2 in

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[hmim][Tf2N] was incorporated in the absorption and stripping process simulation.

116 117 118

Multistage Equilibrium Process Model for Gas Absorption Gas absorption using [hmim][Tf2N] for CO2 capture is treated as a steady-state vapor-liquid

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process that consists of multiple equilibrium stages. For any stage in a countercurrent cascade,

120

vapor and liquid streams leave in phase equilibrium. Then, a multi-stage equilibrium model is

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developed to delineate the absorption process, including the mass balance (M), phase equilibrium

122

(E), summation (S), and enthalpy balance (H) at each stage:

123

Mass balance for each component at stage (j): (11)

Lj-1 xi,j-1 -Lj xi,j +Vj+1 yi,j+1 -Vj yi,j =0 124

Equilibrium for each component at stage (j): (12)

yi,j =K ·xi,j i,j

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The phase equilibrium is predicted using the RK EOS-based model defined in Equation (10).

126

Summation based on mole fractions at stage (j):

xi,j = yi,j =1 127

(13)

Enthalpy balance at stage (j):



Lj-1 HL, j-1 -Lj HL, j +Vj+1 HV, j+1 -Vj HV, j -Q=0

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(14)

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in which H# is the enthalpy of liquid flow (kJ/kmole); H$ is the enthalpy of gas flow (kJ/kmole);

129

L is the solvent molar flow rate (kmole/hr); K is the phase equilibrium constant (ratio); Q is the

130

cooling duty (kJ/hr); V is the gas molar flow rate (kmole/hr); x is the mole fraction in liquid

131

phase; y is the mole fraction in gas phase. An inside-out (I/O) algorithm is applied to solve the

132

coupled MESH equations, which employs two sets of thermodynamic property models: a simple,

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approximate, empirical set used frequently to converge inner-loop calculations, and a rigorous

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set used less often in the outer loop.31-32

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The pressure drop across an absorber is estimated as the product of pressure drop rate and

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absorber height. The generalized Sherwood/Leva/Eckert correlation is used to estimate the

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pressure drop rate.33-34 The absorber height is estimated based on the overall gas-phase mass

138

transfer coefficient of physical absorption, using mass transfer correlations developed by Onda et

139

al. for randomly packed vessels.35 Details of the pressure drop estimation are given in Section S3

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of the SI.

141 142 143

Flashing for Solvent Regeneration For solvent regeneration, the CO2-rich solvent stream exiting the absorber with a much high

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pressure can enter multiple flash drums with lower pressures, which is similar to the design

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adopted for Selexol-based CO2 capture.36 The solvent is regenerated and recirculated to the

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absorber, while the CO2 is released from the solvent in a series of flash drums and further

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compressed for transport to a storage site. The calculation of TP-flash based on the phase

148

equilibrium model discussed above is performed to assess the stripping process.

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Key Power Equipment

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There is no thermal energy required for the IL-based capture process. However, electric

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power is required for process and CO2 product streams compression and solvent pumping. In the

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pressure swing for solvent regeneration, some energy from high-pressure solvent streams can be

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recovered by hydraulic power turbines.37 Equipment power use calculations are detailed in

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Section S4 of the SI.

156 157

Engineering-Economic Analysis of CO2 Capture System This study adopts the costing method presented in the Electric Power Research Institute’s

158 159

Technical Assessment Guide.38-39 The performance models discussed above are integrated with

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engineering-economic models that estimate the capital cost, annual operating and maintenance

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(O&M) costs, and total annual levelized cost of an IL-based capture system. Table 1 summarizes

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the costing method.

163

The total capital requirement (TCR) includes the direct process facilities cost (PFC) plus a

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number of indirect cost categories. Major direct cost components are given in Table 1. The cost

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model developed by Chen for a Selexol-based CO2 capture system was adopted to estimate the

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major direct costs.37 In the model, the direct costs of absorber and flash drums per train are

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estimated as a function of stream mass flow rates and process pressure if applicable, while the

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direct costs of compressor, pump, and hydraulic power turbine per train are estimated as a

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function of equipment power use.37 These individual direct cost estimates were further updated

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mainly based on more recent cost studies by the U.S. National Energy Technology Laboratory.40-

171

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Details of the direct cost estimation are available in Section S5 of the SI.



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Indirect costs such as the general facilities cost (GFC) and engineering and home office fees

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are estimated empirically as a percentage of PFC. The process contingency depends on the status

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of technology development and is estimated as a percentage of PFC, whereas the project

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contingency depends on the project design level and is estimated as a percentage of the sum of

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PFC, GFC and process contingency.38 Other indirect cost categories considered include interest

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charges, royalty fees, preproduction cost, and inventory capital cost.

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Table 1 also summarizes variable and fixed O&M cost components. Variable O&M costs

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are estimated as the product of the quantity used times the unit price. Operating labor is

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estimated in terms of hourly labor rate, personnel per shift, and number of shifts. Total

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maintenance cost is estimated as a percentage of total plant cost, while administrative and

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support labor is estimated as a percentage of operating plus maintenance labor.

183

The LCOE of an overall plant or a system is then calculated as:45 LCOE=

TCR·FCF+FOM+VOM +HR·FC CF·8760·MW

(15)

184

Where LCOE is the levelized cost of electricity ($/MWh); TCR is the total capital requirement

185

($); CF is the capacity factor (%); FCF is the fixed charge factor (fraction/yr); FOM is the fixed

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O&M costs ($/yr); VOM is the non-fuel variable O&M costs ($/yr); HR is the net heat rate

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(MBtu/MWh); FC is the unit fuel cost ($/MBtu); and MW is the net power output (MW).

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The LCOE is also used to calculate the cost of CO2 avoided— a key measure of the overall cost

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of carbon capture and storage (CCS): 45 Cost of CO2 avoided ($/t CO2 )=

LCOE%%& − LCOERef ERRef − ER%%&


(16)

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where ER is the CO2 emission rate (t/MWh). The subscripts of “CCS” and “Ref” indicate an

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IGCC plant with CCS and a reference IGCC plant without CCS, respectively. A plant-level

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assessment will be conducted to estimate LCOE and ER for both plants with and without CCS.

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Some studies of CO2 capture processes also report the cost of CO2 separation ($/t CO2). This

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is typically based on a “stand-alone” analysis of a capture process in which the cost of energy to

195

operate the system is assumed to be purchased externally. This measure of cost is given by: Cost of CO* separation =

TCR·FCF+FOM+VOM CF·8760·mCO2

(17)

196 197 198

Integrated Environmental Control Model for Power Plant Assessments The Integrated Environmental Control Model (IECM) is a computer-modeling tool

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developed by Carnegie Mellon University for power plant assessments.44 The IECM provides

200

systematic estimates of the performance, emissions, and costs for a variety of fossil-fuel fired

201

power generation systems with and without CCS, including IGCC systems.36-37,40,46 To evaluate

202

the effects of CCS deployment on the overall plant performance and cost, the 2017 release of

203

IECM (Version 9.5) was employed for this study.44 The new performance and cost models

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developed for IL-based CO2 capture were integrated with the IECM framework and used to

205

conduct techno-economic assessments at both process and plant levels.

206 207

BASE CASE RESULTS

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Here we present results for base case assumptions using current IL properties. First we show

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results for the stand-alone IL-based capture process. Then we show results for a complete power

210

plant with pre-combustion capture. Following this, we look at potential process improvements

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that can enhance the viability of ILs for pre-combustion CO2 capture applications.



212 213 214

Phase Equilibrium of CO2 and H2 in [hmim][Tf2N] The calibrated VLE model based on the modified RK EOS can predict the CO2 and H2 phase

215

behavior under different temperatures and pressures. Figures 1(a) and 1(b) present the PT-x

216

phase diagram for CO2 and H2 in a binary system, respectively. CO2 has a much larger solubility

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than H2. When the interaction between CO2 and H2 in a tertiary system is considered, the EOS-

218

based VLE model needs to be further adjusted by incorporating the binary interaction parameters

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of CO2 and H2 to predict the simultaneous solubility of CO2 and H2 in [hmim][Tf2N]. The

220

predicted pressure of the tertiary system matches well with an experimental study conducted by

221

Kumelan et al.,47 resulting in an R2 value of 0.996.

222 223 224

Process-Level Evaluation An IL-based process based on the newly developed models was assumed to capture CO2

225

from the shifted syngas of an IGCC plant. A parametric analysis was further conducted to

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explore the effects of process design variables on process performance and cost for a “stand-

227

alone” capture system.

228 229 230

Base Case Studies Basha et al. proposed a process configuration for pre-combustion CO2 capture using ILs.17-18

231

Figure 2 shows a similar process configuration: CO2 is absorbed by [hmim][Tf2N] in a packed

232

column, then the CO2-rich solvent stream exiting the absorber is regenerated in pressure-swing

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flash drums arranged in series. Since the gas stream out of the high-pressure flash drum has a

234

high hydrogen concentration, it is recirculated to the absorber to avoid hydrogen losses. The CO2


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streams out of the medium- and low-pressure flash drums for solvent regeneration are combined,

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while the CO2-lean solvent stream out of the low-pressure flash drum is recirculated to the

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absorber. A hydraulic turbine is used to recover energy from pressure changes in the CO2-rich

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solvent stream.37 Table 2 summarizes the major technical and economic parameters and

239

assumptions for the pre-combustion IL-based capture system based on the syngas characteristics

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of an IGCC plant employing Shell gasifiers.37,44 Our process simulation results of stream mass

241

flow rates and compositions for a single train are presented in Figure 2 (the second train is

242

identical).

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The capture system is designed as an isothermal process (at about 29oC) with two trains for

244

90% CO2 capture. The designated absorption pressure is 2960 kPa. The CO2-rich solvent is

245

regenerated in three flash drums with pressures that decrease from 1000 kPa to 80 kPa. Figure 2

246

shows the mass flow rate and composition for each stream in each train. The H2-rich stream

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delivered as fuel to the gas turbines has a H2 purity of 94%, while the final CO2 product stream

248

has a purity of over 99% CO2. Figure 2 also shows that the hydrogen loss and [hmim][Tf2N] in

249

the CO2 product stream are negligible.

250

In each train, two compressors are used to increase the pressure of the recirculated gas

251

stream out of the HP flash drum from 1000 kPa to 2960 kPa, and of the CO2 product stream out

252

of the LP flash drum from 80 kPa to 500 kPa. The regenerated solvent is pumped back to the

253

absorber. The final CO2 product stream is then further compressed to a supercritical fluid phase

254

for transport and storage (T&S). Table 3 summarizes the equipment power use and recovered

255

energy. CO2 product compression is the largest power user.

256

Table 3 also presents the capital and O&M costs of the IL-based capture system, excluding

257

CO2 T&S costs. The major capital cost components are the absorbers, process compressors, and



258

product compressors, which account for 19%, 22%, and 36% of the TCR, respectively. Given the

259

assumptions in Table 2, capital cost accounts for more than a half the total annualized cost and

260

the separation cost for the IL-based system is about $17 per tonne of CO2 separated. We note,

261

however, that the separation cost for energy-intensive processes is strongly affected by the

262

assumed price of purchased electricity for a stand-alone process.49 In this example, doubling the

263

assumed price of electricity increases the separation cost to about $24 per tonne of CO2

264

separated.

265 266 267

Effects of Process Parameters A sensitivity analysis also was conducted to evaluate the effects of major process parameters

268

on the net process power requirement and the separation cost for CO2 capture. The parameters

269

considered include the CO2 capture efficiency, process operating temperature, and absorption

270

and stripping pressures. When a parameter was assessed, other parameters were held at their base

271

case values given in Table 2, unless otherwise noted.

272

The process operating temperature affects the solvent properties such as CO2 solubility and

273

solvent viscosity. When it increases, the solvent requirement increases for the given CO2

274

removal efficiency due to the decreased solubility. As shown in Figure 3(a), both the power use

275

and separation cost increase when the temperature increases. When the pressures of the HP and

276

MP flash drums increase from 800/400 kPa to 1050/525 kPa, the gas stream flow rate

277

recirculated from the HP flash drum to the absorber decreases and the required pressure ratio for

278

the process compressor decreases as well, which collectively lower the corresponding

279

compression power use. Figure 3(c) shows that both the power use and separation cost decrease

280

when the HP and MP stripping pressures increase.


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The absorption pressure affects the capture system performance. When the absorption

282

pressure varies from 2960 kPa to 5000 kPa, the mass flow rate from the HP flash drum to the

283

absorber increases by a factor of more than five and the CO2 concentration of the gas stream

284

recycled increases from 78.8% to 92.2%. These lead to increases in the compression power use

285

for the recycled gas stream, the total gas flow rate into the absorber, and the CO2-rich solvent

286

flow rate. Figure 3(c) shows that for the given three-stage stripping designs, elevating the

287

absorption pressure from 2960 kPa to 5000 kPa increases the system power use and separation

288

cost by 28% and 38%, respectively. This result indicates that when an absorption process occurs

289

at a much high pressure, stripping pressures need to be elevated accordingly.

290

The CO2 removal requirement largely determines the capture system size, power use, and

291

cost. Figure 3(d) shows that both the net normalized power use and separation cost decrease

292

when the CO2 removal efficiency increases from 50% to 95%. To achieve 95% CO2 capture, the

293

three stripping pressures have to be lowered from the base designs to 800 kPa, 400 kPa, and 60

294

kPa, respectively. Otherwise, the process simulation could not converge. As a result, both the

295

normalized power use and separation cost increase when the CO2 removal efficiency is elevated

296

from 90% to 95%. Figure 3(d) shows that the optimal removal efficiency is 90%.

297 298 299

Effects of Process and Project Contingencies The process and project contingencies represent indirect capital costs that are expected to

300

occur. The process contingency depends on the level of technology maturity and lies within the

301

range from 30% to 70% of PFC for a conceptual system with bench-scale data and 5% to 20%

302

for an operational full-size system. The project contingency is related to the level of project

303

design and typically lies within 30% to 50% of the sum of process capital, engineering and home



304

office fees, and process contingency for a simplified project design, and 10% to 20% for a

305

detailed project design.38 Our base case assumes a hypothetical mature plant, akin to analyses by

306

the U.S. Department of Energy’s National Energy Technology Laboratory and others.50 To

307

illustrate cost sensitivity, both the contingency factors are increased from their base values to

308

30%. As a result, the total capital requirement increases by 29% and the separation cost increases

309

by 17% in comparison to the base case.

310 311 312

Plant-Level Evaluation To analyze a complete IGCC plant with CO2 capture and storage (CCS), we use IECM 9.5

313

to model a plant employing Shell gasifiers. The plant design includes a two-stage water gas shift

314

(WGS) reactor and a combined cycle power plant employing GE 7FB gas turbines.37 Because the

315

cost of CO2 avoided offers a more rigorous metric than the separation cost calculated for a stand-

316

alone process with purchased energy,49 it is used to measure the overall cost of CCS. To estimate

317

the avoidance cost, the IECM was employed to calculate plant-level CO2 emission rates and

318

LCOE for Shell-IGCC plants with and without CCS. For the plant with CCS, the designated CO

319

to CO2 conversion efficiency is 95% for the WGS reactor, while the designated CO2 removal

320

efficiency is 90% for the CO2 capture system. The parametric values in default were adopted for

321

other power plant components. For the assumptions given in Table 2, the addition of the IL-

322

based CCS system decreases the net plant efficiency by 9.3 percentage points. The resulting cost

323

of CO2 avoided is $63 per ton of CO2 in constant 2011 dollars.

324

The IL-based technology is further compared to a Selexol-based two-stage system that also

325

employs physical absorption for CO2 capture.36-37,44 IECM 9.5 also was applied to evaluate the

326

performance and cost of an IGCC plant with Selexol-based CCS. The plant-level assessment


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327

results are provided in Table 4. The addition of the Selexol-based CCS for 90% CO2 capture also

328

leads to significant reductions in the net plant efficiency. The cost of CO2 avoided by the

329

Selexol-based CCS system is $65 per ton of CO2. So, the two capture systems have similar

330

effects on overall plant performance and cost of electricity generation. Thus, based on current

331

solvents, this analysis suggests that there is no compelling case for ILs replacing the more

332

established Selexol system.

333 334

ANALYSIS OF HYPOTHETICAL IMPROVED SOLVENTS

335

ILs have tunable properties for gas absorption. To help guide the development of novel materials

336

for CO2 capture, there is a need for quantitative targets for material properties based on a system-

337

level analysis. The model described in this paper provides that capability and is applied to

338

investigate the material properties needed to improve the IL technology's viability.

339

Given that CO2 solubility is the most important property affecting capture process

340

performance, we evaluate a hypothetical IL with improved solubility for CO2. As shown in

341

Figure S1 in the SI, the concentration of CO2 in equilibrium with the hypothetical IL is assumed

342

to be 10 percentage points more than that of [hmim][Tf2N] on an absolute basis for a given

343

temperature and pressure, while other properties are kept the same as those of [hmim][Tf2N]. No

344

deviations from the ideality also are assumed.51-52 A new phase equilibrium model was

345

developed based on the hypothetical solubility and incorporated in the process model for

346

performance and cost assessments.

347

Because of the improved solubility, the CO2-rich solvent stream exiting the absorber has a

348

higher CO2 concentration compared to the base case. If the three flash drums in Figure 2 are

349

employed for solvent regeneration, the gas stream out of the HP drum has a higher CO2 purity.



350

Recycling this CO2-rich gas stream back to the absorber would waste a lot of the separation

351

work. Thus, the HP flash drum and its associated compressor can be removed from the process

352

configuration. Figure S2 in the SI shows the simplified capture process employing the

353

hypothetical solvent. The design temperatures and pressures for the absorber and two flash

354

drums are held the same as the base case system.

355

Results of the process simulation show that the CO2 concentration of the solvent stream

356

leaving the absorber increases to 46% while the CO2 purity of the final product stream is 98%

357

(comparable to the base case). However, the CO2-lean solvent requirement decreases by about

358

70%, which significantly reduces the power requirement for solvent pumping. The resulting

359

system power use and separation cost are 34.6 MW and $12 per tonne of CO2 separated, which

360

are 35% and 29% less than the respective base case values. These results illustrate how an

361

increase of 10 percentage points in CO2 solubility can significantly improve the economic

362

viability of this technology, offering targets for future IL developments. Whether and when such

363

improvements can indeed be achieved, however, remains an open question.

364 365

DISCUSSION

366

This study has analyzed the use of [hmim][Tf2N] as a physical solvent for pre-combustion CO2

367

capture at IGCC power plants. The energy penalty of the IL-based system comes mainly from

368

the process and product compression and solvent pumping, while the major capital cost

369

components are the compressors and absorbers. Both the parasitic load and cost vary to some

370

extent with process design choices. The minimum cost of separation ($/ton CO2 separated from

371

the syngas stream) for the base case design was found to occur at a CO2 removal efficiency of

372

90%, indicating that a bypass design treating only a portion of the syngas would be cost-effective


Page 18 of 33

Page 19 of 33


373

for an IL-based system that required only partial CO2 capture (e.g. 50% removal). Comparisons

374

of plant-level performance and cost between IGCC plants employing [hmim][Tf2N]-based and

375

Selexol-based capture systems showed very similar results, implying that an IL-based CCS

376

system could be a viable alternative to current solvent processes.

377

Advances in multiple areas are needed to improve the viability of IL-based capture

378

technology. Improvements in the CO2 solubility of ILs are needed to simplify the capture system

379

configuration and lower the energy penalty and cost for CO2 capture. Novel compression

380

technology also is needed to further reduce power needs and system cost. In addition, since the

381

CO2 capture system accounts for less than 10% of an IGCC plant’s total capital requirement (see

382

Table 4), large-scale deployment of IGCC plants with CCS will require system-wide cost

383

reductions in other major plant components as well as in overall plant integration, configuration,

384

and design.

385

Two further caveats accompany this study. First, the models presented here do not consider

386

the potential influence on process performance of other minor gas species such as water vapor

387

and carbon monoxide. Second, to estimate the absorber pressure drop, this study employed mass

388

transfer and pressure drop rate correlations derived empirically for conventional solvents with

389

lower viscosity than ILs. Thus, there is a need for additional data in these areas to more

390

rigorously evaluate the effectiveness of ionic liquids as a cost-effective sorbent for pre-

391

combustion CO2 capture.

392 393

Supporting Information

394

Supporting Information includes additional text, tables, and figures on solvent and gases

395

properties, phase equilibrium model parameters and validation, absorber design, major



396

equipment power use or recovery, direct capital cost estimation, and modeling and analysis of a

397

hypothetical improved solvent. This material is available free of charge via the Internet at

398

http://pubs.acs.org.

399 400

ACKNOWLEDGEMENTS

401

This work was supported in part by Stanford University’s Global Climate and Energy Project.

402

All opinions, findings, conclusions and recommendations expressed in this paper are those of the

403

authors alone and do not reflect the views of any agencies.

404 405

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406

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542 543

Table 1. Costing Method and Nomenclature Category Cost Component Capital Cost Process facilities capital

Variable O&M Cost

Fixed O&M Cost

absorbers high-pressure flash drums middle-pressure flash drums low-pressure flash drums solvent pumps high-pressure stream compressors low-pressure stream compressors CO2 product compressors hydraulic power recovery turbines sump tanks heat exchangers General facilities capital Engineering and home office fees Project contingency Process contingency Interest charges Royalty fees Preproduction cost Inventory capital Solvent makeup Electricity CO2 transport and storage when applicable Operating labor Maintenance labor Maintenance material Administrative and support labor

544 545 546 547 548 549 550



551 552 553

554 555

Table 2. Major Technical and Economic Parameters and Assumptions for IL-based Pre-combustion CO2 Capture Technical Parameter Unit Value Economic Parameter Unit Number of trains # 2 Discount rate % Capacity factor % 75 Fixed charge factor fraction Total syngas flow rate kmol/hr 27900 Construction time yr Syngas pressure kPa 2960 General facilities capital % PFC o Syngas temperature C 29.4 Engr. and overhead fees % PFC CO2 concentration in syngas % 37 Process contingency % PFC CO2 removal efficiency % 90 Project contingency % (PFC+Overhead +Process Conting.) Absorber pressure kPa 2960 Royalty fees %PFC High/middle/low flashing kPa 1000/500/80 Misc. capital cost %TPIb pressures o Flashing temperature C 29.4 Inventory capital %TPCb Solvent loss ratea kg IL/tonne CO2 0.23 Total maintenance cost %TPCb Compressor efficiency % 80 Labor fee $/hr Pump efficiency % 75 Makeup solvent price $/tonne Recovery turbine efficiency % 75 Reclaimer waste disposal $/tonne CO2 product pressure MPa 15.3 CO2 transport and storage $/tonne Electricity price $/MWh a This is derived based on the study by Maginn et al.48 b TPI is the total plant investment. TPC is the total plant cost.

556 557 558


Page 28 of 33

Value 7.09 0.113 3 15 10 10 15 0.5 2 0.5 5.0 34.65 10000 260 10 50

Page 29 of 33

559 560 561


Table 3. Performance and Costs of IL-based CO2 Capture System Category System power use (MW) Solvent pumping Process compression CO2 product compression Hydraulic turbine recovery power Total net system power use Process facilities capital (2011 M$) Absorbers Flash drums Solvent pumps Heat exchangers Slump tanks Process compressors CO2 product compressors Power recovery turbines Total capital requirement (2011 M$) Annualized capital cost (2011 M$/yr) Fixed O&M cost (2011 M$/yr) Variable O&M cost, excluding CO2 T&S (2011 M$/yr) CO2 separation cost, excluding CO2 T&S (2011 $/tonne CO2)

562 563 564


Value 18.1 14.0 30.4 9.4 53.1 99.3 19.1 7.9 4.9 2.4 3.9 22.1 35.5 3.5 168.3 19.0 9.4 17.4 17.1


565 566 567

568 569 570

Table 4. Comparisons of Performance and Costs of IGCC plants with and without CCS Parameter IGCC IGCC IGCC w/o CCS w/ Selexolw/ IL-CCS CCS Gross power output (MW) 681 656 656 Net power output (MW) 595 533 534 CO2 capture system Power use (MW) 54 53 TCR(2011$/kWnet) 365 315 Net plant efficiency (HHV, %) 44.3 34.9 35.0 Plant CO2 emission rate (kg/kWh) 0.710 0.131 0.130 Plant TCR (2011$/kWnet) 3211 4381 4323 a Plant LCOE (2011$/MWh) 86.6 124.4 123.2 Added LCOE for CCSb (2011$/MWh) 37.8 36.6 (relative %) 44 42 b Cost of CO2 avoided (2011$/tonne CO2) 65 63 a The assumed coal price is $1.62/GJ. b This also includes CO2 transport and storage costs.

571 572 573 574 575 576 577 578 579 580 581


Page 30 of 33

Page 31 of 33


582 583 584 585 14000

14000 298.15 K 313.15 K 373.15 K 413.15 K

10000

10000

8000 6000

8000 6000

4000

4000

2000

2000 0

0 0%

586 587 588 589 590

293.15 K 313.15 K 333.15 K 353.15 K

12000

Pressure (kPa)

Pressure (kPa)

12000

20% 40% CO2 mole% in [hmim][Tf2N]

60%

0%

1%

2%

3%

4%

H2 mole% in [hmim][Tf2N]

(a) (b) Figure 1. PT-x phase diagram for a binary system (a) CO2 in [hmim][Tf2N]; (b) H2 in [hmim][Tf2N]

591 592 593 594 595 596 597 598 599 600 601 602 603 604 605


5%

6%


606 607 608 609

610 611

Figure 2. Pre-combustion CO2 capture system using [hmim][Tf2N]

612 613


Page 32 of 33

Page 33 of 33


614 615 616

15 5.0 10 Net Power Use Separation Cost 2.5 35 Operating Temperature (oC)

5.0 10 Net Power Use Separation Cost

40

Absorption Pressure (kPa)

(a)

(b)

20 7.5 15 5.0 10 Net Power Use Separation Cost 2.5 800

850

900

950

High Flash Pressure (kPa)

5 1000

10.0 Net Power Use (kWh/kmol CO2)

25 Cost of CO2 Separation ($/t CO2)

Net Power Use (kWh/kmol CO2)

10.0

619 620 621 622 623

15

25

20 7.5 15 5.0 Net Power Use Separation Cost

10

2.5 5 50% 55% 60% 65% 70% 75% 80% 85% 90% 95% CO2 Removal Efficiency

(c) (d) Figure 3. Sensitivity of capture process power use and separation cost to process design (a) operating temperature; (b) absorption pressure; (c) regeneration pressure; (d) CO2 removal efficiency


Cost of CO2 Separation ($/t CO2)

30

20 7.5

2.5 5 2950 3250 3550 3850 4150 4450 4750 5050 5350

5

617 618

25 Cost of CO2 Separation ($/t CO2)

20 7.5

10.0 Net Power Use (kWh/kmol CO2)

25

Cost of CO2 Separation ($/t CO2)

Net Power Use (kWh/kmol CO2)

10.0

Systems Analysis of Physical Absorption of CO2 in ... - ACS Publications

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