Critical Assessment of Using an Ionic Liquid as Entrainer via Extractive

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Critical Assessment of Using Ionic Liquid as Entrainer via Extractive Distillation Hung-Hsin Chen, Meng-Kai Chen, Bor-Chang Chen, and I-Lung Chien Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.7b01223 • Publication Date (Web): 21 Jun 2017 Downloaded from http://pubs.acs.org on June 22, 2017

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Industrial & Engineering Chemistry Research

Critical Assessment of Using Ionic Liquid as Entrainer via Extractive Distillation

Hung-Hsin Chen, Meng-Kai Chen, Bor-Chang Chen, and I-Lung Chien*

Department of Chemical Engineering National Taiwan University Taipei 10617, Taiwan

ABSTRACT Ionic liquid (IL) has received much attention in the last two decades. One of the important applications is to use this relatively new class of compounds for the separation of azeotropic mixtures via extractive distillation. In this paper, overall extractive distillation processes of two azeotropic separation systems using favorable ionic liquid as entrainer are rigorously developed. The optimized design flowsheets are compared with the conventional processes using industrial entrainer. The two ionic liquid extractive distillation systems include separating acetone and methanol using 1,3-dimethylimidazolium dimethylphosphate ([MMIM][DMP]) as entrainer and *

Corresponding author. I-Lung Chien, Tel: +886-3-3366-3063; Fax: +886-2-2362-3040; E-mail: [email protected]

1 Environment ACS Paragon Plus


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another

system

of

separating

isopropanol

(IPA)

and

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water

using

1-ethyl-3-methylimidazolium acetate ([EMIM][OAC]) as entrainer. The potential drawbacks of using ionic liquid in the extractive distillation systems are given in the paper. It is found that degradation temperature and high normal boiling point temperature of the ionic liquids in these two systems make the IL processes requiring highly vacuum operating condition, thus, are only economically comparable to the conventional processes.

Keywords: Ionic Liquid, Azeotrope, Extractive Distillation, Entrainer, Process Design.


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1. Introduction Ionic liquid (IL) is composed with a cation and an anion and has melting point lower than 100 °C. Ionic liquids have been found for more than 100 years1, 2, but, until the last two decades, they started to get attention in many applications such as in azeotropic separation3. In 1992, Wilkes and Zaworortko4 found ILs composed with imidazolium-based cations are very stable in water and air. In 2006, Earle et al.5 investigated the volatility of ILs and found they can be separated under low pressure condition. In 2014, Cao and Mu6 studied the thermal stability of 66 ILs by thermogravimetric analysis and divided them into five groups according to their thermal stability. In 2015, Heym et al.7 studied liquid vapor pressure and reaction mechanism of thermal decomposition of 9 ILs. Besides the studies above on the thermodynamic properties of ILs, quite a few researchers proposed the extractive distillation systems using ILs as entrainer8-12 and the extraction process using ILs as solvent13-17. Pereiro et al.3 gave a thorough review about ionic liquids in separation of azeotropic systems. In 2016, Aniya et al.12 compared the process economics of using conventional organic entrainer, triethylene glycol (TEG), and IL, 1-ethyl-3-methy-limidazolium chloride ([EMIM][Cl]), for dehydration of tert-butyl alcohol via extractive distillation. The result showed that using [EMIM][Cl] as entrainer reduced 13.9% of total annual cost (TAC) compared with using TEG as entrainer. However, this study did not consider the thermal stability of ionic liquid. The operating pressure in their entrainer recovery column was 0.95 bar. Owing to low vapor pressure of IL, such a high operating pressure led to the bottom temperature in entrainer recovery column near to the normal boiling point of IL, which is higher than the degradation temperature of IL.



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For this ionic liquid, the normal boiling point and the degradation temperature of [EMIM][Cl] was found to be 239.12 oC18 and 196 oC19, respectively. To perform more practical simulation and economic analysis, vacuum system should be designed for the entrainer recovery section to avoid the operating temperature becoming higher than the degradation temperature. Jongmans et al.20 studied 5 entrainer recovery configurations to separate ionic liquid [4-mebupy][BF4] from product styrene in ethylbenzene/styrene extractive distillation system. It was found that only one flash drum under mild vacuum is not enough to recover all styrene, and it required another unit operation after the first flash drum to have complete separation. Nitrogen stripping, ethylbenzene stripping, super critical carbon dioxide extraction and another flash drum at very low pressure were investigated.

Another

configuration

was

to

apply

mixed

entrainer,

N-methylpyrrolidinone/IL and sulfolane/IL, in the extractive distillation system to recover entrainer by another distillation column. From their economic evaluation, stripping with ethylbenzene after a flash drum and two vacuum flash drums in series are the most promising IL recovery configurations. In this paper, two azeotropic separation systems via extractive distillation are studied to compare the economic performances of using conventional entrainer to that of using ionic liquid as entrainer. The first system is to separate acetone from methanol. Both components are common industrial solvents and their mixture appears in the water phase product of Fischer-Tropseh synthesis.21, 22 Kossack et al.23 proposed an entrainer screening methodology in extractive distillation and found three better candidates in acetone/methanol system: water, DMSO (dimethyl sulfoxide), chlorobenzene. Luyben24 compared the performance of acetone/methanol system with


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these solvents. DMSO showed the best result in both steady-state economic analysis and dynamic controllability. Chen et al.25 measured the isobaric vapor-liquid equilibrium data for acetone/methanol/phosphate ionic liquids. Among of three ionic liquids discussed, 1,3-dimethylimidazolium dimethylphosphate ([MMIM][DMP]), 1-ethyl-3-methylimidazolium

diethylphosphate

([EMIM][DEP]),

and

1-butyl-3-methylimidazolium dibutylphosphate ([BMIM][DBP]), the relative volatility between acetone and methanol was enhanced the most by adding [MMIM][DMP] as entrainer. The ability of relative volatility enhancement of this IL is also much better than DMSO. The second system discussed in this paper is dehydration of isopropanol (IPA). IPA is widely used in the semiconductor industry as a cleaning agent, thus the recovery of IPA from the waste solvent stream, composed of water and IPA, is an important task. Two usual practices are used in industry to separate IPA from water: (1) adding a light component cyclohexane as entrainer via heterogeneous azeotropic distillation26, 27; (2) adding a heavy component DMSO as entrainer via extractive distillation.28 Arifin and Chien28 studied these two processes and found that the latter saved 32.7% TAC. Liang et al.29 combined preconcentration column and entrainer recovery column in extractive distillation using DMSO as entrainer. It is noteworthy that the feed and product purity specifications in Liang et al.’s work were different from the ones in Arifin and Chien’s work. Thus, two sets of specifications will be discussed in this work. Zhang et al.30 measured the isobaric vapor-liquid equilibrium for IPA/water/ILs. Among of six ILs discussed, composed of an anion chosen from acetate ([OAC]-) or chloride ([Cl]-), and a

cation

chosen

from

1-ethyl-3-methylimidazolium

([EMIM]+)

or

1-butyl-3-methylimidazolium ([BMIM]+) or 1-hexyl-3-methylimidazolium ([HMIM]+),



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the relative volatility between IPA and water was enhanced the most by adding [EMIM][OAC] as entrainer, even much better than the conventional entrainer DMSO. In the following section, the required thermodynamics models and parameters are determined before simulation. The optimized process flowsheet and economic performances

of

two

extractive

distillation

systems

are

discussed.

For

acetone/methanol system, using DMSO or [MMIM][DMP] as entrainer are compared. As for IPA dehydration system, using DMSO or [EMIM][OAC] as entrainer are compared considering two sets of feed and product purity specifications. The objective of this paper is to determine whether or not the usage of IL as entrainer in extractive distillation would be more economically favorable than using conventional entrainer, considering all the necessary costs including the costs associated with vacuum operating condition.

2. Thermodynamic models and process simulation issues In this work, the simulation software, Aspen Plus V8.4, is used. The nonrandom two-liquid (NRTL) model is used to describe the non-ideal behavior in liquid phase. The behavior in gas phase is assumed to be ideal (IG). Many researchers adopted NRTL-IG model to describe the non-ideal behavior of the mixture containing IL and the accuracy were quite good.25, 30 2.1. Ionic liquids Due to its complex and uncommon structure, no component data of IL is built in the Aspen databank. Proper definition of new component in the simulation software is essential. The necessary thermodynamic properties for NRTL-IG model in Aspen Plus include: critical properties, ideal gas heat capacity, liquid vapor pressure, and heat of


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vaporization. Valderrama et al.31, 32 estimated critical properties of ionic liquids by the group contribution method. Ge et al.33 estimated ideal gas heat capacity by the extended Joback method. Heym et al.7 measured liquid vapor pressure of nine ionic liquids with imidazolium type cation, one of them was [MMIM][DMP], and then regressed into simplified Antoine equation. Clausius-Clapeyron equation is applied to predict heat of vaporization with liquid vapor pressure data, and then the model based on Watson equation can be built. The corresponding thermodynamic models built in Aspen Plus are CPIG for ideal gas heat capacity, PLXANT for liquid vapor pressure, and DHVLWT for heat of vaporization. Unfortunately, there is no experimental data of liquid vapor pressure of [EMIM][OAC] in open literature. In Aspen Plus, Riedel equation can be applied to predict liquid vapor pressure of unknown compound based on its normal boiling point, critical temperature and critical pressure. From the experimental liquid vapor pressure data in Heym et al.’s work7, normal boiling points of nine ionic liquids discussed could be derived based on the simplified Antoine equation, and the range of normal boiling points is from 834 to 1075 K. In this paper, three different values of normal boiling point of [EMIM][OAC], 800, 900 and 1000 K, are assumed to study the sensitivity of this assumption. Besides normal boiling point, critical temperature is also needed for NRTL-IG model in Aspen Plus. Valderrama et al.31, 32 used the group contribution method to estimate the critical properties of ionic liquids. The estimated critical temperature of [EMIM][OAC] by the group contribution method is 807.10 K, which is lower than the normal boiling point we assumed if assumed at 900 or 1000 K. Figure S1a, S1c and S1e (shown in Supporting Information) shows the predicted PLXANT models of



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[EMIM][OAC] by extended Antoine equation based on the critical temperature from the group contribution method. The predicted vapor pressures for all three assumed normal boiling points could not describe the low volatility of ionic liquids. To solve this problem, another method to predict critical temperature is necessary. In Aspen Plus, Klincewicz method can be applied to predict critical temperature with information of molecular weight and normal boiling point. Figure S1b, S1d and S1f shows the predicted PLXANT models of [EMIM][OAC] by extended Antoine equation based on the estimated critical temperature from Klincewicz method. Result shows that application of Klincewicz method are better to describe the very low liquid vapor pressure of this ionic liquid, which makes the using of Riedel equation and Klincewicz method seem quite reasonable. Note that in the PLXANT model, the range of valid temperature needs to be reset to include all operating temperatures in the following design study avoid extrapolation. All necessary thermodynamic properties and model parameters of [MMIM][DMP] and [EMIM][OAC] are calculated by the above-mentioned methods and summarized in Table 1. Considering the thermal stability of ionic liquid, the operating temperature should be limited to a maximum value. In Heym et al.’s work7, the maximum operating temperature for 1% conversion by decomposition of [MMIM][DMP] is 105.85 oC (379 K). Therefore, the maximum operating temperature of the process involving [MMIM][DMP] is conservatively set as 100.85 oC. Cau and Mu6 indicated that [EMIM][OAC] started decomposition at 140 oC, which is set as the maximum operating temperature of the process involving [EMIM][OAC] in this paper. 2.2. Acetone/methanol/[MMIM][DMP] system NRTL-IG model is used because ionic liquid is involved in this system. We use the


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vapor-liquid equilibrium data of acetone/methanol/[MMIM][DMP] system measured by Chen et al.25 to determine the binary parameters of NRTL model, as listed in Table 2a. Comparison of the NRTL model prediction and experimental data can be found in Figure 1. 2.3. IPA/water/[EMIM][OAC] system Zhang

et

al.30

measured

the

vapor-liquid

equilibrium

data

of

IPA/water/[EMIM][OAC] system and regressed the binary parameters of NRTL model. However, the binary parameters of IPA/water pair from Zhang et al.’s work could not predict the binary vapor-liquid behavior well. In Zhang et al.’s work, they made a mistake in retrieving the binary parameters of IPA/water from Marzal et al.’s work34, which could fit the binary experimental data34 well. Therefore, we did the regression again to find the binary parameters of the other two pairs with the experimental data from Zhang et al.’s work and use the correct binary parameters of IPA/water from Marzal et al.’s work, as shown in Table 2b. The prediction of vapor-liquid behavior on the basis of new NRTL parameter set and the comparison with experimental data can be found in Figure 2. 2.4. Process simulation issues In Aspen Plus, several modules are used in the following process simulation: Flash2 for flash drum, Radfrac for distillation column and stripping column, Heater and HX for heat exchanger. For vacuum system, detailed description can be found in Seider et al.’s book.36 Air would leak into the equipment operating under vacuum condition via gasket joints, porous welds, cracks and fissures in vessel walls. Then the air leakage leaves the equipment accompanied by volatile process components. To recover most of those



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volatile components and reduce the loading on vacuum system, the exiting gas from equipment should pass through a pre-condenser before entering vacuum system. Flow rate of air leakage could be estimated by the following equation36:

[

]

W = 5 + 0.0298 + 0.03088(ln P ) − 0.0005733(ln P ) V 0.66 2

(1)

where W is the air leakage rate in lb/hr, P is the absolute operating pressure in torr assuming a barometric pressure of 760 torr, and V is the vessel volume in ft3. The recovery of volatile component in pre-condenser could be simulated by a Flash2 module, as displayed in Figure 3. Steam-jet ejector, liquid ring pump and screw compressor are three common types of vacuum system in industrial practice. Lower limit of suction pressure for these three types of vacuum system are 2, 10, 0.1 torr (267, 1333 and 13 Pa), respectively35. In fact, suction pressure much lower than 2 torr can be obtained by multi-stage ejectors up to six-stage, however, cost multiplying factors for only one-stage to three-stage steam-jet ejector are listed in Seider et al’s book36. Therefore, maximum number of stages of steam-jet ejector is limited to three in our study. For the following sections, type of vacuum system is determined by the operating pressure and minimum annual cost. 2.5. Economic analysis The TAC includes the annualized capital cost and the operating cost with payback period for the capital cost assuming to be 3 years. In this work, TAC calculations are mainly based on the formulas from Luyben’s book.37 For the vacuum system, the capital cost and operating cost are based on the formulas from Seider et al.’s book.36 Detailed information about economic analysis is listed on Table S1 in Supporting Information.


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3. Case study: Acetone and methanol separation 3.1. DMSO as entrainer Figure S2 shows the retrofit of the simulation result from Luyben’s work24 with an additional feed-effluent heat exchanger for additional heat integration. 540 kmol/hr 50 mol% acetone/50 mol% methanol mixture and 750 kmol/hr DMSO are fed into an extractive distillation column with rectifying section above the entrainer feed, extractive section between the two feeds and the stripping section below the feed stage of the acetone/methanol mixture. DMSO enhances the relative volatility between acetone and methanol, making acetone evaporating more easily and further purified via rectifying section. The bottom stream of extractive distillation column, composed of DMSO and methanol, is further sent to the second distillation column to separate entrainer and methanol. Methanol is collected in the distillate, and DMSO in bottom stream is recycled to heat up fresh acetone/methanol mixture before entering the extractive distillation column. To compensate for DMSO loss accompanied with two product streams, small amount of makeup stream is added. Doherty and Malone39 gave a guideline of the favorable entrainer temperature before entering the extractive distillation column to be 5-10 oC lower than the temperature of distillate. Thus, the recover entrainer after the feed-effluent heat exchanger is further cooled to 46.85 oC (320 K) by cooling water. In this conventional design flowsheet, both product purity specifications are 99.95 mol%, and both distillation columns are operated at atmospheric pressure. 3.2. Ionic liquid as entrainer [MMIM][DMP] enhances the relative volatility between acetone and methanol even more than the conventional DMSO entrainer, thus, one would intuitively think to



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make the separation easier. Both IL and acetone/methanol mixture are fed to an extractive distillation column to purified acetone to 99.95 mol% first. To make sure that acetone would not leave with methanol too much at bottoms of extractive distillation column, a 0.01 mol% of acetone at bottom stream is also specified as in the conventional design flowsheet. Entrainer is recovered from the bottom stream of extractive distillation column through two flash drums in series operating in vacuum. First flash drum operates under a mild low pressure (76 kPa) with steam-jet ejector and at temperature lower than the maximum operating temperature (100.85 oC). Vapor from the first flash drum is sent to a pre-condenser and cooled down to 46.85 oC (320 K) by cooling water. Liquid from the first flash drum is sent to the second flash drum operating at very low pressure (40 Pa) with screw compressor and at temperature not exceeding 100.85

o

C. Vapor from the second flash drum is sent to another

pre-condenser and cooled down to -57.78 oC by refrigerant at -67.78 oC (-90 oF refrigent36). Condensates from two flash drums consist of at least 99.95 mol% methanol. Recovered IL is sent to a feed-effluent heat exchanger and further cooled down to 46.85 oC by cooling water. Methanol residue in the IL recycle stream would reduce the separation ability in extractive distillation column. Maximum methanol impurity mole fraction (Xmmax) in IL recycle stream to purify acetone in extractive distillation column to 99.95 mol% may be influenced by entrainer-feed ratio (E/F), total number of stages of extractive distillation column (NT), entrainer feed stage (NE) and acteone/methanol mixture feed stage (NF). Sensitivity analysis for these four variables is shown in Figure S5. Both NT and E/F have no obvious effect on Xmmax. Tolerance of methanol impurity increases as stage number of extractive section (NF-NE) increases and then stays almost constant


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at a larger total stage of extractive section. Also larger stage number of the rectifying section (e.g. NE) actually makes the tolerance of methanol impurity to be decreased. This may seem counterintuitive. However, in the rectifying section with less entrainer concentration, the methanol tends to go up with the minimum-boiling azeotrope. Methanol impurity amount is mainly determined by the operating condition of the second flash drum. High operating temperature and low operating pressure would lead to purer [MMIM][DMP] in IL recycle stream. Operating temperature is set at the upper limit, 100.85 oC, and operating pressure is determined by optimization based on lower TAC. Operating condition of the first flash drum would affect methanol loss through vapor from the first pre-condense and also affect loading of the second flash drum. Operating pressure is determined by 0.1% loss of methanol, and operating temperature should be maximized (100.85 oC) to reduce loading of the second flash drum. The resulting process flowsheet is shown in Figure 4a. From Figure 4a, we can see that [MMIM][DMP] is purified merely from 23.0 to 33.5 mol% in the first flash drum, and not enough methanol is vaporized from the first flash drum. If chilled water can be used as the cooling agent of the first pre-condenser, the operating pressure of the first flash drum could be much lower to vaporize more methanol. If chilled water is used to cool down vapor from the first flash drum to 25 oC, [MMIM][DMP] could be purified from 23.0 to 47.4 mol%, as shown in Figure 4b. Table 3 lists detailed capital cost and operating cost of these two cases. Although chilled water (4.43 $/GJ) is much more expensive than cooling water (0.345 $/GJ), loading of the second flash drum is greatly reduced and use much less refrigerant at the second pre-condenser. From Table 3, TAC can be reduced by 7.92% with using of chilled water in the first pre-condenser; therefore, optimization is then applied to this



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process flowsheet as in Figure 4b. Sequential iterative optimization to minimize TAC is applied, and the procedure is outlined in Figure 5. There are six design variables in this process. However, NT is excluded in the optimization for simplification. It is fixed at 37, the same as in the conventional design flowsheet. The remaining five design variables are E/F, operating pressure of two flash drums (P1 and P2), locations of entrainer and feed (NE and NF). We combine costs of extractive distillation column, feed-effluent heat exchanger and cooler as TAC1, and combine costs of two flash drums, two pre-condensers, steam-jet ejector for the first flash drum and screw compressor for the second flash drum (vacuum system are not shown on the flowsheet) as TAC2. The former represents total annual cost of extractive distillation section, and the latter represents the total annual cost of entrainer recovery section. NE and NF have no effect on TAC2, P1 has no effect on TAC1, but E/F and P2 affect both TAC1 and TAC2. Larger entrainer amount will increase loading of entrainer recovery section, but may make the relative volatility of acetone enhanced more. P2 affects methanol impurity amount in IL recycle stream directly and also contributes to cost of vacuum system. As P2 is larger than 44 Pa, the simulation would not converge. From optimization result shown in Figure 6a and 6b, when E/F is 0.15, P2 is 44 Pa, NE is 2, NF is 11 and P1 is 22 kPa would lead to lowest TAC. Figure 7 shows the optimal process flowsheet of this system. 3.3. Discussion From thermal degradation kinetics of [MMIM][DMP] in Heym et al.’s work7, cost of IL loss could be evaluated. In the optimal case, IL would loss 16.8 kg per year though thermal degradation, and IL lost cost is about 2610 $/yr, which is much lower than other equipment capital cost. Therefore, IL loss cost is not considered in TAC


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calculation. Due to extremely low vapor pressure of IL, it is not necessary to add an entrainer makeup stream like DMSO process. However, methanol would loss with vapor stream from flash drums when vacuum system is applied. Both of these additional costs should not be ignored and are considered as operating cost in TAC calculation. Table 4 lists detailed terms of TAC calculation of both DMSO process and optimal IL process. In extractive distillation section, IL process saves significant 38.69% operating cost and 37.37% capital cost because [MMIM][DMP] enhance relative volatility of acetone more than DMSO does. However, in entrainer recovery section, IL process leads to 10.76% more operating cost and 123.14% more capital cost. Although it seems that methanol/[MMIM][DMP] is easier to separate than methanol/DMSO, additional chilled water and expensive refrigerant makes operating cost higher. To operate at extreme low pressure condition (44 Pa), screw compressor should be applied, which accounts for over half (53.58%) of capital cost in the entrainer recovery section. To sum up, application of IL in acetone/methanol separation via extractive distillation is still more economical with saving of 10.07% TAC than DMSO process, even though expensive screw compressor and refrigerant are applied. However, the saving is not significant as first thought. Using acetone stripping after a flash drum to recover [MMIM][DMP] is not considered, because IL is purified merely from 23.0 to 49.5 mol% by the first flash drum, which leads the required amount of stripping agent too large for stripping out the remaining methanol.

4. Case study: IPA dehydration 4.1. DMSO as entrainer



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In Arifin and Chien’s work28, product specifications are 99.9999 mol% of IPA and 99.9 mol% of water, and feed is 100 kmol/hr with 50 mol% IPA/50 mol% water at 25 o

C. In Liang et al’s work29, product specifications are 99.99 mol% of IPA and 99.99

mol% of water and feed is 76.8965 kmol/hr with 65 mol% IPA/35 mol% water at 80.18 oC. Figure S3 duplicates Arifin and Chien’s simulation result, and Figure S4 retrofits Liang et al’s work with a feed-effluent heat exchanger added for heat integration. DMSO and IPA/water mixture are fed to an extractive distillation column to purify IPA in the distillate. The bottom stream of extractive distillation column, composed of DMSO and water, is further sent to the second distillation column to separate DMSO and water. Water is collected in the distillate, and DMSO at bottom stream is recycled to heat up fresh IPA/water mixture before entering the extractive distillation column. To compensate DMSO loss through two product streams, small amount of makeup stream is added. 4.2. Ionic liquid as entrainer [EMIM][OAC] enhances more of the relative volatility between IPA and water, making separation even easier in the extractive distillation column. Both IL and IPA/water mixture are fed to an extractive distillation column to purified IPA first. Ionic liquid is also recovered from the bottom stream of extractive distillation column through two flash drums in series operating in vacuum. Recovered IL is sent to a feed-effluent heat exchanger to heat up IP/water mixture to save energy. 4.2.1. 99.9999 mol% IPA product purity To make sure that IPA would not leave with water too much so that making the water product specification of 99.9 mol% not reachable, ratio of IPA to (IPA+water) at


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bottom stream of extractive distillation column is specified as 0.09 mol%. Also, water impurity in the IL recycle stream would reduce the separation capability in extractive distillation column. Maximum mole fraction of water impurity mole in IL recycle stream (Xwmax) to purify IPA in extractive distillation column to 99.9999 mol% may be influenced by E/F, NT, NE and NF. Sensitivity analysis for these four variables can be seen in Figure S6. Similar trends as in the acetone/methanol system were observed. Water impurity amount is mainly determined by the operating condition of the second flash drum. High operating temperature and low operating pressure would lead to purer [EMIM][OAC] in IL recycle stream. Operating temperature is set as the upper limit, 140 oC. Operating pressure is specified to obtain 5×10-4 mol% water in IL recycle stream, which is a conservatively low value of Xwmax shown in Figure S6 so that stringent IPA product purity can be met. The specifying of Xwmax to a conservative value simplifies the latter iterations of optimization scheme by reducing one design variable (operating pressure of the second flash drum) without affecting much of the optimization result. It is noted that the stringent dryness of the ionic liquid (5×10-4 mol% water) may be hard to achieve in practice. Figure 8a shows a base case condition of the IL separation system. Owing to the very low operating pressure of the second flash drum, some amounts of [EMIM][OAC] vaporize and recovered by the following pre-condenser. The condensate, composed of 0.07 mol% [EMIM][OAC], from the second pre-condenser is sent back to the bottom of extractive distillation column to prevent loss of IL. Operating condition of the first flash drum would affect water loss through vapor from the first pre-condenser and also affect the loading of the second vacuum system. For simplicity of latter optimization of this separation system, operating pressure is determined by 1% loss of water and



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operating temperature should be at 140 oC to reduce loading of the second vacuum system. Vapor from the first flash drum is cooled down to 25 oC by chilled water, as discussed in Section 3, and vapor from the second flash drum is cooled down to -91.11 o

C by refrigerant at -101.11 oC (-150 oF refrigent36). However, the lowest suction pressure of screw compressor is merely 13 Pa, which

is mentioned in Seider et al’s book36. Therefore, two-flash-drum entrainer recovery system proposed here seems impractical because the vacuum pressure of the second flash drum is lower than the limit. Figure 8a shows that [EMIM][OAC] is purified to 94.5 mol% by the first flash drum, and the majority of water is already vaporized in this flash drum. Thus, another low pressure stripping column can be applied to replace the second flash drum. Stripping agent for this system is IPA vapor drawn from the second stage of extractive distillation column, and this operation is similar to ethylbenzene stripping in Jongmans et al.’s work.20 Figure 8b shows the process flowsheet with a flash drum followed by a stripping column to recover IL. The IPA stripping vapor is heated to 244 oC by high pressure steam before entering the stripping column. Then the IPA stripping vapor contacted with liquid stream from the first flash drum counter-currently in the stripping column. Vapor from the stripping column is sent to the second pre-condenser and cooled down to -40 oC by refrigerant. The stripped [EMIM][OAC] from the stripping column is sent to feed-effluent heat exchanger before entering extractive distillation column. IPA impurity in the IL recycle stream does not have a negative impact on extractive distillation performance, because IL recycle stream is sent to the top of extractive distillation column and IPA would vaporize to distillate rather than flowing down to extractive section. Both stripping IPA vapor flow rate and operating pressure in


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stripping column affect the water impurity in IL recycle stream. Stripping IPA vapor flow rate is determined by specifying 5×10-4 mol% of water in IL recovery stream, which is the same specification as in the two-flash-drum configuration. Operating pressure in stripping column is determined by 0.025% loss of IPA from the vapor stream of second pre-condenser. This leads to the higher operating pressure of stripping column. The second pre-condenser recovers the majority of IPA from the vapor of stripping column, and the condensate is sent back to extractive distillation column at the stage with similar composition. Using stripping column instead of the second flash drum makes operating pressure within the operating range of screw compressor. 4.2.2. 99.99 mol% IPA product purity For this case, only operating conditions and specifications are different from the 99.9999 mol% IPA case. To make sure that IPA would not leave with water too much so that making the purer water product specification of 99.99 mol% not reachable, ratio of IPA to (IPA+water) in bottom stream of extractive distillation column is specified as 0.009 mol%. Sensitivity analysis for E/F, NT, NE and NF with Xwmax is shown in Figure S7. Tendency of each variable is similar to the case of 99.9999 mol% IPA. Two-flash-drum configuration is considered first, and process flowsheet is shown in Figure 9a. For the second flash drum, operating pressure is specified to obtain 0.05 mol% water impurity in IL recycle stream, which is a conservatively low value of Xwmax shown in Figure S7 so that IPA product purity can be met. The second pre-condenser is cooled down to -40 oC by refrigerant. The resulting operating pressure of the second flash drum is within the operating range of three-stage steam-jet ejector. With this higher operating pressure (32 Pa) as compared to Figure 8a (0.32 Pa), the ionic liquid in V2 stream is



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Page 20 of 53

negligible, thus, the liquid stream after the second pre-condenser can be drawn out as another water product stream. Process flowsheet using one-flash-drum-one-stripping-column configuration is shown in Figure 9b. The stripping IPA vapor flow rate is specified to obtain 0.05 mol% water in IL recycle stream, and operating pressure of stripping column is determined by 0.01% loss of IPA. The effect of changing this value of loss of IPA to the total TAC is shown in Figure 10 which evidences the selection of 0.01% loss of IPA. 4.3. Discussion Process flowsheets shown in Figure 8b, 9a and 9b are all optimized, followed the similar optimization procedure discussed in Section 3. Detailed terms in TAC calculation for 99.9999 mol% and 99.99 mol% IPA cases are listed in Table 5a and 5b, respectively. Cost calculation did not include 99.9999 mol% IPA case using two-flash-drum configuration, because this case is impractical with too extreme lower pressure at second flash drum. Comparing with acetone/methanol system in Section 3, the remaining IL composition in the liquid phase after the first flash drum is much larger (0.945 versus 0.495 for the acetone/methanol system). It is mainly determined by the maximum allowable operating temperature of IL, [MMIM][DMP] is 100.85

o

C while

[EMIM][OAC] is 140 oC. The higher the operating temperature, the better separation performance could be obtained at the first flash drum. For one-flash-drum-one-stripping-column configuration, some IPA would loss as vapor from the second pre-condenser. Also, another recycle stream is needed to send back the liquid stream from second pre-condenser to extractive distillation column to


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prevent large product loss. However, by introducing stripping column into the process, the second vacuum system could operate at a much more moderate vacuum condition. For processes which have to use screw compressor, two-flash-drum configuration for 99.99 mol% IPA and one-flash-drum-one-stripping-column configuration for 99.9999 mol% IPA, capital cost increases obviously, even though better entrainer (IL) would help saving cost in extractive distillation section. Only the configuration of one-flash-drum-one-stripping-column for 99.99 mol% IPA could use cheaper three-stage steam-jet ejector with operating pressure higher than 267 Pa. In this case, a small saving of 8.36% TAC can be obtained in comparison with the conventional system. It is important to note that we do not have cost multiplying factor for steam-jet ejectors with more than three stages from Seider et al’s book36. If cost of those steam-jet ejectors with more stages could be evaluated, using of screw compressor could be avoided so that some savings of the capital cost can be made. So far, all the process flowsheets were simulated under the assumption that the normal boiling point of [EMIM][OAC] is 1000K. Figure 11a, 11b, and 11c show the result of sensitivity analysis of different normal boiling point assumptions at 800K, 900K, or 1000K. For one-flash-drum-one-stripping-column configuration, the values of TAC, TAC1 and TAC2 are all very close despite different reasonable boiling point assumptions. The maximum TAC differences are all smaller than 1%, thus, the effect of different boiling point assumptions can be negligible. As for two-flash-drum configuration, the results in Figure 11 display that both TAC1 and TAC2 for assumption at 900K are all very close to the ones based on 1000K assumption. However, TAC2, which is the total annual cost of entrainer recovery section, shows obvious increase when the normal boiling point assumption is at 800K.



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From previous Figure S1b, the liquid vapor pressure evaluated at 800K assumption could not be ignored under highly vacuum condition in the entrainer recovery section. The amount of IL vaporizing from the highly vacuum second flash drum is more than the ones based on the other two assumptions. Thus, the ionic liquid in the liquid stream of the second pre-condenser cannot be ignored. This will also make the water purity of this liquid stream not satisfying product specification. An extra recycle stream is needed to send back this stream to extractive distillation column. Since the total flow rate of the treated mixture in the entrainer recovery section becomes larger, TAC2 is increased with 800K assumption.

5. Conclusions Two extractive distillation systems using IL as entrainer and two kinds of entrainer recovery configurations are discussed in this paper. In our study, practical issue of thermal stability of IL and rigorous calculation of costs associated with necessary vacuum system are considered. For both separation systems, the costs associated with the extractive distillation column are all considerably reduced by using ionic liquid as entrainer. However by rigorously calculating the overall total annual cost including the entrainer recovery section, the economics become only comparable to the systems by using conventional entrainer. For acetone/methanol system, process using [MMIM][DMP] as entrainer and two-flash-drum entrainer recovery configuration only saves 10.07% TAC because of needing an expensive screw compressor in the vacuum system. For IPA dehydration system, two sets of product specification are discussed. For the 99.9999 mol% IPA case using [EMIM][OAC] as entrainer, operating pressure at the second flash drum needs to be too low to even reject the use screw compressor in this vacuum system.


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Therefore, only one-flash-drum-one-stripping-column entrainer recovery configuration could be applied. This case leads to 20.07% more TAC than the conventional process of using DMSO as entrainer. For the 99.99 mol% IPA case, both entrainer recovery configurations

could

be

one-flash-drum-one-stripping-column

applied, configuration

but is

only more

IL economical

with than

conventional case. However, the saving of 8.36% TAC is also not very significant. Ionic liquids have received much attention for their versatility; however, low thermal stability and high normal boiling point of ILs make its application in azeotropic separation not as appealing as expected. Ionic liquid with higher thermal degradation temperature is favorable to be applied in the extractive distillation system. Choosing one-flash-drum-one-stripping-column configuration can alleviate the highly vacuum requirement in the second-stage of entrainer recovery section. However, it is only economically favorable if the first flash drum can separate out most of the product component from ionic liquid. Otherwise, too much stripping agent would be needed and circulated in the system causing the total annual cost to go up considerably.

Acknowledgements The research funding from the Ministry of Science and Technology of the R. O. C. under grant no. MOST 105-2221-E-002-210 is greatly appreciated.



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Nomenclature Capital Letters [MMIM][DMP]

1,3-dimethylimidazolium dimethylphosphate

DMSO

Dimethyl sulfoxide

[EMIM][OAc]

1-ethyl-3-methylimidazolium acetate

E

Entrainer feed flowrate (kmol/hr)

F

Fresh feed flowrate (kmol/hr)

HX

Heat exchanger

IL

Ionic liquid

IPA

Isopropanol

NE

Entrainer feed stage

NF

Fresh mixture feed stage

NT

Total number of stages of extractive distillation column

P1

Operating pressure of vacuum system 1 (Pa)

P2

Operating pressure of vacuum system 2 (Pa)

P’

Suction pressure of vacuum system (torr)

STR

Stripping column

TAC

Total annual cost (k$/year)

TAC1

Total annual cost of extractive distillation section (k$/year)

TAC2

Total annual cost od entrainer recovery section (k$/year)

V

vessel volume (ft3)

W

air leakage rate (lb/hr)

Xmmax

Maximum methanol impurity mole fraction in IL recycle stream

Xwmax

Maximum water impurity mole fraction in IL recycle stream


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Lower case Letters xi

Liquid mole fraction of component i

yi

Vapor mole fraction of component i

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water + 2-propanol system at 30, 60, and 100 kPa, J. Chem. Eng. Data. 1996, 41, 608-611. (35) Li, Q.; Zhang, J.; Lei, Z.; Zhu, J.; Wang, B.; Huang, X. Isobaric Vapor-Liquid

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Table 1 Summarized thermodynamic properties and model parameters. Ionic liquid

[MMIM][DMP]

[EMIM][OAC]

Mw

222.20

170.20

Pc (bar)

27.18

29.19

Vc (cc/mol)

626.80

544.00

Zc

0.25

0.24

Tb (K)

1074.79

800

900

1000

Tc (K)

1089.95

1150.97

1291.97

1432.97

*1

CPIG C1

99.69

C2

0.25

93.33 0.28

C3

1.04 × 10

-3

5.27 × 10-4

C4

-6.69 × 10-7

-4.23 × 10-7 PLXANT*2

C1

23.98

74.11

76.00

77.63

C2

-13387.06

-14382.34

-16353.03

-18325.73

C5

0

-6.68

-6.81

-6.92

C6

0

C7

0

1.49 × 10

-19

6.00 DHVLWT

7.56 × 10

-20

4.10 × 10-20

6.00

6.00

*3

C1

109.106

135.749

119.103

103.146

C2

1074.79

273.15

273.15

273.15

C3

3.13 × 10

0.40

0.42

0.42

C4

0.033

0.13

-0.09

-0.13

-3

*1 C op = C1 + C 2 T + C 3T 2 + C 4 T 3 + C 5 T 4 + C 6 T 5 , 0 ≤ T ≤ 1000 (K), where C op is ideal gas heat capacity (J/mol-K), the defaults of C4 and C5 are 0. *2 ln P vap = C1 +

C2 + C 4 T + C5 ln T + C 6 T C 7 , 0 ≤ T ≤ 1000 (K), T + C3

where Pvap is liquid vapor pressure (Pa), the defaults of C3 and C4 are 0. Parameters of [MMIM][DMP] are retrieved from Heym et al.’s work7. The range for temperature in the above two equations is to make sure that all calculations at operating temperatures in the design flowsheets are based on the above two equations without using any extrapolation.



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

 1 − T / Tc   *3 ∆ vap H = C1   1 − C 2 / Tc 

C3+C4(1−T / Tc )

,

where ∆ vapH is heat of vaporization (kJ/mol).


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Table 2a NRTL model parameters of acetone/methanol/[MMIM][DMP] system Component i

Acetone

Component j

Methanol

Source

Chen et al.

25

Acetone

Methanol

[MMIM][DMP]

[MMIM][DMP]

Regression

Regression

aij

0

0

0

aji

0

0

0

bij (K)

111.16

1126.46

4477.79

bji (K)

103.81

-779.08

-3183.50

cij

0.300

0.300

0.103

Table 2b NRTL model parameters of IPA/water/[EMIM][OAC] system Component i

Isopropanol

Isopropanol

Water

Component j

Water

[EMIM][OAC]

[EMIM][OAC]

Regression

Regression

Source

Marzal et al.

34

aij

0

0

0

aji

0

0

0

bij (K)

9.32

-371.15

1763.71

bji (K)

830.02

1066.95

-1201.51

cij

0.30

0.30

0.30



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Page 34 of 53

Table 3 TAC comparison for cooling water (Base case 1) and chilled water (Base case 2) as cooling agent at the first pre-condenser for the acetone/methanol system. Base case 1

Base case 2

C1

Flash1

Flash2

C1

Flash1

Flash2

shell capital cost (k$)

400.44

55.18

44.51

400.44

50.69

44.51

condenser/pre-condenser capital cost (k$)

302.53

284.01

259.05

302.53

203.76

176.75

reboiler/heater capital cost (k$)

131.87

81.15

112.87

131.87

113.22

80.73

vacuum system capital cost (k$)

-

1.49

869.24

-

2.26

868.95

vacuum system operating cost (k$/yr)

-

16.91

1.58

-

16.01

1.58

38.21

12.14

-

38.21

263.08

-

-

-

1070.26

-

-

594.39

699.09

306.35

508.93

699.09

511.37

303.91

cooling/chilled water cost (k$/yr) refrigerant cost (k$/yr) steam cost (k$/yr) △Capital cost (%)

-

-

0

-9.75

△Operating cost (%)

-

-

0

-11.79

cooler capital cost (k$)

29.02

29.02

HX capital cost (k$)

45.21

45.21

cooler cooling water cost (k$/yr)

0.66

0.66

methanol lost cost (k$/yr)

38.96

38.14

2616.57

2449.96

-

-6.37

total operating cost (k$/yr)

2693.10

2466.44

△Total operating cost (%)

-

-8.42

TAC (k$/yr)

3565.29

3283.09

△TAC (%)

-

-7.92

total capital cost (k$) △Total capital cost (%)


Page 35 of 53

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


Table 4 TAC comparison for acetone/methanol system. DMSO

[MMIM][DMP]

C1

C2

C1

Flash1

Flash2


620.45

247.30

401.06

50.23

44.52


433.75

204.06

302.97

208.90

167.30


280.90

192.56

132.16

116.40

76.78


-

-

-

2.80

769.89


-

-

-

22.79

1.27

66.54

33.11

38.29

273.36

--

-

-

-

-

546.17

1139.69

1464.34

701.21

533.65

281.36

△Capital cost (%)

-

-

-37.37

123.14


-

-

-38.69

10.76

cooling/chilled water cost (k$/yr) refrigerant cost (k$/yr) steam cost (k$/yr)


25.25

29.03


69.27

45.24


2.89

0.66

128.65

-

-

46.89

2262.58

2445.65

-

8.09


2835.23

2347.27


-

-17.21

TAC (k$/yr)

3589.42

3228.07

△TAC (%)

-

-10.07

entrainer makeup cost (k$/yr) methanol lost cost (k$/yr) total capital cost (k$) △Total capital cost (%)



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

Page 36 of 53

Table 5a TAC comparison for 99.9999 mol% IPA dehydration case. DMSO

[EMIM][OAC] one flash drum + stripping column

C1

C2

C1

Flash1

STR


206.17

130.00

176.32

26.62

30.17


66.01

22.83

50.81

97.02

29.17


105.57

68.95

104.01

102.23

4.20


-

-

-

8.85

640.89


-

-

-

15.73

0.11

10.34

8.39

6.91

84.01

-

-

-

-

-

27.13

250.29

292.81

197.84

147.76

2.75

△Capital cost (%)

-

-

-12.34

323.46


-

-

-21.44

-7.87



36.62

13.22

entrainer makeup cost (k$/yr)

0.03

-

-

19.14

636.15

1283.51

-

101.76


561.86

501.38


-

-10.76

TAC (k$/yr)

773.91

929.22

△TAC (%)

-

20.07

IPA lost cost (k$/yr) total capital cost (k$) △Total capital cost (%)


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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


Table 5b TAC comparison for 99.99 mol% IPA dehydration case. [EMIM][OAC]

[EMIM][OAC]

two flash drums

flash drum + stripping column

DMSO C1

C2

C1

Flash1

Flash2

C1

Flash1

STR


217.63

91.62

203.91

22.70

22.68

209.83

21.60

0.05


61.81

31.42

54.09

62.87

11.49

53.88

64.89

16.38


108.14

47.29

119.26

64.70

9.69

111.04

67.43

1.75


-

-

-

69.51

838.60

-

7.00

16.88


-

-

-

5.80

0.13

-

10.03

3.80

9.35

4.77

7.61

43.09

-

7.56

45.25

-

-

-

-

-

13.19

-

-

11.17

204.91

164.94

162.31

73.10

3.94

171.36

77.90

0.74

△Capital cost (%)

-

-

-2.66

547.13

-3.31

15.06


-

-

-20.69

-17.94

-16.49

-12.27



30.87

17.08

15.07


9.44

7.10

6.74


0.50

0.25

0.23

entrainer makeup cost (k$/yr)

0.42

-

-

-

-

9.04

588.79

1441.18

585.81

-

144.77

-0.51


384.38

313.70

336.86


-

-18.39

-12.36

TAC (k$/yr)

580.64

794.10

532.12

△TAC (%)

-

36.76

-8.36

IPA lost cost (k$/yr) total capital cost (k$) △Total capital cost (%)



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

Figure Captions Figure 1

NRTL model prediction and experimental data of acetone (1), methanol (2) and [MMIM][DMP] (3) mixture, parameters from Table 2a.

Figure 2

NRTL model prediction and experimental data of IPA (1), water (2) and [EMIM][OAC] (3) mixture, parameters from Table 2b.

Figure 3

Illustration of how to simulate volatile component recovery system in pre-condenser. (a) steam-jet ejector (b) liquid ring pump/screw compressor

Figure 4

Process flowsheet with (4a) cooling water and (4b) chilled water as the cooing agent at the first pre-condenser for acetone/methanol separation.

Figure 5

Sequential iterative optimization procedures for acetone/methanol separation.

Figure 6

Optimization result (a) summarized plot of total TAC (b) TAC1 and TAC2 when E/F = 0.15.

Figure 7

Optimal process flowsheet of acetone/methanol separation with [MMIM][DMP].

Figure 8

Process flowsheet for 99.9999 mol% IPA case using (8a) two flash drums and (8b) one flash drum and one stripping column to recover IL, assuming the normal boiling point of [EMIM][OAC] is 1000K.

Figure 9

Process flowsheet for 99.99 mol% IPA case using (9a) two flash drums and (9b) one flash drums and one stripping column to recover IL, assuming the normal boiling point of [EMIM][OAC] is 1000K.

Figure 10 Relationship between TAC and IPA loss fraction for 99.99 mol% IPA case. Figure 11 Sensitivity analysis with different assumption of normal boiling point of [EMIM][OAC] (a) TAC (b) TAC1 (c) TAC2.


Page 38 of 53

Page 39 of 53

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


Figure 1

Where x'1 =

x3 x1 x2 , x '2 = , E/F= x1 + x 2 x1 + x 2 x1 + x 2



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

Page 40 of 53

Figure 2

Where x'1 =

x3 x1 x2 , x '2 = , E/F= x1 + x 2 x1 + x 2 x1 + x 2


Page 41 of 53


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Figure 3



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

Page 42 of 53

Figure 4a

Figure 4b


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Figure 5



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

Page 44 of 53

Figure 6a


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Figure 6b



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

Page 46 of 53

Figure 7


Page 47 of 53


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Figure 8a



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

Page 48 of 53

Figure 8b


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Figure 9a



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

Page 50 of 53

Figure 9b


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Figure 10



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

Page 52 of 53

Figure 11a

Figure 11b

Figure 11c


Page 53 of 53

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Critical Assessment of Using an Ionic Liquid as Entrainer via Extractive

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