Thermodynamic Modeling and Assessment of Ionic Liquid-Based CO2

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Thermodynamic Modeling and Assessment of Ionic Liquid-based CO2 Capture Processes Ying Huang, Xiangping Zhang, Xin Zhang, Haifeng Dong, and Suojiang Zhang Ind. Eng. Chem. Res., Just Accepted Manuscript • Publication Date (Web): 17 Jun 2014 Downloaded from http://pubs.acs.org on June 18, 2014

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Industrial & Engineering Chemistry Research

Thermodynamic Modeling and Assessment of Ionic Liquid-based CO2 Capture Processes Ying Huang1,2, Xiangping Zhang1,*, Xin Zhang1,2, Haifeng Dong1,2, Suojiang Zhang1,* 1

Beijing Key Laboratory of Ionic Liquids Clean Process, State Key Laboratory of Multiphase

Complex Systems, Key Laboratory of Green Process and Engineering, Institute of Process Engineering, Chinese Academy of Sciences, Beijing 100190, China 2

School of Chemistry and Chemical Engineering, University of Chinese Academy of Sciences,

Beijing 100049, China ABSTRACT Ionic liquid (IL)-amine hybrid solvents has been experimentally proved to be effective for CO2 capture. This paper provided rigorous thermodynamic models, process simulation and cost estimation of a potential design of IL-based CO2 capture processes. Three ILs ([Bmim][BF4], [Bmim][DCA] and [Bpy][BF4]) were investigated to blend with MEA aqueous solution. The physicochemical properties of the ILs were predicted by several temperature-dependent correlations. Phase equilibria were modeled based on Henry’s law and NRTL equation, and the calculated values were in good agreement with the experimental data. The simulation results show that the [Bpy][BF4]-MEA process can save about 15% regeneration heat duty compared to the conventional MEA process, which is attributed to the reduction of sensible and latent heat. Moreover, a modified [Bpy][BF4]-MEA process via adding inter-cooling and lean vapor recompression presents 12% and 13.5% reduction in overall equivalent energy penalty and capture cost compared with the conventional MEA process, respectively.

1

ACS Paragon Plus Environment


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1. INTRODUCTION Carbon capture and storage (CCS) is recognized as an option to reduce the greenhouse gas emissions1, including three typical technologies, i.e., pre-combustion capture, post-combustion capture and the oxyfuel process2. Among which post combustion capture via amine absorption such as monoethanolamine (MEA) has been commercially used for decades3-5 with several disadvantages, including the loss of solvent, large stripping energy consumption, unfavourable degradation6 and corrosion problems. Thus it is highly required to develop new solvents and processes. In recent years, ionic liquids (ILs) have been paid more attentions7-12 because of their nonvolatility, thermal stability and tunable chemistry, which are perceived to be helpful in reducing energy consumption and environmental pollution. Conventional ILs such as 1-butyl-3-methylimidazolium

tetrafluoroborate

([Bmim][BF4])13

and

1-butyl-3-methylimidazolium bis((trifluoromethyl)sulfonyl)imide ([Bmim][Tf2N])14 present relatively well CO2 solubilities under high pressure, thus they are not suitable for CO2 capture from flue gas under atmospheric pressure. Task-specific ILs with amine groups can chemically absorb up to 1 mole CO2 per mole IL15. However, it may be difficult to employ the task-specific ILs in industrial applications because of complex synthesis and purification steps, high cost16 and increased viscosity after the absorption reaction17. Alternatively, IL-amine hybrid solvents seem to be effective for CO2 capture, such as IL-MEA18 and IL-MDEA19-21 aqueous solutions. They are prior to current amine absorption in terms of reducing solvent volatility, energy consumption and the corrosion rate22, 23. Recently, Yang et al.24 conducted a CO2 capture experiment in an absorption−desorption loop system using an IL-MEA hybrid 2


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solvent (30 wt % MEA + 40 wt % [Bmim][BF4] + 30 wt % H2O). They observed that the MEA loss for the mixed solution was much lower than that for the MEA solution. Besides, no ionic liquid loss was detected. These results proved the possibility and feasibility of IL-amine solvents for CO2 capture, including the CO2 removal efficiency and the solvents stability. LaFrate et al.

25

detected the degradation of two IL-MEA solvents. The results revealed that

[Emim][OTf]

(1-ethyl-3-methylimidazolium

trifluoromethanesulfonate)

make

the

degradation of the solvent faster, which was induced by amine alkylation by the anion([OTf]). Moreover, Janiczek et al.26 set up a technical-scale-plant for CO2 absorption from natural gas streams utilizing an imidazolium-based IL, which indicated the potential of industrial application of ILs. Process simulation and assessment is essential for a new process to be developed in industry. However, such research related to IL-based processes was rarely reported probably due to the shortage of rigorous thermodynamic models for complex IL-containing systems27. Shiflett et al.28 simulated a capture process using pure [Bmim][Ac] and claimed a 16% reduction in energy consumption compared to the commercial MEA process. Basha et al.29 developed a conceptual CO2 capture process applying three flash drums as regenerators. They considered [Hmim][Tf2N] could be used as a physical solvent for high efficient CO2 capture from warm shifted fuel gas streams with relatively high CO2 content (23.87 mol.%). The above pioneering works provide ideas and methods for the simulation of CO2 capture process using pure IL, while they may be not available for flue gas with lower CO2 concentration. Thus it is imperative to simulate and assess the capture processes using IL-amine hybrid solvents. Besides developing new solvents, efforts were also concentrated on the process 3



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modification to improve the process performance. Numerous modifications can be made on the basis of a standard process such as split flow, rich split, vapor recompression, and absorber inter-stage cooling30. Karimi et al.31 found that vapor recompression was the best configuration with the lowest increase in total capture cost and the plant complexity. In general, in order to achieve the energy benefits, the capital cost and the complexity of the plant will probably increase32. So the satisfactory process should the one with the best balance between additional cost and decreased energy penalty33. Taken together, it is meaningful to employ an IL-amine hybrid solvent in a modified process configuration to develop an efficient capture process. The aim of this work is to provide thermodynamic models and system assessment method for new ionic liquid-based CO2 capture processes. Ionic liquids were created as new (pseudo) components in Aspen Plus (version 8.4), requiring the import of the basic properties and parameters such as critical properties and heat capacities, and the parameters of thermodynamic models which were regressed via experimental data from literature and our study. The rigorous thermodynamic models of the MEA-H2O-IL systems were verified to ensure the reliabilities of the simulation results. Based on the results of process simulation, the energetic analysis and cost estimation of IL-based processes were performed, taking the conventional MEA-based process as the reference case. 2. SELECTION OF ILS The anion of an ionic liquid has a greater influence on the CO2 solubility than the cation34. In this work, three hydrophilic ionic liquids ([Bmim][BF4] , [Bmim][DCA] and [Bpy][BF4]) were selected as component of the IL-amine hybrid solvents based on the following reasons. [Bmim][BF4] has been studied frequently for mixing with MEA. Taib et al.18 measured CO2 4


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solubilities in [Bmim][BF4]+MEA solution at pressure from 100 to 1600 kPa and observed that the solubilities were higher than conventional MEA solution. Based on a kinetic experiment, Lu et al.35 found that the values of the enhancement factor and the second-order reaction rate constant for the [Bmim][BF4]+MEA solution were higher than those in pure MEA aqueous. [Bmim][DCA] has very low viscosity (16.8mPa·s at 40℃ 36) which is favorable for the absorption process. Experiments results including the CO2 solubilities, density and viscosity of the [Bmim][DCA]-MEA hybrid solvent have been reported in our previous work37. [Bpy][BF4] is perceived to be more potential for large scale utility with lower cost38, toxicity39 and higher biodegradability40 than imidazolium-based ILs. 3. METHODOLOGY A four-step procedure of this study was conducted: 1) The critical properties of the ILs were calculated based on the FC-CS method41 proposed in our previous work. The temperature-dependent properties, including densities, viscosities, surface tensions, heat capacities and thermal conductivities were predicted and compared with the experimental data from literature. 2) Gas-liquid equilibrium (GLE) of CO2-IL was correlated using the experimental data from literature. Vapor-liquid equilibria (VLE) of the IL-MEA-H2O systems were measured for the thermodynamic study of the whole absorption system. 3) Four IL-based CO2 capture processes and the reference MEA process were simulated, including an IL-based process with configuration modification. 4) The MEA and IL-based processes were assessed and compared on the aspects of energy consumption and capture cost. 3.1 Experimental section 3.1.1 Materials.

All the above ILs were supplied by Linzhou Keneng Materials 5



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Technology Co., Ltd. China and dried under vacuum for 48 hours at 333 K before used. The water contents of the ILs after drying were ensured ＜200 ppm by Karl-Fisher titration. The residual chloride in the ILs were ＜ 0.002mol·L-1, which were measured by the ion chromatography (Dionex ICS-900). MEA (purity>99.8%) was purchased from Beijng Chemical Works and used without further purification. The deionized water was supplied by the local laboratory. 3.1.2 Apparatus and Operation.

The experimental apparatus was based on the

quasi-static ebulliometry method which was the similar to some researchers 42,43-45, as shown in Figure 1. It was composed by two ebulliometers (a working ebulliometer and a reference ebulliometer) connected to a buffering vessel, two condensers to minimize the composition variation, two temperature control and measurement systems to make the liquid boiling and record the temperature, two magnetic stirrers, and a pressure transducer system. The total variation of the liquid phase composition caused by the holdup amount of volatile component in the condenser was within 0.2%42. First, the hybrid solvents were obtained by mixing definite weight of the corresponding components weighted by an electronic analytical balance (BS124S, Sartorius Scientific Instrument Co. Ltd., China) with precision of ± 0.0001 g. Second, the sample was added into the working ebulliometer and the deionized water was added into the reference ebulliometer. Third, the system was evacuated to an absolute pressure of approximately 7 kPa, and then the solution was heated and stirred. When VLE was reached, the temperature of two ebulliometers was unchanged and recorded. A series of VLE data were obtained by increasing the system pressure. 6


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3.2 Prediction of Physicochemical Properties The physicochemical properties, phase equilibrium and chemical reaction of the MEA-H2O-CO2 system were modeled based on the electrolyte nonrandom two-liquid model. This property method has been commonly adopted for simulating the MEA scrubbing process27, 46-48

. For ionic liquids, the property prediction and phase equilibrium modeling were described

in detail in the subsequent sections. 3.2.1 Scalar Properties.

Critical properties are used in many corresponding states

correlations for volumetric, thermodynamic, and transport properties of gases and liquids49. They are also essential for a process simulation. Critical properties of the ILs were obtained from the FC-CS method reported in our previous work41. Other properties were obtained from literature50,51,52, 53,54. 3.2.2 Temperature-dependent Properties.

The temperature-dependent properties such as

heat capacity, density and viscosity were calculated by the equations (Eq.1-Eq.6) listed in Table 1. Ionic liquid is well-known as non-volatile liquid whose vapor pressure is difficult to be observed. Liquid heat capacity is a basic thermodynamic property used for specifying the amount of heat required to change the temperature of a liquid by a given amount, which is important in analyzing energy demand of the whole process. All the equation coefficients were obtained by minimizing the objective function (Eq. 7) based on available experimental data. The properties were calculated by the FC-CS method41 if there was no available experimental data. NP

O.F . = ∑ ( X i cal − X i exp )

2

(7)

i =1

The average absolute relative deviation (AARD) was defined as follows: 7



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NP

AARD(%) = 100 × ∑ X i cal / X i exp − 1.0 / N P

(8)

i =1

where X represents the temperature-dependent property, subscript ‘i’ denotes different investigated IL, Np denotes the total number of data points, and the superscripts ‘cal’ and ‘exp’ are the calculated values and the experimental data, respectively. 3.3 Phase Equilibria Modeling The phase equilibrium and kinetic reactions among the MEA-H2O-CO2 system were modeled using a so-called KEMEA thermodynamic package, which is available within the Aspen Plus modeling system based on the electrolyte nonrandom two-liquid (ELECNRTL) model. This property method has been commonly adopted for simulating the MEA scrubbing CO2 process by many researchers46-48. IL was created as a new component whose interaction parameters with other components were required and determined as the following described. Solubilities of N2 and O2 in ILs were very low55 at the magnitude of 10-4. Thus N2 and O2 were assumed as no solubility in ILs in the simulation. Moreover, their solubilities in MEA and H2O were also very low which have been included in the ‘KEMEA’ property package. 3.3.1 GLE of the CO2-IL Binary System.

The phase equilibrium relationship for

dissolved gases was modeled as following56:

ϕiV yi p = xiγ i* H iA

(9)

γ i* = (γ i / γ i∞ )

(10)

where γ i∞ is the infinite dilution activity coefficient of component i in the mixture. The Henry’s law constant of gas component i was defined as

fi ϕiV yi p H iA = lim = lim xi → 0 x xi →0 xi i

(11)

8


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where p, yi, and xi are total pressure, mole fraction of component i in vapor phase and liquid phase, respectively. ϕ iV is the fugacity coefficient in vapor phase. HiA is the Henry constants in Pa of component i in solvent A. The temperature dependence of the Henry constants was described by the following equation56:

ln H iA = aiA + biA / T + ciA ln T + d iAT

(12)

where aiA, biA, ciA and diA are the equation parameters. Due to the negligible vapor pressure of ILs, the gaseous phase was assumed to be pure CO2. The fugacity coefficient in vapor phase ϕ iV can be obtained by Redlich–Kwong (R–K) equation of state57, given in the following form.

ln ϕ = Z − 1 − ln( Z − bp / RT ) − (

a / R 2T 2.5 ) ln(1 + bp / ZRT ) b / RT

Z = pV / RT

p=

RT a − 0.5 V − b T V (V + b) a = 0.42748

(13) (14) (15)

R 2Tc2.5 pc

(16)

RTc pc

(17)

b = 0.08664

where ϕ is the fugacity coefficient, Z is the compressibility factor, R is the gas constant, a and b are the EOS constants, Pc and Tc are the critical pressure and critical temperature, respectively, and V is the molar volume. * Based on the calculated H CO2 and γ CO , the activity coefficient γ CO2 in the mixture were 2

modeled by the NRTL model58: 9



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δ

lnγ i =

∑ xτ j =1

j

ji

G ji

δ

∑xG k =1

k

ki

 δ  x j Gij + ∑ δ j =1   ∑ xk Gkj  k =1

  τ ij −   

δ

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 Gmj     xk Gkj   

∑xτ

m =1

δ

∑ k =1

m mj

(18)

Gij = exp( −α ijτ ij ) ; τ ij = ( gij − g jj ) / RT = aij + bij / T ; α ij = α ji (i ≠ j ) where δ is the number of components; x represents the mole fraction; R is the gas constant; T is the absolute temperature; G ij is a dimensionless interaction parameter depending on the energy interaction parameter (gij) and the nonrandomness factor (αij). α ij was fixed at 0.3 used in most polar systems. The parameters of Henry constants and NRTL binary interaction parameters between CO2 and IL were generated based on the above models using experimental solubility data from literature.

3.3.2 VLE of the MEA-H2O-IL System.

The following equation (modified Raoult’s law)

was used for VLE calculation at the pressures less than 2 bars.

yi p = xiγ i pis

(19)

where p and pis are total pressure of the system and vapor pressure of the pure component i, respectively. yi and xi represent mole fraction of component i in the vapor phase and liquid phase, respectively. Due to the vapor mole fraction of IL was assumed as zero, Eq. 19 can be simplified to Eq. 20 and Eq. 21 for binary system and ternary system, respectively.

p = x1γ 1 p1s

(20)

p = x1γ 1 p1s + x3γ 3 p3s

(21)

where subscript ‘1’ denotes H2O, ‘3’ denotes MEA.

10


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Vapor pressure of pure H2O and MEA can be calculated by the Antoine Equation. Based on the vapor pressure of liquid mixture of different composition, the phase equilibrium criterion (Eq. 20 and Eq.21) combining with the NRTL model (Eq. 18) were used to calculate binary parameters of IL-H2O and IL-MEA systems.

3.4 Process Assessment 3.4.1 Process simulation.

The MEA and IL-based processes were all simulated in Aspen

Plus, among which the columns were modeled by RadFrac (an equilibrium stage model). The flue gas specification from the Shanxi Coke Plant in China was as following, temperature: 200 ℃ ; pressure: 1.01bar; mole flow rate: 24000kmol/h; mole composition: 6.37% CO2, 69.46% N2, 3.66% O2, 20.51% H2O. The plant produces methanol from CH4 reforming, thus the captured CO2 can be utilized as an additive to adjust the H/C ratio to an appropriate value for methanol production, which was perceived to be beneficial both in economy and environment, similar to the idea of Taghdisian et al.59. The flue gas has been desulphurized before it is sent to the CO2 capture unit. Figure 2 illustrates the conventional capture process, including the following main parts: (1) Gas pretreatment. The flue gas is pressurized to 1.1 bar using a blower and cooled down to approximately 308K after passing the washing column. (2) CO2 removal step. The cooled flue gas then contacts the solvent in a packed counter-current absorber. The purified gas is vented to the atmosphere after a MEA recovery section at the top of the absorber. (3) Heat exchanger section. The rich solvent from the absorber bottom is preheated by the lean solvent from the stripper bottom via a lean/rich heat exchanger and then pumped into the stripper. (4) Solvent regeneration step. The preheated rich solvent is regenerated in the stripper to release CO2. CO2 11



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and steam are cooled in a condensate reflux section at the top of the column, with condensed water being recycled to the top of the stripping column. The lean solvent from the bottom is cooled to 313K by passing the lean/rich heat exchanger and lean solvent cooler before recycling to the absorber. (5) Compression step. CO2 is dried and compressed to 2.5 MPa to enter into the methanol production plant. A modified capture process adding an absorber inter-cooling (AIC) and a lean vapor recompression (LVR) is depicted in Figure 3. Such modifications have been proved to be efficient in reducing energy consumption both by experiment and simulation31,60. The AIC section draws a side stream from the absorber and feeds it to the intercooler to cool down to a specific temperature. The cooled liquid is returned to the column at a location just beneath where the liquid is taken. Lower temperature of the solvent may be benefit in increasing absorption capacity thus decrease the solvent demand. In LVR scheme, semi-lean solvent is extracted from the stripper and flashed. The flashed lean vapor is recompressed to the stripper to substitute part of steam. The stage numbers of absorber and stripper were determined by sensitivity analysis depicted in Supporting Information. Murphree efficiency of the absorber was set as 25% (typically Murphree efficiency can be set as 10-30%)61. The flow rate of the inlet washing water of the MEA recovery section was varied to maintain the water balance of the system, and to lower the MEA slip to an acceptable value of 1 ± 0.05 ppmV. The equipment specifications of the modified process were set referring to literature31,62 as shown in Table S11 of Supporting Information.

12


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3.4.2 Energetic analysis.

The specific regeneration heat duty is often an indicator of

process performance, but it is not sufficient to rank different capture process. Thus the overall energy consumption was expressed by the equivalent energy penalty considering both electrical power and thermal energy. As adopted by other researchers31,

63

, the turbine

efficiency to produce electrical power from steam was assumed to be 75%, the temperature of steam was assumed to be 10 K higher than the reboiler temperature (Treb) and that steam condensed at 313 K in the turbine.

 313  Weq = 0.75Qreb 1 −  + Wcomp + Wpump  Treb + 10  3.4.3 Cost estimation methodology.

(22)

There are several metrics to measure the cost of CCS

system, including the cost of CO2 avoided, cost of CO2 captured, cost of CO2 abated (or reduced), and the increased cost of electricity64. In this study, cost of CO2 captured was selected. The total capture cost (TCC) was calculated from the sum of annual capital cost (ACC) and total operating cost (TOC) based on Table 2 which were common used in the literature33, 65. The purchased equipment cost (PEC) for columns, heat exchangers and pumps were obtained according to the NETL report66. Cost for those equipment out of the capacity scale given in the report were computed using the six tenths rule67 as depicted in Eq. 23. All the equipment costs were updated to the year 2008 with the Chemical Engineering Cost Index. Then TCI was computed on the basis of Table S1 in Supporting Information, referring to Abu-Zahra et al.65 and Schach et al.33

S  CB = C A  B   SA 

0.6

(23)

13



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where CB is the cost of equipment having size SB, CA is the known cost of equipment having corresponding size SA (same units as SB), and SB/SA is the ratio known as the size factor, dimensionless. Annual capital cost (ACC) was computed by Eq. 24, which converted TCI to a constant series of payments for every year of the project life.

ACC =

TCI ((1 + ir ) − 1) / ir (1 + ir )n n

(24)

Where ir denotes the interest rate and n denotes project life time. The total operating cost was divided into two parts. One was variable operating cost from MEA make up and utility consumption such as steam, cooling water and electricity etc. Another was fixed operating cost, mainly composed by operating labor, plant overheads, research and development (R&D) cost etc. Detailed computation basis was shown in Table S2 in Supporting Information. The utility price and calculation percentage for fixed operating cost (FOC) were taken from Mores et al.68, Peters et al.69 and Karimi et al.31. The price of MEA was 1250 US$/t68, while The price of IL was estimated at 6600$/t under the assumption of industrial production in the future based on the estimation by Linzhou Keneng Materials Technology Co., Ltd. China.

4. RESULTS AND DISCUSSION 4.1Validation of Thermodynamic Property Models Scalar properties from literature were listed in Table S4 of Supporting Information. The equation coefficients for vapor pressures, heat capacities, densities, viscosities, surface tensions and thermal conductivities were listed in Table S3 of Supporting Information. As shown in Figure 4, the predicted value for these properties display high accordance with the 14


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experimental data with the AARD less than 0.5%.

4.2 Phase Equilibrium Calculation Results Calculated Henry constants were listed in Table 3. NRTL binary interaction parameters of CO2, H2O and MEA with different IL were obtained as shown in Table 4. As shown in Figure 5, the calculated CO2 solubilities agree well with the experimental solubilities with an AARD of 0.9%. Experimental and calculated vapor pressure data can be found in Tables S5−S10 of Supporting Information with an overall AARD of 0.8%. The good accordance between experimental and calculated values in Figures 6−8 validate the applicability of the binary interaction parameters for all systems investigated. All the experimental results indicate that ILs decrease the vapor pressures of solvents due to diluting effect and affinity to H2O. So IL can improve the boiling temperature of the solvent thus increase the reboiler temperature of the IL-based processes.

4.3 Comparison of the MEA and IL-based Processes The following processes were simulated: (1) reference case of 30wt % conventional MEA process; (2) three conventional process configuration using IL-MEA solvents ([Bmim][BF4]-MEA, [Bmim][DCA]-MEA and [Bpy][BF4]-MEA). After that the IL-based solvent with lowest regeneration heat duty was used in the modified process. The main simulation results to assess a CO2 capture process were summarized in Table 5. The following discussion compared the [Bpy][BF4]-based processes with the MEA process in term of lean loading, solvent specific demand, cooling water demand, energy consumption and capture cost.

4.3.1 Lean loading (αlean).

Lean loadings of the MEA process and the IL-based 15



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processes were studied to achieve the minimal regeneration heat duty. At the same time, the lean solvent circulation rate was varied to achieve the same CO2 removal capacity. As shown in Figure 9, regeneration heats of different processes present the similar variation with the increase of αlean. The minimum regeneration heat duty can be reached at αlean of 0.25 for the MEA process, 0.22 for the [Bmim][DCA]-MEA process, 0.2 for the [Bmim][BF4]-MEA and the [Bpy][BF4]-MEA process, which were defined as the optimum lean solvent loadings. The reboiler heat duty of 3.7 GJth/t CO2 for the MEA process from this work is close to the value reported by many researches62, 70, 71, which indicates the simulation measures and results are valid. The [Bpy][BF4]-MEA hybrid solvent was used in the modified IL-based process due to the lowest heat duty. The main reasons for such changing trend are as follows. At low values of αlean, the latent energy of generating steam is dominant in the overall heat consumption. With the increase of

αlean, the latent heat decreases till constant, while the sensible heat increases and begins to play a key role in the total regeneration heat consumption due to higher solvent flow rate required for a fixed CO2 removal ratio. According to the experimental results by Kim et al.72, heat of absorption almost keeps constant within the αlean range of 0.16 to 0.5. Thus the minimum heat duty can be reached at the coordinate of sensible heat and latent heat.

4.3.2 Solvent Specific Demand.

Higher value of ∆CO2 loading results in a significant

decrease of solvent specific demand. The solvent circulation rates of the [Bpy][BF4]-MEA process and the modified [Bpy][BF4]-MEA decrease by 22% and 28% compared to the MEA process, which result in lower regeneration sensible heat. Besides, the powers of solvent pumps reduce thus saving a small amount of electrical power. 16


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4.3.3 Cooling Water Demand.

The cooling water demand of the [Bpy][BF4]-MEA

process is less than that of the MEA process. The reduction of circulating solvent flow rate contributes to the reduction of cooling water demand of MEA recovery section and lean cooler. The reduction of lean solvent temperature out of lean/rich exchanger also cut down the cooling water demand of lean cooler. The reduction of vaporized water causes the decrease of cooling water demand of stripper condenser. Ultimately, the cooling water specific demand of the [Bpy][BF4]-MEA process decreases by about 20% compared to the MEA process. Although the modified [Bpy][BF4]-MEA process require more cooling water than the [Bpy][BF4]-MEA process due to the additional inter-cooler, the amount is still less than the reference MEA process.

4.3.4 Energy Consumption.

Regeneration heat duty and total equivalent energy penalty

of the MEA and IL-based processes are described in Figure 10. It is observed that the IL-based processes require less regeneration heat duty than the MEA process, which is attributed to the decrease of sensible and latent heat during regeneration. The sensible heat is decreased because of two factors: (1) The heat capacity of the IL-MEA hybrid solvent is lower than that of 30% MEA solution because the heat capacity of IL is obviously lower than water; (2) The solvent flow rate of the IL-MEA hybrid solvent is lower than MEA solution, which is resulted from higher value of ∆CO2 loading determined by the lean loading optimization. The latent heat decreases due to the reduction of the amount of water vaporized, which is achieved by using IL to reduce vapor pressure of the solvent or increase the reboiler temperature. The modified [Bpy][BF4]-MEA process cut down thermal energy by 31% than the conventional MEA process, because LVR section can provide additional stripping steam to reduce the latent heat, 17



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and AIC section decrease the solvent demand to reduce the sensible heat. Besides the thermal energy consumption, the electrical power required for the [Bpy][BF4]-MEA process is a little lower than the MEA process because of lower pump power. However, power consumption of the modified [Bpy][BF4]-MEA process is higher than the MEA process because of the additional compressor. All in all, the [Bpy][BF4]-MEA process and the modified [Bpy][BF4]-MEA process saved the overall 5.8% and 12% of the equivalent energy compared to the reference MEA process, which indicate the IL-based processes are energy-saving compared to the MEA process.

4.3.5 Cost analysis.

The differences of calculation criterion, process configuration,

disposal capacity and CO2 concentration should be considered when comparing the capture cost with other researches. These will lead to different capture cost results, for instance, 55$ by Hassan et al.73, 74€ by Raynal et al.71 and 112$ by Mores et al68. The capture cost of 70$ per ton CO2 by this work is reasonably among the above values. The total capital investment (TCI) is depicted in Figure 11. It is observed that the fixed capital investment (reflecting equipment cost) of the [Bpy][BF4]-MEA process is slightly lower than the MEA process, which is resulted from the reduced cost of pumps and heat exchangers. Even so, TCI of the [Bpy][BF4]-MEA process is a little higher than the MEA process due to increased solvent investment. For the modified [Bpy][BF4]-MEA process, each item of the capital investment is the highest among the three processes, which can be attributed to solvent investment and high purchased equipment cost due to additional flash tank and compressor. Figure 12 presents the distribution of total capture cost (TCC) per ton CO2 captured. It can be firstly observed that total operating cost occupy more than 70% of the total CO2 capture cost. 18


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The main sinks for electrical power cost are the compressor, flue gas inlet blower and pumps. Steam is consumed as a heat source during regeneration. The above costs are the dominating contributors (above 80%) to the total operational cost (TOC). This illustrates the importance of reducing regeneration energy. With comparison to the reference MEA process, the [Bpy][BF4]-MEA process and the modified [Bpy][BF4]-MEA process offer steam cost reduction of 21.8% and 36.1%, respectively. However, the total capital investments (TCI) of IL-based processes are larger than MEA process mainly due to the high cost of IL. Even so, the overall capture cost of IL-based and modified IL-based process are reduced by 11% and 13.5%, respectively compared to the MEA process. This can be illustrated by the low percentage of solvent cost in the total capture cost (1.3% and 1.2% for the IL-based and modified IL-based process, respectively). In general, the cost analysis results indicate the IL-based processes are cost-efficient.

5. CONCLUSIONS The present work assessed the techno-economic performance of IL-based CO2 capture processes based on the rigorous thermodynamic modeling and process simulation results, taking a conventional MEA-based process as the reference case. The thermodynamic models related to the IL-based capture system, including the physicochemical properties and phase equilibria were developed. Phase equilibria of the CO2-IL binary systems were modeled by the R-K equation of state, Henry's law and NRTL model. Experiments were conducted to obtain the VLE data of the H2O-IL-MEA systems. The calculated results displayed good accordance with the experimental data, which indicated the models and parameters were reliable to be used in process simulation. 19



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The [Bpy][BF4]-MEA process display 15% and 11% reduction in regeneration heat duty and capture cost respectively compared to the reference MEA process. Similarly, the modified [Bpy][BF4]- MEA process with inter-cooling and lean vapor recompression can accomplish 31% and 13.5% reduction in regeneration heat consumption and capture cost, respectively. Although ionic liquid has high price, total capture cost of the IL-based processes are still reduced because solvent investment merely occupy less than 1.5% of the total capture cost. Detailed calculation of energy distribution explains that the reduction in regeneration heat duty is resulted from the contribution of both sensible and latent heat. Besides, the above two IL-based processes are better than the MEA process in term of circulated solvent, cooling water demand and total equivalent energy penalty. All in all, the IL-based solvent combining with the process modification could realize energy-saving and cost-efficient carbon capture, which provides a perspective capture technology in the future.

ASSOCIATED CONTENT Supporting Information Supporting Information provides the following tables and discussion: the calculation method and basis of capital and operation cost, obtained parameters of correlations for physicochemical properties prediction, experimental and calculated vapor pressures of three IL-MEA hybrid solvents, determination of stage numbers of the absorber and stripper, column profiles of the absorber and stripper of the MEA and IL-based processes, mass and energy balance information of the process streams. This information is available free of charge via the internet at http://pubs.acs.org. 20


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AUTHOR INFORMATION Corresponding Author * E-mail: [email protected] and [email protected] Tel./fax: +86-010-82544875.

ACKNOWLEDGEMENTS This research was supported by the Key Program of National Natural Science Foundation of China (21036007), the National Basic Research Program of China (973 Program) (No. 2013CB733506) and the National Natural Science Foundation of China (51274183, 51306194). We also appreciate the support from the Science and Technology Innovation Team of Cross and Cooperation of Chinese Academy of Sciences.

NOTATIONS List of symbols pis

P T Cp Z i*,RA

Z Ai Bi Ci Vi Tr M Np Pc Tb Tbr Tc Vc fi HiA

vapor pressure of pure component i pressure, Pa Temperature, K liquid heat capacity, J·mol-1·K-1 parameters of racket liquid molar volume model compressibility factor the first coefficient of Andrade liquid viscosity equation the second coefficient of Andrade liquid viscosity equation the third coefficient of Andrade liquid viscosity equation liquid molar volume, cm3·mol-1 reduced temperature molar mass, g·mol-1 total number of data points critical pressure, bar normal boiling temperature, K reduced temperature at the normal boiling point critical temperature, K critical volume, cm3·mol-1 fugacity of component i Henry’s law constant of component i in solvent A, Pa 21



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αlean αrich

the first coefficient of Henry’s law constant the second coefficient of Henry’s law constant the third coefficient of Henry’s law constant the fourth coefficient of Henry’s law constant composition of component i in liquid composition of component i in vapor activity coefficient the infinite dilution activity coefficient of component i total equivalent work, kJ ·kg-1 CO2 the reboiler temperature, K heat duty of reboiler, GJ ·ton-1 CO2 work compressor and blower, kJ ·kg-1 CO2 work of pumps, kJ ·kg-1 CO2 CO2 loading in lean solvent CO2 loading in rich solvent

Abbreviations IL MEA [Bmim][BF4] [Bmim][DCA] [Bpy][BF4] [Bmim][Tf2N] [Bmim][Ac] [Emim][OTf] GLE VLE FC-CS O.F. AIC LVR AARD Nabs Ns TCC TCI TOC VOC FOC

ionic liquid monoethanolamine 1-butyl-3-methylimidazolium tetrafluoroborate 1-butyl-3-methylimidazolium dicyanamide 1-butylpyridinium tetrafluoroborate 1-butyl-3-methylimidazolium bis((trifluoromethyl)sulfonyl)imide 1-butyl-3-methylimidazolium Acetate 1-ethyl-3-methylimidazolium trifluoromethanesulfonate gas-liquid phase equilibrium vapor-liquid phase equilibrium the fragment contribution-corresponding states method objective function absorber inter-cooling lean vapor recompression average absolute relative deviation stage number of absorber stage numbers of stripper total capture cost total capital investment total operating cost variable operating cost fixed operating cost

Greek letters ρ η σ

density, g·cm-3 viscosity, mPa·s surface tension, mN·m-1

aiA, biA ciA diA xi yi

γ

γ i∞

Weq Treb Qreb Wcomp Wpump

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λ ω

liquid thermal conductivity, W·m-1·K-1 acentric factor

Superscripts cal exp

calculated property experimental property

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FIGURE CAPTIONS Figure 1. The experimental apparatus for saturated vapor pressure measurement. Figure 2. Conventional process flow diagram. Figure 3. Modified process flow diagram. Figure 4. Experimental and calculated heat capacities(a), densities(b), viscosities(c), surface tensions(d) and thermal conductivities(e) of the ILs: dots and lines denote experimental data from literature and calculated value, respectively. Figure 5. Total pressures of the systems of CO2-[Bmim][BF4](a), CO2-[Bmim][DCA](b) and CO2-[Bpy][BF4](c): dots and lines denote experimental data from literature and calculated value. Figure 6. Experimental and predicted vapor pressure data for the binary system [H2O(1)-[Bmim][BF4] (2)] and ternary system [H2O(1)-[Bmim][BF4](2)-MEA(3)]: dots and lines denote experimental data reported here and calculated value, respectively. Figure 7. Experimental and predicted vapor pressure data for the binary system [H2O(1)-[Bmim][DCA] (2)] and ternary system [H2O(1)-[Bmim][DCA](2)-MEA(3)]: dots and lines denote experimental data reported here and calculated value, respectively. Figure 8. Experimental and predicted vapor pressure data for the binary system [H2O(1)-[Bpy][BF4](2)] and ternary system [H2O(1)-[Bpy][BF4](2)-MEA(3)]: dots and lines denote experimental data reported here and calculated value, respectively. Figure 9. Effect of αlean on the regeneration thermal energy requirement of MEA and IL-based processes. Figure 10. Regeneration heat duty and total equivalent energy penalty of three CO2 capture processes. Figure 11. Distribution of TCI of the three CO2 capture processes. Figure 12. Distribution of TCC of the three CO2 capture processes.

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Figure1. The experimental apparatus for saturated vapor pressure measurement. 1-a: working ebulliometer (Type CL-II); 1-b: reference ebulliometer (Type CL-II); 2: magnetic stirrer; 3: heating rod; 4: heating mantle; 5: thermometer; 6:condenser; 7:vacuum jacket; 8:inner casing; 9:buffer bottle; 10- manometer; 11-vacuum pump; 12-evacuation valve; 13-pressure adjusting valve

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Figure 2. Conventional process flow diagram.

Figure 3. Modified process flow diagram.

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Figure 4(a)

Figure 4(b)

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Figure 4(c)

Figure 4(d)

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Figure 4(e)

Figure 4. Experimental and calculated heat capacities(a), densities(b), viscosities(c), surface tensions(d) and thermal conductivities(e) of the ILs: dots and lines denote experimental data from literature and calculated value, respectively.

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Figure 5(a)

Figure 5(b)

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Figure 5(c)

Figure 5. Total pressures of the systems of CO2-[Bmim][BF4](a), CO2-[Bmim][DCA](b) and CO2-[Bpy][BF4](c): dots and lines denote experimental data from literature11,70,71 and calculated value.

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Figure 6. Experimental and predicted vapor pressure data for the binary system [H2O(1) - [Bmim][BF4] (2)] and ternary system [H2O(1)-[Bmim][BF4](2)-MEA(3)]: dots and lines denote experimental data reported here and calculated value, respectively.

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Figure 7. Experimental and predicted vapor pressure data for the binary system [H2O(1) [Bmim][DCA] (2)] and ternary system [H2O(1) -[Bmim][DCA](2)-MEA(3)]: dots and lines denote experimental data reported here and calculated value, respectively.

Figure 8. Experimental and predicted vapor pressure data for the binary system [H2O(1) - [Bpy][BF4] (2)] and ternary system [H2O(1)-[Bpy][BF4](2)-MEA(3)]: dots and lines denote experimental data reported here and calculated value, respectively. 36


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Figure 9. Effect of αlean on the regeneration thermal energy requirement of MEA and IL-based processes.

Figure 10. Regeneration heat duty and total equivalent energy penalty of three CO2 capture processes.

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Figure 11. Distribution of TCI of the three CO2 capture processes.

Figure 12. Distribution of TCC of the three CO2 capture processes.

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TABLE CAPTIONS Table 1. Prediction Equations for Some Temperature Dependent Properties Table 2. Calculation Measures of Total CO2 Capture Cost (TCC) Table 3. The Parameters of Temperature Dependent of Henry Constants Table 4. NRTL Binary Parameters of CO2, H2O and MEA with Different IL Table 5. Key Parameters for the MEA and IL-based Processes

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Table 1. Prediction Equations for Some Temperature Dependent Properties property

specific equation

No.

vapor pressure

C2 i ln Pi s = C1i + T + C3i

(1)

heat capacity

C pi = C1'i + C 2' i T + C3' i T 2

(2)

notation description

ref.

Pi s is vapor pressure in Pa and 46

T is in K, C1i−C3i are equation coefficients. Cpi is liquid molar heat capacity in J·kmol-1·K-1, C ' − C 3' i are

72

1i

equation coefficients. -3 ρ i is density in g·cm , Tci and

density

ρi =

M i Pci RTci  Z

*, RA i

(1 + d i (1 − Tr ) 

viscosity

ln ηi = Ai + Bi / T + Ci ln T

surface tension

σ i = C '' (1 − T / Tci ) ( C

thermal conductivity

1+ (1−T )2/7  r  

'' '' '' 2 '' 3 2 i + C3 i Tri + C4 iTri + C5 iTri )

1i

(3)

Pci are critical temperature and

73

pressure, respectively. Z i*,RA and

(4)

di are equation coefficients. ηi is the liquid viscosity in mPa·s, Ai, Bi and Ci are equation coefficients. σi is the liquid surface tension in

(5)

mN·m-1, C1''i − C 5''i are equation

46

72

coefficients. λi is the liquid thermal conductivity in W·m-1·K-1,

λi = C ''' + C ''' T + C ''' T 2 1i

2i

3i

(6)

C1'''i − C3'''i are equation

72

coefficients.

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Table 2. Calculation Measures of Total CO2 Capture Cost (TCC) cost item

calculation basis

TCC

annual capital cost (ACC), total operating cost (TOC)

ACC

total capital investment (TCI)

TCI

fixed capital investment (FCI),working capital, startup cost, initial solvent cost

FCI

direct cost (DC), indirect cost(IC)

DC, IC

purchased equipment cost (PEC) with a coefficient

TOC

variable operating cost (VOC), fixed operating cost (FOC)

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Table 3. The Parameters of Temperature Dependent of Henry Constants parametersa ILs

aiA

biA

ciA

d iA

[Bmim][BF4]

43.962

-10864.735

7.708

-0.119

[Bmim][DCA]

150.2171

-9598.69

-15.9565

-0.0379

[Bpy][BF4]

21.259

7113

-11.1

0.1146

parametersa: i denotes CO2 and A denotes various ionic liquid.

Table 4. NRTL Binary Parameters of CO2, H2O and MEA with Different IL component i CO2 H2O MEA CO2 H2O MEA CO2 H2O MEA

parameters

component j

[Bmim][BF4]

[Bmim][DCA]

[Bpy][BF4]

aij

aji

bij

bji

6.713 -0.9689 -0.265 -3.9246 1.1886 -0.365 81.456 -0.8891 0.165

-16.928 1.114 -0.771 -126.974 -5.514 -20.1705 6.1485 -6.084 -20.471

-2577.397 85 3.495 1178.485 -199 -271.995 -23369.5 248 185.9951

6048.75 732 10.866 32278.18 11510 28964.8 -2021.2 4962.67 9884.8

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Table 5. Key Parameters for the MEA and IL-based Processes process

MEA

a

solvent specific demand, m3·ton-1 CO2

21.88

17.07

15.85

CO2 lean loading(αlean)

0.25

0.20

0.20

CO2 rich loading (αrich)

0.469

0.454

0.478

∆ loading of rich and lean solvent

0.219

0.254

0.278

cooled lean solvent temperature, K

334.45

331.15

319.35

heated rich solvent temperature, K

376.85

379.85

366.85

reboiler temperature, K

392.75

399.95

398.85

regeneration heat duty, GJth·ton CO2

3.72

3.17

2.51

electrical power consumption, kWh·ton-1CO2

117.86

115.14

133.09

1047.01

985.89

920.41

143.46

114.53

132.08

-1

-1

total equivalent energy penalty, kJ·kg CO2 -1

cooling water specific requirement, ton·ton CO2

IL

b

M-IL

a

IL: the [Bpy]BF4-MEA capture process M-IL: the modified [Bpy]BF4-MEA capture process

b

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Thermodynamic Modeling and Assessment of Ionic Liquid-Based CO2

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