CO2 capture by methanol, ionic liquid, and their binary ...download.xuebalib.com/ljsE2hxk9lq.pdf · Rectisol, Fluor, Purisol, and Ifpexol.4-7 Methanol (MeOH) as in the Rectisol process

CO2 Capture by Methanol, Ionic Liquid, and Their Binary

Mixtures: Experiments, Modeling, and Process Simulation

Mohsen Taheri, Chengna Dai, and Zhigang Lei*

State Key Laboratory of Chemical Resource Engineering, Beijing University of Chemical

Technology, Box 266, Beijing 100029, China

ABSTRACT: The CO2 solubility data in the ionic liquid (IL) 1-allyl-3-methylimidazolium

bis(trifluoromethyl sulfonyl)imide, methanol (MeOH), and their mixture with different

combinations at temperatures of (313.2, 333.2, and 353.2 K) and pressures up to 6.50 MPa

were measured experimentally. New group binary interaction parameters of the predictive

UNIFAC-Lei model, which has been continually advanced by our group, were introduced by

correlating the experimental data of this work and the literature. The consistency between

experimental data and predicted results proves the reliability of UNIFAC-Lei model for

CO2-IL-organic solvent systems. The newly obtained parameters were incorporated into the

UNIFAC property model of Aspen Plus software to optimize a conceptual process developed

for the purification of a CO2-containing gas stream. The simulation results indicate that the

use of IL either mixed with MeOH or purely considerably lowers the process power

consumption, and improves the process performance in terms of CO2 capture rate and solvent

loss.

Keywords: CO2 capture, physical absorption, ionic liquid (IL), methanol (MeOH),

UNIFAC-Lei model, process simulation

Additional Supporting Information may be found in the online version of this article. ∗ Correspondence concerning this article should be addressed to Z. Lei at [email protected].

Separations: Materials, Devices and Processes AIChE JournalDOI 10.1002/aic.16070

This article has been accepted for publication and undergone full peer review but has not beenthrough the copyediting, typesetting, pagination and proofreading process which may lead todifferences between this version and the Version of Record. Please cite this article asdoi: 10.1002/aic.16070© 2018 American Institute of Chemical Engineers (AIChE)Received: Jan 19, 2017; Revised: Jul 27, 2017; Accepted: Dec 24, 2017

This article is protected by copyright. All rights reserved.

http://orcid.org/0000-0001-7838-7207

2

Introduction

The observation of carbon dioxide (CO2) emissions in recent years shows the severe

destructive effects of CO2 as the primary long-lived greenhouse gas on the global warming

and climate change issues clearly.1 In particular, a significant amount of CO2 is sent out to the

atmosphere annually in the accustomed processes, involving fuels combustion and

purification.2 The conventional industrial approaches for CO2 capture and separation mainly

include solvent absorption, sorbent adsorption and membrane filtration.3 Among others,

physical absorption processes are the most common methods for acid gas removal, where

CO2 concentration (partial pressure) in feed gas stream is high. The main advantage of a

physical solvent is its low required regeneration energy due to the weak interaction between

the solvent and the gas to be captured. The famous physical solvent processes include Selexol,

Rectisol, Fluor, Purisol, and Ifpexol.4-7

Methanol (MeOH) as in the Rectisol process is widely

applied in industry for synthesis gas (syngas) purification with some advantages over other

physical solvents, such as low viscosity, availability, non-corrosivity, and deep CO2 removal

along with other impurities removal. However, a high vapor pressure of MeOH can cause

substantial reduction in the solvent. Thus, the process must be operated at very low

temperature that results in enormous elevation of power consumption.2,7-11

On the other hand, in the past few years, the non-volatile sustainable ionic liquids (ILs)

have been recognized by many researchers as propitious physical absorbents for CO2 and

other acid gases as well.12-20

Although the previous experimental data showed that the lower

operating temperatures and higher pressures bring about more CO2 dissolved in ILs, CO2

solubility in ILs doesn’t decline as sharp as it does in MeOH when the temperature increases

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at a constant pressure. For example, at the pressure 1 MPa, when the temperature increases

from 298.2 to 323.2 K, the solubility of CO2 in MeOH drops by 43% (Ref. 21-26), while the

values are 30%, 30%, 29%, 33%, 28%, and 27% for [BMIM]+[BF4]

- (Ref. 27),

[BMIM]+[Tf2N]

- (Ref. 27), [BMPYR]

+[Tf2N]

- (Ref. 27), [BMIM]

+[PF6]

- (Ref. 28),

[HMIM]+[Tf2N]

- (Ref. 29), and [HMPY]

+[Tf2N]

- (Ref. 29), respectively. This suggests that

the fall in CO2 solubility due to the exothermic CO2 absorption process for ILs is lower when

compared to MeOH. Therefore, adding an IL to MeOH and capturing CO2 at a higher

temperature can compensate the rise in temperature.

To provide the concept of ameliorated CO2 capture processes characterized with less gas

stream contamination and lower operating costs due to the higher operating temperatures, an

imidazolium-based IL was added into MeOH to have both advantages of organic solvents and

ILs. The IL 1-allyl-3-methylimidazolium bis(trifluoromethyl sulfonyl)imide

([AMIM]+[Tf2N]

-) was selected in this work, because (1) the [Tf2N]

- anion has a strong

affinity towards CO2;30

(2) the [Tf2N]-based ILs with imidazolium cation belong to the most

thermally stable ILs;31

and (3) [AMIM]+[Tf2N]

- has been reported as a promising IL for CO2

capture by Nonthanasin et al.32

Thus, the solubility of CO2 in pure [AMIM]+[Tf2N]

-, pure

MeOH, and their binary mixtures (80 wt% [AMIM]+[Tf2N]

- + 20 wt% MeOH, 50 wt%

[AMIM]+[Tf2N]

- + 50 wt% MeOH, and 20 wt% [AMIM]

+[Tf2N]

- + 80 wt% MeOH) at

temperatures 313.2, 333.2, and 353 K and pressures up to 6.5 MPa was measured in this work

by using a static gas-liquid equilibrium view-cell technique.

A number of predictive thermodynamic models are available for modeling the solubility

of CO2 in ILs.33-42

Among others, the universal quasichemical functional-group activity

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coefficients model for ILs constantly developed by our group (UNIFAC-Lei model)43-50

has

attracted wide attention, and it produces solid outcomes in various specific applications due

to the unique advantages: (1) it provides quantitative accurate predictions;51

(2) its

formulation is simple, and thus the calculation speed is rapid; (3) its parameters can be

readily utilized in the modern simulation packages (e.g., Aspen Plus, PROII, and ChemCAD)

for simulation and optimization of the desired industrial processes;52-54

and (4) in terms of the

UNIFAC-Lei preferable decomposition approach, the introduced main groups and subgroups

together with their derived parameters are appropriate for the ILs-based designing

methodologies for specific applications.55-58

Thus, the UNIFAC-Lei model has been extended

in this work to predict the solubility of CO2 in the proposed solvents.

The remainder of this article is arranged as follows: firstly, the experimental data of this

work, along with those collected from the literature, were correlated by the UNIFAC-Lei

equations and its data fitting procedure, and the new group binary interaction parameters (αmn

and αnm) between CO2 and MeOH/[AMIM]+[Tf2N]

- groups were obtained. The predicted

results were compared with experimental data to justify the applicability of UNIFAC-Lei

model for CO2-IL-organic solvent integrated systems. Then, a semi-Rectisol flow

configuration with rigorous RadFrac equilibrium models was set up in the Aspen Plus to

simulate and analyze the processes including IL solvents. The simulations were compared

with those just using conventional pure MeOH solvent from the perspectives of energy

consumption, CO2 capture ability, and evaporative loss. For this purpose, the

[AMIM]+[Tf2N]

- compound was added to the Aspen Plus module, and the newly obtained

UNIFAC-Lei parameters were input so that process simulations were implemented based on

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the Aspen Plus UNIFAC property model.

Experimental Section

Materials

The IL [AMIM]+[Tf2N]

- (molecular mass 403.32 g.mol

-1, and mass fraction purity >

99%) was purchased from the Shanghai ChengJie Chemical Co. Ltd. Before the experiment,

under a vacuum rotary evaporator device, [AMIM]+[Tf2N]

- was degassed and dried for 24

hours at 335.2 K to remove its water traces and other volatile impurities. After this treatment,

its water content was reduced to below 200 ppm as measured by the Karl Fischer titration

(KLFX device, model KLS701). MeOH (mass fraction purity > 99.5%) was purchased from

the Tianjin Damao Chemical Reagent Factory. CO2 (mass fraction purity 99.995%) was

purchased from the Beijing Longkou City Gas Plant. Sulfuric acid (H2SO4) and sodium

sulfate (Na2SO4) were purchased from the Beijing Chemical Plant with mass fraction purities

of 95% to 98% for H2SO4 and > 99% for Na2SO4. The four latter chemicals were used

directly without further purification.

Apparatus and procedure

A high-temperature high-pressure view-cell apparatus, the detailed schematic flowsheet

of which has been depicted in our previous work,59

was used in this work to conduct the

experimental vapor-liquid equilibrium measurements of CO2 and the solvents. The operating

procedure is explained in what follows. In each experiment, after the operating temperature

was set to the predetermined value, a vacuum pump (2XZ-1) was used to purge out the air

inside the equilibrium cell. After evacuation, a certain amount of solvent (about 35 mL) was

injected into the cell. Then, CO2 was introduced to the solvent by a syringe pump, while the

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magnetic stirring system was working at a speed of about 200 rpm. The CO2 started to

dissolve in the solvent, and thus the pressure began to drop down gradually. In accordance

with Duhem’s theorem,60

it was presumed that gas-liquid equilibrium is established when at

each fixed operating temperature, the system pressure becomes invariable for more than 6

hours. After reaching the equilibrium, a small amount of solvent containing the dissolved

CO2 (3 to 5 g) was extracted from the lower part of the equilibrium cell (liquid phase), and

was trapped in a stainless steel sample bomb. Subsequently, the Na2SO4 aqueous solution

with pH between 2.5 and 3 adjusted by H2SO4, was used to degas the CO2 from the sampler

solution, and also to keep the volatile organic components which may emit along with CO2.

As in the gravimetric method, based on the weights of the unloaded sampler (m1), loaded

sampler (m2), sampler after CO2 release (m3), Na2SO4 solution flask before exposing to the

sampler (m4), and Na2SO4 solution flask after exposing to the sampler (m5), the mole fraction

of CO2 in the liquid phase was measured. Moreover, a small charge of gas mixture (about 400

µL), which was taken out by a sealed sampler from the upper part of equilibrium cell (gas

phase), was fed to a gas chromatograph (GC 4000A) equipped with a thermal conductivity

detector (TCD) and a 3 m × 3 mm Porapak-Q column to analyze the equilibrium gas phase

composition. The GC column temperature, injector temperature, and hydrogen flow rate -

acting as the carrier gas - were set at 403.2 K, 423.2 K, and 30 mL·min-1

, respectively.

A pressure gauge (XSE, 0-10 MPa) and a Pt100 thermocouple (XMTD-751W) with the

digital displays and accuracies of ±0.06% FS and ±0.5%, respectively, had been installed on

the view-cell apparatus to monitor the pressure and temperature of equilibrium cell. These

errors don’t bring about any significant variation in the CO2 solubility measurement. An

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electronic balance (CPA 1003S, Sartorius) with a precision of ±0.001 g was used for

weighing the sampler and the flask. To ensure the accuracy of experimental data, at least

three repeated runs were performed for each equilibrium data point. By employing the

following equations together with the so-called Kragten spreadsheet method,61

the expanded

uncertainties with a level of confidence of approximately 95% (coverage factor k = 2) were

calculated for CO2 mole fractions in the liquid phase (x1) and CO2 mole fractions in the gas

phase (y1), which are termed as U(x1) and U(y1), respectively;

1 1 2 3 4 5( ) ( , , , , )nx F m m m m m= (1)

22 2

1 1 1c 1 1 2 5

1 2 5

( ) ( ) ( )( ) ( ) ( ) ... ( )n n n

n

x x xu x u m u m u m

m m m

∂ ∂ ∂= + + + ∂ ∂ ∂

(2)

1 1 1 2 1 1 1 1 2 11 1

( ) ( ) ... ( ) ( ) ( ) ... ( ),n nx x x y y y

x yn n

+ + + + + += = (3)

2 2

1 1 1 1

1 11 1

(( ) ) (( ) )

( ) , ( )1 1

n n

i i

i i

x x y y

u x u yn n

= =

− −= =

− −

∑ ∑ (4)

( )2 2

1 1 1 1 1( ) ( ) ( ( ) ) , ( ) ( )c nU x u x MAX u x k U y u y k= + = (5)

where n represents the number of experimental runs, uc stands for the combined standard

uncertainty due to the electric balance precision, and u is the standard uncertainty due to

repeatability of the experiments. These values are listed in the tables in Supporting

Information along with the experimentally measured results.

UNIFAC-Lei Model

Model description

Upon combining the solution-of-functional-groups concept with a model based on the

UNIQUAC (universal quasi-chemical) theory, the original UNIFAC (UNIQUAC

functional-group activity coefficient) group-contribution method was initially proposed by

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Fredenslund et al. in 197562

for the estimation of liquid-phase activity coefficients. Since then,

it became notably popular among chemists and chemical engineers as a powerful predictive

tool of thermodynamic properties of solvent-containing systems. In 2005, it was applied to

ILs for the first time by Kato and Gmehling.63

However, they only introduced two main

groups with an inappropriate fragmentation approach. Later on in 2007, in an attempt by Lei

et al.,64

the UNIFAC model for ILs was improved in some degree. Afterwards, because of its

high potential and the lack of group parameters, we decided to comprehensively develop the

original UNIFAC model and extensively extend it to ILs, in a manner that up to now over 700

come-in-handy group interaction parameters for ILs have been published by our group.44-50

Like all the other group contribution methods, in the UNIFAC-Lei model the existing

molecules in mixture should be decomposed into distinct functional groups. Three

approaches to ILs decomposition have been put forth by researchers. (1) The IL molecule is

divided into two groups: one cation group, and one anion group.65,66

(2) As what Kato and

Gmehling63

did, the IL molecule is divided into several groups: one anion group, one cation

core group (e.g., imidazolium or pyrrolidinium ring), and several other groups like CH2 or

CH3 groups.67-70

(3) The IL is also fragmented into several but electrically neutral groups: one

group comprised of the anion and the skeleton of the cation, and several groups on the

side-chain of the cation.33,34,39,44-50,71

The first approach is rejected since it doesn’t consider

the structural variation of substituents on both the cation and the anion. In the second

approach, the electrically charged functional groups are introduced. However, the original

UNIFAC model has been established based on the Derr and Deal’s ASOG

(analytical-solution-of-groups) model72

and the Abrams and Prausnitz’s UNIQUAC model,73

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which both had been proposed for the solutions with nonelectrolyte molecules. More

specifically, the UNIFAC activity coefficient equation doesn’t contain a term such as the

Debye-Hückel accounting for the electrostatic contribution of electrically charged functional

groups. In this regard, the UNIFAC-Lei model adopts the third approach to dissociate the IL

into electrically neutral functional groups. Accordingly, the [AMIM]+[Tf2N]

- molecule here is

decomposed into three individual functional groups as shown in Figure 1. Meanwhile, CO2

and MeOH are considered as two separate groups as in the original UNIFAC model.

Activity coefficient formulation and parameters

In the UNIFAC model, the activity coefficient of species i ( iγ ), as a function of

temperature and composition, in a nonideal liquid mixture is calculated by

C Rln ln lni i iγ γ γ= + (6)

where Cln iγ accounts for the combinatorial contribution to activity coefficient, essentially

due to differences in shape and size of the groups; and Rln iγ stands for the residual

contribution to activity coefficient, essentially due to energetic interactions between the

functional groups. The combinatorial part Cln iγ includes the group volume (Rk) and the

group surface area (Qk) parameters, which either can be found directly in the references or, in

case of unavailability, can be determined by the van der Waals group volume (Vk) and the van

der Waals group surface area (Ak), respectively. By using either the molar volume correlation

method44

or the COSMO-RS method45

as proposed in our previous work, the Vk and Ak of the

functional groups of ILs can be calculated. The residual part Rln iγ includes two

temperature-independent group binary interaction parameters (αmn and αnm, αmn ≠ αnm) for

each pair of groups m and n. In principle, the group interaction parameters must be computed

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by means of correlating the experimental phase equilibrium data.

As listed in Table 1, 20 group binary interaction parameters in addition to 10 group

volume and surface area parameters are required in this modeling work. However, the group

interaction parameters between CH3OH and CO2 as well as those between C=C and CO2

groups had not been reported in the literature yet. Thus, the parameters for CH3OH and CO2

were first fixed by fitting the experimental solubility data of CO2 in MeOH collected from the

literature,21-26,74-82

and presented in Table S1 in Supporting Information. Afterwards, by

correlating the experimental solubility data of CO2 in [AMIM]+[Tf2N]

- from the literature

32

and from this work as presented in Tables S2 and S3 in Supporting Information, the

parameters for C=C and CO2 were estimated. Other group interaction, group volume and

surface area parameters are collected from the references.44,46,83-85

Procedure for estimating the group interaction parameters

For the CO2-IL or -MeOH binary and CO2-IL-MeOH ternary systems, the gas-liquid

equilibrium states at low to moderate pressures are expressed as

s

1 1 1 1 1 1( , , )y P T P y x Pϕ γ= (7)

given CO2 as the component number 1. 1y and 1x are the mole fractions of CO2 in the gas

and liquid phases, respectively; P is the system pressure; 1 1( , , )T P yϕ refers to the fugacity

coefficient of CO2 in the gas phase evaluated using the Span and Wagner’s equation of

state;86

1γ is the activity coefficient of CO2 in the liquid phase calculated by the

UNIFAC-Lei model; and s

1P is the saturated vapor pressure of pure CO2 obtained by using

the extrapolated Antoine equation as proposed by Shiflett and Yokozeki.87

The group binary interaction parameters were estimated by minimizing the average

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relative deviation (ARD) of experimental data and calculated values via the following

objective function:

cal exp

1 exp

1min

N x xOF

N x

− =

∑ (8)

where expx is the mole fraction of CO2 in the liquid phase obtained experimentally; calx is

the one calculated by the UNIFAC-Lei model; and N is the number of data points employed

in the fitting procedure which is performed by using the generalized reduced gradient (GRG)

nonlinear optimization algorithm of the Solver function of Microsoft Excel 2016.

Process Simulation

Pure MeOH processes operate at reduced temperatures, sometimes as low as 200 K.88

Besides complex refrigeration schemes and refrigerants demand, there are two CO2 emission

sources in such a kind of processes. (1) Direct CO2 emission: coming from the CO2 escaped

from the process itself. (2) Indirect CO2 emission: coming from the fuel combustion

processes in power plants to supply the required energy of the process, especially the heavy

power requirement of refrigeration system.10,89

Therefore, by focusing on the higher- or

ambient-temperature processes, the indirect CO2 emission can be diminished or even omitted.

Hereby, by an adaptation of the Rectisol-based processes layouts,10,89,90

a conceptual CO2

capture process configuration has been developed, and its simulations have been carried out

in Aspen Plus environment. To assess and interpret the separation performance at different

temperatures for different solvents, five absorption processes as for pure MeOH at 243.2 K,

mixed MeOH-IL at 243.2 K, pure IL at 243.2 K, pure IL at 265.7 K, and pure IL at 288.2 K

were simulated and the results were compared with each other.

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Process scheme

The process flowsheet depicted in Figure 2, is made up of one absorption tower (TW1),

one rectification column (TW2), one flash drum (TW3), two feed mixers (MX1, and MX2),

four heat exchangers (HE1, HE2, HE3, and HE4), one compressor (CM1), and one solvent

recycle pump (PM1). In all the five cases investigated, the raw gas to be treated, with 47% H2,

34% CO2, and 19% CO mole composition, enters the MX2 at 6 MPa and 298.2 K. After

mixing with the retrieved gas from the top of TW2, it is cooled down by HE2 and fed to the

bottom of the TW1. In each process, all the heat exchangers are set at the corresponding

operating temperature, i.e., 243.2, 265.7, or 288.2 K. Makeup solvent is added to the recycled

solvent pumped from the bottom of TW3, and after being cooled down, the solvent is fed to

the top of TW1. TW1, in which the CO2 is absorbed in an exothermic process, is equipped

with one side-draw liquid cooling loop. The purified gas is withdrawn from the top of TW1,

while the CO2-loaded solvent enters the top stage of TW2. The CO and H2 gases co-absorbed

with CO2 are recovered in this rectification column equipped with reboiler, and are recycled

back to the TW1. The solvent is then conveyed to the vacuum flash drum to desorb CO2.

For each of the four low temperature processes with defined flow rates of raw gas and

solvent feeds, parametric sensitivity analysis of the operative variables was performed to

acquire their optimal values. First, the runnable range and optimum pattern of every

parameter are elucidated, setting the captured CO2 rate leaving out the flash drum as the main

performance measure of the process. From Tables S4a to S4d in Supporting Information,

theoretical stage number of TW1, pressure of TW1, pressure of TW3, and temperature of

TW3 are perceived as the decisive parameters since they influence the production of captured

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CO2 to a greater extent than other three parameters, i.e., number of stages, pressure, and

boilup ratio of TW2. Nevertheless, TW2 and TW3 variables play a dual role as the retrieved

gas yield relies on the former and the solvent loss strongly depends on the latter. Therefore,

under multioptimal consideration, fixing TW1 pressure and number of stages, TW2 and TW3

variables are then defined by way of compensation. However, this approach doesn’t hold for

the 288.2 K pure IL process since its higher solvent flow rate prevents the simulation runs

from convergence in most of the cases. In this context, taking the knowledge of the low

temperature processes into account, trial and error was used to reach the proximity of its

optimal conditions. Ultimately, the operating conditions and design specifications of five CO2

capture processes established in this study are given in Table 2.

UNIFAC property method

The IL [AMIM]+[Tf2N]

- was not available in the Aspen Plus databank. Thus, it was

added by the User Defined as a conventional component. The [AMIM]+[Tf2N]

- critical

properties were obtained using the group contribution method developed by Valderrama et

al.91

Its viscosity and liquid molar volume data, and liquid surface tension data were collected

from Ref. 92 and Ref. 93, respectively. Its ideal gas heat capacities and liquid heat capacities

were calculated using the Joback group contribution model extended for ILs by Ge et al.94

During the whole simulation and modelling procedures using Aspen Plus software,

UNIFAC model was selected. The parameters listed in Table 1 in addition to those for H2 and

CO functional groups are a prerequisite to run the simulations. H2 and CO group volume and

surface area parameters as well as their binary interaction parameters with all the existing

functional groups except those with C=C and CH3OH groups have been published in our

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previous work.47

However, the group interaction parameters between H2/CO groups and C=C

group cannot be acquired since there is no relevant data available in the literature to enable to

calculate them. Therefore, likewise the binary interaction parameters between two gaseous

groups, they were input as zero. Nevertheless, the references experimental solubility data of

H2 in MeOH25,82,95-98

and CO in MeOH25,99-101

collected and presented in Tables S5 and S6 in

Supporting Information were used to derive the group interaction parameters of CH3OH/H2

and CH3OH/CO. The ARDs between experimental data and calculated results for MeOH-H2

and MeOH-CO systems are 9.35% and 8.94%, respectively. In this case, the fugacity

coefficient of H2 and CO gases in Eq. 7 were calculated using the Peng-Robinson (PR)

equation of state, and the saturated vapor pressure of H2 and CO were calculated using the

extrapolated Antoine equation as proposed by Fogg and Gerrard.102

These additional

parameters used in this work are given in Table 3. The most recent UNIFAC-Lei parameter

matrix for IL and non-IL systems is illustrated in Figure 3.

Results and Discussion

Solubility of CO2 in pure IL and pure MeOH: experiments and predictions

The solubility data of CO2 in pure [AMIM]+[Tf2N]

- and pure MeOH at three

temperatures (313.2, 333.2, and 353.2 K) and pressures up to 6.5 MPa were measured

experimentally, and the solubility data expressed on the dimensionless mole-fraction basis are

listed in Tables S3 and S7, respectively. To validate the accuracy of experimental data, our

solubility data were compared with those available in the literature. The results fall within the

acceptable norms with the highest ARD of 6.30%, as depicted in Figures S1 to S4 in

Supporting Information. The solubility of CO2 in pure IL and pure MeOH follows the normal

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trend, that is, increasing with decreasing the temperature and increasing the pressure.

Furthermore, while the IL is assumed as non-vaporizable, MeOH appears in the gas phase. As

shown in Figure S5 in Supporting Information, the amount of MeOH vaporized into the gas

phase increases as the temperature goes up. This would add energy consumption to the

subsequent MeOH recovery in the solvent regeneration step.

The previously obtained parameters from the experimental data correlation in

UNIFAC-Lei model are used here to predict the CO2 mole fractions in pure MeOH and pure

[AMIM]+[Tf2N]

-. The comparisons between experimental data and predicted values for the

IL-CO2 and MeOH-CO2 systems are illustrated in Figures 4a and 4b, respectively. They agree

well with an ARD of 3.22% for IL-CO2 system and an ARD of 9.87% for MeOH-CO2 system.

In addition, 312 CO2-MeOH data points presented in Table S1 and 39 CO2-[AMIM]+[Tf2N]

-

data points collected from the references presented in Table S2 give the ARDs equal to 8.31%

and 3.61%, respectively. This confirmed that the UNIFAC-Lei model can reliably predict the

solubility of CO2 in pure ILs and pure organic solvents.

Solubility of CO2 in binary mixtures of IL and MeOH: experiments and predictions

The solubility data of CO2 in the binary mixtures of [AMIM]+[Tf2N]

- and MeOH at

three different mass fraction combinations (0.8 IL + 0.2 MeOH, 0.5 IL + 0.5 MeOH, and 0.2

IL + 0.8 MeOH) at temperatures of 313.2, 333.2, and 353.2 K and pressures up to 6.5 MPa

were measured experimentally, as listed in Table S8 in Supporting Information. Under the

same temperature and pressure, the solubility of CO2 in mixed solvents goes up as the amount

of IL increases. Furthermore, Figures S6 to S8 in Supporting Information give the

information on the influence of the addition of IL into MeOH on the gas phase. It can be seen

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that adding a certain amount of IL into MeOH can lead to a relative decrease of CH3OH in

the gas phase. This means that solvent loss and consequently the required regeneration energy

for the mixed IL-MeOH solvents are less than those for pure MeOH.

By incorporating the predetermined CH3OH/CO2 and C=C/CO2 group parameters into

the UNIFAC equations and evaluating the activity coefficients, the predicted CO2 solubility

data for the binary IL-MeOH solvents are obtained. Comparison between experimental data

and predicted values is depicted in Figure 5, showing that they are in good agreement with an

ARD of 4.45%. Therefore, the UNIFAC-Lei parameters obtained by correlation of the pure

solvent experimental data can be effectively applied to the binary solvents, enabling to

extrapolate from the current single solvent CO2 capture processes to the optimized more

efficient mixed solvents processes.

Prediction of Henry’s constants of CO2

The magnitude of solubility of CO2 gas in each solvent at respective operating

temperature can be described by Henry’s constant. Expressing on the mole fraction basis, it

has the unit of pressure and is defined as

1 1

s s1 1 11 1 1 1 1

0 01

( , , )( ) lim lim ( )

x x

y P T P yH T P P T

x

ϕγ γ ∞

→ →= = ≅ (9)

where 1γ∞ is the activity coefficient of CO2 in solvent as it approaches infinite dilution. A

smaller value of the Henry’s constant implies a larger gas solubility.

Herein, the Henry’s constants of CO2 in five solvents (pure IL, 80 wt% IL + 20 wt%

MeOH, 50 wt% IL + 50 wt% MeOH, 20 wt% IL + 80 wt% MeOH, and pure MeOH) at

different temperatures (313.2, 333.2, and 353.2 K) have been obtained experimentally using

linear extrapolation at x1→0 from the fugacity divided by mole fraction. Whereupon, the

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activity coefficients at infinite dilution predicted by UNIFAC-Lei model were used to obtain

the Henry’s constants. As shown in Figure 6, good agreement between the experimental and

predicted CO2 Henry’s constants for [AMIM]+[Tf2N]

-, MeOH, and their binary mixtures is

observed. Moreover, the low Henry’s constants for pure IL even at temperature of 353.2 K

are notable, demonstrating the good absorption capability of ILs in capturing CO2.

Simulation results and analysis

Simulation results of the purified gas stream leaving the absorber TW1, and the captured

CO2 gas stream leaving the flash drum TW3 are given in Table 4. Table 5 gives the

information on power penalty of the processes, which is comprised of heat duties of four heat

exchangers, heat duty of TW2 reboiler, heat duty of TW3 flash drum, and electricity demand

of CM1 and PM1. Table 6 provides the information on separation performance of the five

processes in terms of H2 and CO fuel gases production rate, CO2 capture rate, CO2 venting

rate in the purified gas stream, and solvent loss rate.

In order to analyze the efficiency of the five simulated processes on producing nearly the

same amount of fuel gases, the pure MeOH process operating at 243.2 K is taken on as the

benchmark. It can be seen that (1) by employing the mixed MeOH-IL solvents and applying

the exactly same design specifications and operating conditions as the pure MeOH solvent,

less energy is required while better separation performance is achieved. Thus, adding IL to

MeOH is in favor for Rectisol-based processes. (2) The pure IL process operating at 243.2 K

is distinctly more efficient than the pure MeOH process since its power penalty of the

energy-intensive blocks, i.e., TW2 reboiler and TW3, along with the CO2 venting rate decline

sharply, while there is more CO2 captured and almost no solvent loss. (3) In the two

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higher-temperature pure IL processes, i.e., at 265.7 and 288.2 K, operating costs go down

substantially since the energy consumption of each block is less, and there is almost no

solvent loss. This is apart from the elimination of the negative environmental impacts caused

by volatilization loss of the organic solvent. (4) The 265.7 and 288.2 K absorption processes

using pure IL as solvent, require 64% and 32% less heat duty for their refrigeration systems,

respectively. This is particularly important since all the pure MeOH processes have

energy-intensive cooling systems in common.

To quantitatively compare the energy requirement of the processes, block specific

energy penalty (BSEP) is defined as

22

[ ]block heat duty or energy demand GJBSEP

kmol captured COflowrateof captured CO= (10)

The BSEPs of the energy-consuming blocks (i.e., heat exchangers, rectification column

reboiler, and flash drum) in the IL + MeOH (50 + 50 wt%) process, in which solvent and raw

gas flow rates are the same as those in the pure MeOH process, decrease from 0.055, 0.171,

and 0.127 GJ/kmol captured CO2 to 0.027, 0.096, and 0.074 GJ/kmol captured CO2,

respectively. Furthermore, by increasing the temperature from 243.2 to 265.7 K (increased

22.5 K) for pure IL process, the BSEPs for heat exchangers, TW2 reboiler, and TW3, varies

from 0.034, 0.018, and 0 to 0.022, 0.070, and 0.050, respectively. However, if the temperature

further increases 22.5 K from 265.7 to 288.2 K, double amount of solvent will be needed, and

the corresponding BSEPs reach 0.014, 0.119, and 0.108, respectively. The values of process

specific energy penalty (PSEP) calculated as the sum of BSEPs for heat exchangers, TW2

reboiler, and TW3 are given in Table 7. As can be seen, in comparison with the pure MeOH

process, all the four processes including IL need much less energy to capture 1 kmol of CO2.

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When considering environmental benignity together, the pure IL process operating at 243.2 K

is the most efficient since its PSEP, vented CO2 rate, and solvent loss rate are at the very least

while its CO2 capture rate is the highest.

We take a further step to study the indirect CO2 emission discharged into the atmosphere

by running the flowsheets. The entire power demand is defined by the process electricity

consumption (PEC):

PM1 CM1 HEs TW2 TW3( )PEC W W Q Q Qη= + + + + (11)

where W is the electricity, and Q is the heat duty required for the designated blocks. η is the

efficiency factor for converting thermal to electrical energy assumed at 0.38. PEC has the

unit of MW. Meanwhile, the average CO2 emission factor (EFCO2) for electricity generation

in USA was reported at 11.296 kmol/MW·h in 2015.107

The process indirect CO2 emission

(PIECO2) is then calculated by

2

1

CO2 CO [kmol h ]PIE PEC EF −= × ⋅ (12)

PIECO2 is the indirect CO2 emission resulting from the electricity supply to the cooling (HEs

and TW3), heating (TW2), and compression blocks (CM1 and PM1). As shown in Figure 7,

using pure IL or mixed with MeOH brings about a massive reduction in CO2 emission from

power plants. Indirect CO2 generated by the 243.2 K pure MeOH process is six times larger

than that of its 243.2 K pure IL counterpart.

Moreover, while elevated pressure is essential to CO2 physical absorption, in an attempt

to examine a lower-pressure condition, pressures of the feed gas and TW1 in the 243.2 K

pure IL process are brought down to 4 MPa. It revealed that by equipping TW1 with another

side-draw cooling loop (7+7+6), the captured CO2 rate decreases only by 1.52%, still

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separating 98.32% (5951.13 kmol·h-1

) of the CO2 fed to the system, with no considerable

change to other performance measures, e.g., PSEP equals to 0.047 GJ/kmol captured CO2.

Viscosity, toxicity, and cost of the deployed IL are of significant practical value to its

industrial exploitation. [AMIM]+[Tf2N]

- belongs to the family of low-viscous ILs since it has

a viscosity of 31.9 mPa·s at 298.2 K, far below the acceptable limit of 100 mPa.s for

industrial applications.92,108-110

High CO2 uptake in a pure IL as well as the addition of MeOH

to an IL also can cause a drastic reduction in viscosity of the fluid. Liu et al. demonstrated

that upon the dissolution of 50 mol% CO2 at 313.2 K and 6 MPa, the viscosity of

[BMIM]+[PF6]

- drops from 92 mPa·s to 32 mPa·s.

111 Therefore, in this work, viscosity of

[AMIM]+[Tf2N]

- doesn’t seem to be a barrier to its implementation. To our knowledge, the

only direct toxicity assessment of [AMIM]+[Tf2N]

- was reported towards human HeLa cell.

112

Out of 40 distinct ILs ranging 48 h EC50 values from 80 to 34000 µmol·L-1

, [AMIM]+[Tf2N]

-

has a value of 1630 µmol.L-1

(higher values denote lower toxicity). In general, although for

some biological targets [Tf2N]- might be regarded as a toxic anion, the allyl-substituted

[AMIM]+ can moderate the harmfulness degree whereby the cations containing smaller

number of carbon atoms in their substituents demonstrate lower toxicity.112-116

Finally, ILs

typically costs up to 100 times more than organic solvents.117

For our experiment, there was

an 80-fold difference in price between 1 kg [AMIM]+[Tf2N]

- and 1 kg MeOH. Accordingly,

with respect to solvent losses of the two 243.2 K pure MeOH and pure IL processes in Table

6, MeOH is 298 times costlier than [AMIM]+[Tf2N]

-. In this regard, the initial cost of IL

solvent is higher; however, the operating expense due to its running losses is much lower. The

vaporized MeOH in the acid gas stream may be further water washed, which adds on the

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process cost in turn. Therefore, the higher price of [AMIM]+[Tf2N]

- doesn’t necessarily

impair the commercial promise of its implementation.

Development and techno-economic analysis of a whole IL-integrated sour gas treating

process requires intermolecular knowledge of other compounds such as AR, CH4, H2S, N2,

and COS with ILs, which will be addressed in the future work.

Conclusions

In this work, the experimental CO2 solubility data in pure IL, pure MeOH, and mixed

IL-MeOH solvents at different temperatures (313.2, 333.2, and 353.2 K) and pressures up to

6.5 MPa were measured. Adding certain quantities of IL to MeOH brings about an increase in

CO2 solubility in the liquid phase and a decrease in the volatile MeOH content in the gas

phase. Mole fraction of CO2 in pure [AMIM]+[Tf2N]

- reaches around 0.6, 0.5, and 0.4 at

pressure of 6 MPa and temperatures of 313.2, 333.2, and 353.2 K, respectively. The

experimental CO2 solubility data in pure [AMIM]+[Tf2N]

- and pure MeOH along with those

exhaustively collected from the literature were correlated with the UNIFAC-Lei model to

obtain the group interaction parameters for further prediction of CO2 solubility in the solvents

investigated. Both experimental data and predicted values agree well with the ARDs below

10%. Moreover, since process simulation involving H2 and CO compounds has to be carried

out in this work, the H2 and CO solubility data in MeOH coming from the literature were

used to correlate the CH3OH/H2 and CH3OH/CO parameters. Together with the CH3OH-CO2

and C=C-CO2 parameters, 8 new group binary interaction parameters were evaluated and

added to the current UNIFAC-Lei parameter matrix.

The newly obtained UNIFAC-Lei parameters are important because, as what we did here,

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they can be easily incorporated into the modern process simulation softwares to design and

optimize the absorption processes. A non-volatile physical solvent with a high CO2 uptake

capacity at a higher temperature would improve the efficiency of the currently widely

industrialized energy-intensive CO2 capture processes, which should be operated at deep low

temperatures. Therefore, the comparison between pure MeOH process and four other

IL-integrated processes for CO2 capture shows the practical applicability and effectiveness of

ILs in capturing CO2. Besides having the advantage of less or almost no solvent loss, the

mixed IL-MeOH process operating at low temperature or the pure IL processes operating at

either low or high temperatures lessen the power consumption substantially, while producing

nearly the same amount of fuel gases. This also leads to a massive reduction in indirect CO2

emission sourcing from the power plants. Moreover, both mixed IL-MeOH and pure IL

processes using the same amount of solvent as the pure MeOH process and operating at the

same temperature, i.e., 243.2 K, capture more CO2 and release less CO2 along with the

purified gases.

Acknowledgements

This work was financially supported by the National Natural Science Foundation of China under

grants (Nos. 21476009, 21406007, and U1462104), and by the Higher Education and High-quality and

World-class Universities (PY201608).

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phase equilibria using a variable-volume view cell. Fluid Phase Equilib. 1996;123:215-230.

82. Bezanehtak K, Combes GB, Dehghani F, Foster NR. Vapor-liquid equilibrium for binary systems of

carbon dioxide + methanol, hydrogen + methanol, and hydrogen + carbon dioxide at high pressures. J

Chem Eng Data. 2002;47:161-168.

83. Nocon G, Weidlich U, Gmehling J, Menke J, Onken U. Prediction of gas solubilities by a modified

UNIFAC equation. Fluid Phase Equilib. 1983;13:381-392.

84. Skjold-Jørgensen S, Kolbe B, Gmehling J, Rasmussen P. Vapor-liquid equilibria by UNIFAC group

contribution. Revision and extension. Ind Eng Chem Process Des Dev. 1979;18:714-722.

85. Hansen HK, Rasmussen P, Fredenslund A, Schiller M, Gmehling J. Vapor-liquid equilibria by

UNIFAC group contribution. 5. Revision and extension. Ind Eng Chem Res. 1991;30:2352-2355.

86. Span R, Wagner W. A new equation of state for carbon dioxide covering the fluid region from the

triple-point temperature to 1100 K at pressures up to 800 MPa. J Phys Chem Ref Data.

1996;25:1509-1596.

87. Shiflett MB, Yokozeki A. Solubility and diffusivity of hydrocarbons in room-temperature ionic liquids.

AIChE J. 2006;52:1205-1219.

88. Kohl AL, Nielsen RB. Gas purification, 5th ed. Houston: Gulf Publishing Co., 1997.

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89. Sun L, Smith R. Rectisol wash process simulation and analysis. J Clean Prod. 2013;39:321-328.

90. Yang S, Qian Y, Yang S. Development of a full CO2 capture process based on the Rectisol was

technology. Ind Eng Chem Res. 2016;55:6186-6193.

91. Valderrama JO, Rojas RE. Critical properties of ionic liquids. Revised. Ind Eng Chem Res.

2009;48:6890-6900.

92. Wlazło M, Karpińska M, Domańska U. Thermodynamics and selectivity of separation based on

activity coefficients at infinite dilution of various solutes in 1-allyl-3-methylimidazolium

bis{(trifluoromethyl)sulfonyl}imide ionic liquid. J Chem Thermodyn. 2016;102:39-47.

93. Domańska U, Królikowski M, Zawadzki M, Wróblewska A. Phase equilibrium investigation with

ionic liquids and selectivity in separation of 2-phenylethanol from water. J Chem Thermodyn.

2016;102:357-366.

94. Ge R, Hardacre C, Jacquemin J, Nancarrow P, Rooney DW. Heat capacities of ionic liquids as a

function of temperature at 0.1 MPa. Measurement and prediction. Ind Eng Chem Res.

2008;53:2148-2153.

95. Descamps C, Coquelet C, Bouallou C, Richon D. Solubility of hydrogen in methanol at temperatures

from 248.41 to 308.20. Thermochima Acta. 2005;430:1-7.

96. Liu Q, Takemura F, Yabe A. Solubility of hydrogen in liquid methanol and methyl formate at 20 oC to

140 oC. J Chem Eng Data. 1996;41:1141-1143.

97. Walnwright MS, Ahn T, Trimm DL. Solubility of hydrogen in alcohols and esters. J Chem Eng Data.

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98. Choudhary VR, Sane MG, Vadgaonkar HG. Solubility of hydrogen in methanol containing reaction

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99. Liu Q. Solubility of carbon monoxide in liquid phase methyl formate and methanol at various

pressures and temperatures. J of the M.E.S.J. 1998;4:300-305 (in Japanese).

100. Liu Q, Takemura F, Yabe A. Solubility and diffusivity of carbon monoxide in liquid methanol. J Chem

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101. Tonner SP, Walnwright MS, Trimm DL. Solubility of carbon monoxide in alcohols. J Chem Eng Data.

1983;28:59-61.

102. Fogg PGT, Gerrard W. Solubility of Gases in Liquids: A Critical Evaluation of Gas/Liquid Systems in

Theory and Practice. New York: Wiley, 1991.

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103. Gmehling J, Rasmussen P, Fredenslund A. Vapor-liquid equilibria by UNIFAC group contribution.

Revision and extension. 2. Ind Eng Chem Process Des Dev. 1982;21:118-127.

104. Kleiber M. An extension to the UNIFAC group assignment for prediction of vapor-liquid equilibria of

mixtures containing refrigerants. Fluid Phase Equilib. 1995;107:161-188.

105. Wittig R, Lohmann J, Gmehling J. Vapor-liquid equilibria by UNIFAC group contribution. 6.

Revision and extension. Ind Eng Chem Res. 2003;42:183-188.

106. Dong L, Zheng D, Wu X. Working pair selection of compression and absorption hybrid cycles

through predicting the activity coefficients of hydrofluorocarbon + ionic liquid systems by the

UNIFAC model. Ind Eng Chem Res. 2012;51:4741-4747.

107. United States Electricity Profile 2015. U.S. Energy Information Administration. Release date: January

17, 2017. Available at: http://www.eia.gov/electricity/state/unitedstates; accessed July 18, 2017.

108. Seiler M, Jork C, Kavarnou A, Arlt W, Hirsch R. Separation of azeotropic mixtures using

hyperbranched polymers or ionic liquids. AIChE J. 2004;50:2439-2454.

109. Orchillés AV, Miguel PJ, Vercher E, Martínez-Andreu A. Using 1-ethyl-3-methylimidazolium

trifluoromethanesulfonate as an entrainer for the extractive distillation of ethanol + water mixtures. J

Chem Eng Data. 2010;55:1669-1674.

110. Chai M, Jin Y, Fang S, Yang L, Hirano S, Tachibana K. Ether-functionalized pyrazolium ionic liquids

as new electrolytes for lithium battery. Electrochimica Acta. 2012;66:67-74.

111. Liu Z, Wu W, Han B, Dong Z, Zhao G, Wang J, Jiang T, Yang G. Study on the phase behaviors,

viscosities, and thermodynamic properties of CO2/[C4mim][PF6]/methanol system at elevated

pressures. Chem. Eur. J. 2003;9:3897-3903.

112. Wang X, Ohlin CA, Lu Q, Fei Z, Hu J, Dyson PJ. Cytotoxicity of ionic liquids and precursor

compounds towards human cell line HeLa. Green Chem. 2007;9:1191-1197.

113. Ventura SPM, Gonçalves AMM, Sintra T, Pereira JL, Gonçalves F, Coutinho JAP. Designing ionic

liquids: the chemical structure role in the toxicity. Ecotoxicology. 2013;22:1-12.

114. Egorova KS, Ananikov VP. Toxicity of ionic liquids: eco(cyto)activity as complicated, but

unavoidable parameter for task-specific optimization. Chemsuschem. 2014;7:336-360.

115. Pretti C, Chiappe C, Baldetti I, Brunini S, Monni G, Intorre L. Acute toxicity of ionic liquids for three

fresh water organisms: Pseudokirchneriella subcapitata, Daphnia magna and Danio rerio. Ecotox.

Environ. Safe. 2009;72:1170-1176.

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32

116. Frade RFM, Matias A, Branco LC, Afonso CAM, Duarte CMM. Effect of ionic liquids on human

colon carcinoma HT-29 and CaCo-2 cell lines. Green Chem. 2007;9:873-877.

117. Plechkova NV, Seddon KR. Applications of ionic liquids in the chemical industry. Chem. Soc. Rev.

2008;37:123-150.

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33

Table captions

Table 1. Group Interaction (αmn and αnm), Group Volume (Rk) and Surface Area (Qk)

Parameters Involved in the UNIFAC-Lei Model

Table 2. Operating Conditions and Design Specifications of Five CO2 Capture Processes

Table 3. Additional UNIFAC-Lei Group Interaction Parameters (αmn and αnm) Used in

the Process Simulation

Table 4. Details of the Purified Gas and Captured CO2 Outflow Streams in Each of the

Five CO2 Capture Processes

Table 5. Power Penalty Comprised of Blocks Energy Demand for Each of the Five CO2

Capture Processes

Table 6. Separation Performance of the Five CO2 Capture Processes

Table 7. Block Specific Energy Penalties (BSEPs) and Process Specific Energy Penalty

(PSEP) for the Five CO2 Capture Processes

Page 33 of 49

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34

Figure captions

Figure 1. Decomposition of the IL [AMIM]+[Tf2N]

- into individual functional groups.

Figure 2. Flowsheet of the CO2 capture processes developed in this work.

Figure 3. The most recent UNIFAC-Lei parameter matrix.

UNIFAC parameters formerly available in the literature;83-85,103-105

parameters published in our

previous works;44-50

parameters have been submitted for publication; parameters added by other

researchers using the UNIFAC-Lei method;106

new parameters evaluated in this work; parameters

are not available.

Figure 4. Solubility of CO2 (1) in pure [AMIM]+[Tf2N]

- (a), and pure MeOH (b).

Lines, predicted results by the UNIFAC-Lei model; scattered points, experimental data. (▲) 313.2 K; (●)

333.2 K; (■) 353.2 K.

Figure 5. Solubility of CO2 in pure [AMIM]+[Tf2N]

-, pure MeOH, and their binary

mixtures at 313.2 K (a), 333.2 K (b), and 353.2 K (c).

Lines, predicted results by the UNIFAC-Lei model; scattered points, experimental data. ( ) pure

[AMIM]+[Tf2N]

-; (■) 80 wt% [AMIM]

+[Tf2N]

- + 20 wt% MeOH; (●) 50 wt% [AMIM]

+[Tf2N]

- + 50 wt%

MeOH; (▲) 20 wt% [AMIM]+[Tf2N]

- + 80 wt% MeOH; (★) pure MeOH.

Figure 6. Henry’s constants of CO2 (1) in pure [AMIM]+[Tf2N]

-, pure MeOH, and their

binary mixtures.

Lines and unfilled scattered points, predicted results by the UNIFAC-Lei model; filled scattered points,

experimental data. (■) 313.2 K; (●) 333.2 K; (▲) 353.2 K.

Figure 7. Indirect CO2 generated from electricity supply to each of the processes

(PIECO2, comprised of compression, cooling, and heating facilities).

Page 34 of 49

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AIChE Journal


Figure 1. Decomposition of the IL [AMIM]+[Tf2N]

- into individual functional groups.

Page 35 of 49

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AIChE Journal


Figure 2. Flowsheet of the CO2 capture processes developed in this work.

Page 36 of 49

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AIChE Journal


29

28

27

26

25

24

23

22

21

20

19

18

17

16

15

14

13

12

11

10

9

8

7

6

5

4

3

2

1C

H2

C=

C

AC

H

AC

CH

2

OH

CH

3O

H

H2O

AC

OH

CH

2C

O

CH

O

CC

OO

HC

OO

CH

2O

CF

2

CH

F

CH

F2

CH

F3

CO

2

C2H

4

CH

4

O2

CO

H2

SO

2

H2S

N2O

DM

F

NM

P

Thio

ph

ene

[MIM

][B

F4]

[MIM

][T

f2N

]

[MIM

][B

OB

]

[MIM

][T

fO]

[MIM

][M

eSO

3]

[MIM

][M

eSO

4]

[MIM

][E

tSO

4]

[MIM

][C

H3O

C2H

4S

O4]

[MIM

][C

l]

[MIM

][D

EP

O4]

[MIM

][D

MP

O4]

[MIM

][M

DE

GS

O4]

[MIM

][N

O3]

[MIM

][O

cSO

4]

[MIM

][P

F6]

[MIM

][S

bF

6]

[MIM

][S

CN

]

[MIM

][T

CB

]

[MIM

][T

FA

]

[MIM

][T

OS

]

[MIM

][A

c]

[MP

IP][

Tf

2N

]

[MP

IP][

SC

N]

[MP

Y][

BF

4]

[MP

Y][

Tf2

N]

[MP

Y][

TfO

]

[MP

Y][

C2H

5O

C2H

4S

O4]

[MP

Y][

SC

N]

[MP

Y][

TO

S]

[MP

YR

][T

f2N

]

[MP

YR

][T

fO]

[MP

YR

][S

CN

]

[N-C

3O

HP

Y][

Tf2

N]

[N-C

3O

HP

Y][

FA

P]

[(O

CH

2)2

IM][

Tf2

N]

[OC

H2M

IM][

Tf2

N]

[Qu

in][

Tf2

N]

[R1R

2R

3R

4N

][T

f2N

]

[R1

R2

R3

R4

N][

TS

]

[R1

R2

R3

R4P

][B

F4]

[R1

R2R

3R

4P

][T

f2N

]

[R1

R2

R3

R4P

][M

eSO

4]

[R1

R2

R3R

4P

](C

8H

17)2

PO

2

[R1R

2R

3R

4P

][C

l]

[R1

R2

R3R

4P

][F

AP

]

[R1

R2R

3R

4S

][T

f2N

]

1 2 3 4 5 6 7 8 9 10

11

12

13

14

15

16

17

18

19

20

21

22

23

24

25

26

27

28

29

30

31

32

33

34

35

36

37

38

39

40

41

42

43

44

45

46

47

48

49

50

51

52

53

54

55

56

57

58

59

60

61

62

63

64

65

66

67

68

69

70

71

72

73

74

75

Figure 3. The most recent UNIFAC-Lei parameter matrix.

UNIFAC parameters formerly available in the literature;83-85,103-105

parameters published in our previous works;44-50

parameters have been submitted for publication;

parameters added by other researchers using the UNIFAC-Lei model;106

new parameters evaluated in this work; parameters are not available.

Page 37 of 49

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AIChE Journal


Figure 4. Solubility of CO2 (1) in pure [AMIM]+[Tf2N]

- (a), and pure MeOH (b).

Lines, predicted results by the UNIFAC-Lei model; scattered points, experimental data. (▲) 313.2 K; (●)

333.2 K; (■) 353.2 K.

0 1 2 3 4 5 6

0.0

0.1

0.2

0.3

0.4

0.5

0.6(a)

x1

P (MPa)

0 1 2 3 4 5 60.0

0.1

0.2

0.3

0.4

(b)

x1

P (MPa)

Page 38 of 49

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0 1 2 3 4 5 60.0

0.1

0.2

0.3

0.4

0.5

0.6(a)

x1

P (MPa)

0 1 2 3 4 5 60.0

0.1

0.2

0.3

0.4

0.5

(b)

x1

P (MPa)

Page 39 of 49

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AIChE Journal


Figure 5. Solubility of CO2 in pure [AMIM]+[Tf2N]

-, pure MeOH, and their binary

mixtures at 313.2 K (a), 333.2 K (b), and 353.2 K (c).

Lines, predicted results by the UNIFAC-Lei model; scattered points, experimental data. ( ) pure

[AMIM]+[Tf2N]

-; (■) 80 wt% [AMIM]

+[Tf2N]

- + 20 wt% MeOH; (●) 50 wt% [AMIM]

+[Tf2N]

- + 50 wt%

MeOH; (▲) 20 wt% [AMIM]+[Tf2N]

- + 80 wt% MeOH; (★) pure MeOH.

0 1 2 3 4 5 60.0

0.1

0.2

0.3

x1

P (MPa)

(c)

Page 40 of 49

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Figure 6. Henry’s constants of CO2 (1) in pure [AMIM]+[Tf2N]

-, pure MeOH, and their

binary mixtures.

Lines and unfilled scattered points, predicted results by the UNIFAC-Lei model; filled scattered points,

experimental data. (■) 313.2 K; (●) 333.2 K; (▲) 353.2 K.

Pu

re I

L

80

wt%

IL

+ 2

0

wt%

MeO

H

50

wt%

IL

+ 5

0

wt%

MeO

H

20

wt%

IL

+ 8

0

wt%

MeO

H

Pu

re M

eOH

0

5

10

15

20

25

30

35

40H

1 (

MP

a)

Page 41 of 49

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AIChE Journal


Figure 7. Indirect CO2 generated from electricity supply to each of the processes

(PIECO2, comprised of compression, cooling, and heating facilities).

243.2 K Pure

Methanol

243.2 K

Methanol+IL

(50-50 wt%)

243.2 K Pure

IL

265.7 K Pure

IL

288.2 K Pure

IL

0

500

1000

1500

2000

2500

Ind

irec

t C

O2 e

mis

sio

n (

km

ol.

h-1

) Compression

Heating

Cooling

PIECO2

Page 42 of 49

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Table 1. Group Interaction (αmn and αnm), Group Volume (Rk) and Surface Area (Qk)

Parameters Involved in the UNIFAC-Lei Model

Group m Group n

[MIM][Tf2N]

Group n

CH=CH2

Group n

CH2

Group n

CH3OH

Group n

CO2

[MIM][Tf2N] 0 -61.17 145.8 301.2 81.4726

CH=CH2 799.84 0 -35.36 787.6 1147.077*

CH2 400.89 86.02 0 697.2 28.05

CH3OH 4.39 -12.52 16.51 0 153.779*

CO2 14.5632 829.404* 672.8 43.002

* 0

Rk 8.3145 1.3454 0.6744 1.4311 1.3

Qk 7.392 1.176 0.54 1.432 1.12

*New group binary interaction parameters obtained in this work; others are collected from the

references.44,46,83-85

Page 43 of 49

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AIChE Journal


Table 2. Operating Conditions and Design Specifications of Five CO2 Capture Processes

Items

243.2 K

Pure

Methanol

243.2 K

Methanol +

IL (50-50

wt%)

243.2 K

Pure IL

265.7 K

Pure IL

288.2 K

Pure IL

Flow rates (kg·s-1

)

Raw gas 105 105 105 105 105

Solvent 395 395 395 395 790

Absorption tower TW1

Temperature (K) 243.2 243.2 243.2 265.7 288.2

Number of stages 10 (5+5) 10 (5+5) 20 (10+10) 22 (11+11) 24 (12+12)

Pressure (MPa) 6 6 6 6 6

Rectification column TW2

Number of stages 4 4 6 4 4

Pressure (MPa) 3 3 2 5 5

Boilup ratio 0.6 0.6 0.6 0.6 0.6

Flash drum TW3

Temperature (K) 253.2 253.2 293.2 293.2 293.2

Pressure (MPa) 0.08 0.08 0.01 0.01 0.01

Page 44 of 49

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AIChE Journal


Table 3. Additional UNIFAC-Lei Group Interaction Parameters (αmn and αnm) Used in

the Process Simulation

Group m Group n αmn αnm

[MIM][Tf2N] H2 998.625 726.156

[MIM][Tf2N] CO 345.409 188.311

CH2 H2 496.446 42.082

CH2 CO 140.162 42.822

CH3OH H2 581.623* 160.548

*

CH3OH CO 127.430* 190.104

*

*New group binary interaction parameters obtained in this work; others are from our previous work.

47

Page 45 of 49

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Table 4. Details of the Purified Gas and Captured CO2 Outflow Streams in Each of the

Five CO2 Capture Processes

Streams description

243.2 K

Pure

Methanol

243.2 K

Methanol +

IL (50-50

wt%)

243.2 K

Pure IL

265.7 K

Pure IL

288.2 K

Pure IL

Purified gas stream

Temperature (K) 246.7 245.9 244.3 273.1 288.9

Mole flow (kmol.h-1

) 11913.4 11910 11759.7 12398.6 11757

Mole fraction

H2 0.702 0.703 0.712 0.675 0.712

CO 0.284 0.284 0.288 0.272 0.286

CO2 0.014 0.013 851 PPM 0.053 0.002

CH3OH 145 PPM 130 PPM 0 0 0

[AMIM]+[Tf2N]

- 0 trace trace trace trace

CO2 gas stream

Mole flow (kmol.h-1

) 5963.45 5962.35 6042.96 5404.29 6045.70

Mole fraction

H2 0 0 0 0 0

CO 35 PPM 167 PPM 16 PPM 0.002 0.003

CO2 0.988 0.988 1 0.998 0.997

CH3OH 0.012 0.011 0 0 0

[AMIM]+[Tf2N]

- 0 trace 40 PPB 40 PPB 40 PPB

HE3 flow rate (kg.s-1

) 434.70 419.45 409.41 431.79 794.17

Page 46 of 49

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Table 5. Power Penalty Comprised of Blocks Energy Demand for Each of the Five CO2

Capture Processes

Block energy demand

243.2 K

Pure

Methanol

243.2 K

Methanol +

IL (50-50

wt%)

243.2 K

Pure IL

265.7 K

Pure IL

288.2 K

Pure IL

Total heat duty of HEs (GJ·h-1

) 321.95 160.15 205.07 119.29 81.47

TW2 reboiler heat duty (GJ·h-1

) 1009.60 564.11 111.75 377.93 720.07

TW3 heat duty(GJ·h-1

) 749.07 437.69 +67.82 269.09 650.10

CM1 net work required (MW) 8.72 1.91 1.78 0.07 0.07

PM1 required electricity (MW) 3.29 2.33 1.86 1.86 3.68

Page 47 of 49

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Table 6. Separation Performance of the Five CO2 Capture Processes

Flow rates

243.2 K

Pure

Methanol

243.2 K

Methanol +

IL (50-50

wt%)

243.2 K

Pure IL

265.7 K

Pure IL

288.2 K

Pure IL

Fuel gas production (kmol·h-1

) 11749.65 11748.81 11749.68 11741.17 11733.95

Captured CO2 (kmol·h-1

) 5890.60 5892.95 6042.86 5395.65 6029.88

Vented CO2 (kg·h-1

) 7132.30 7026.08 440.21 28932.50 1013.90

Solvent loss (kg·h-1

) 2382.99 2241.29 0.10 0.09 0.10

Page 48 of 49

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AIChE Journal


Table 7. Block Specific Energy Penalties (BSEPs) and Process Specific Energy Penalty

(PSEP) for the Five CO2 Capture Processes

Items

243.2 K

Pure

Methanol

243.2 K

Methanol +

IL (50-50

wt%)

243.2 K

Pure IL

265.7 K

Pure IL

288.2 K

Pure IL

BSEPs

Heat exchangers 0.055 0.027 0.034 0.022 0.014

TW2 reboiler 0.171 0.096 0.018 0.070 0.119

TW3 flash drum 0.127 0.074 0 0.050 0.108

PSEP (

2

GJ energy demand

kmol captured CO) 0.353 0.197 0.052 0.142 0.241

Page 49 of 49

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AIChE Journal


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