Vapor-liquid equilibrium experiment and process simulation of extractive distillation for separating diisopropyl ether-isopropyl alcohol using ionic liquid

Vapor-liquid equilibrium experiment and process simulation of extractive distillation for separating diisopropyl ether-isopropyl alcohol using ionic liquid

Journal of Molecular Liquids 293 (2019) 111406 Contents lists available at ScienceDirect Journal of Molecular Liquids journal homepage: www.elsevier...

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Journal of Molecular Liquids 293 (2019) 111406

Contents lists available at ScienceDirect

Journal of Molecular Liquids journal homepage: www.elsevier.com/locate/molliq

Vapor-liquid equilibrium experiment and process simulation of extractive distillation for separating diisopropyl ether-isopropyl alcohol using ionic liquid Jun Qi a, Qingjun Zhang b, Xiaoyan Han a, Qingpeng Wu a, Yafang Li a, Qunsheng Li a,⁎ a b

State Key Laboratory of Chemical Resource Engineering, Beijing University of Chemical Technology, Box 266, Beijing 100029, China State Key Laboratory of Chemical Engineering, Tianjin University, School of Chemical Engineering and Technology, Tianjin 30072, China

a r t i c l e

i n f o

Article history: Received 7 June 2019 Received in revised form 17 July 2019 Accepted 19 July 2019 Available online 07 August 2019 Keywords: Vapor-liquid equilibrium Diisopropyl ether-isopropyl alcohol Ionic liquids Extractive distillation Process simulation

a b s t r a c t The formation of azeotrope makes the separation of diisopropyl ether (DIPE) and isopropanol (IPA) a challenging task. Ethylene glycol (EG) and ionic liquid 1,3-dimethylimidazolium dimethyl phosphate ([Mmim] [DMP]) were screened as entrainers for the extractive distillation process. Isobaric vapor-liquid equilibrium data of DIPE + IPA + EG and DIPE + IPA + [Mmim] [DMP] was measured through experiment. The NRTL model was applied to correlate the experiment data and predict the effect of entrainers on the phase behavior of azeotrope in the full concentration range. The results indicate that [Mmim] [DMP] performs better ability than EG in enhancing the relative volatility of DIPE/IPA azeotrope and eliminating azeotropic point. The design and simulation of the two extractive distillation schemes were carried out after embedding the model parameters and the physical properties of IL into Aspen Plus. The optimal operating parameters were obtained though the sequential iterative procedure based on the total annual cost (TAC). The results of the economic analysis showed that the scheme 2 using the flash tank instead of the solvent recovery column has 24.582% reduction of TAC and 31.241% energy savings compared to the scheme 1. © 2019 Elsevier B.V. All rights reserved.

1. Introduction Diisopropyl ether (DIPE), an excellent solvent for oils and fats, can be used as gasoline additive due to its high octane number and antifreeze performance. In addition, it can also be used as a standard solvent for chromatographic analysis and anesthetic [1–3]. Isopropyl alcohol (IPA), which is a widely used organic raw material and solvent, has demonstrated the potential as a green bioenergy with the gradual depleting of fossil energy [4–5]. In the process of producing IPA by propylene hydration method, DIPE is produced as a by-product. Therefore, realizing the separation of DIPE/IPA has important economic and environmental values. It should be noted that the formation of homogeneous azeotrope makes this energy-intensive separation process extremely challenging, which makes conventional distillation difficult to meet the separation requirement. Several special alternatives (such as extractive distillation [6–8], azeotropic distillation [9–10], pressure swing distillation [11–14], membrane separation) have been proven as effective methods to separate azeotropes. Extractive distillation is the most common way to separate DIPE and IPA. The addition of a heavy entrainer enhances the relative volatility of DIPE and IPA and eliminates the azeotropic point. Luo ⁎ Corresponding author. E-mail address: [email protected] (Q. Li).

https://doi.org/10.1016/j.molliq.2019.111406 0167-7322/© 2019 Elsevier B.V. All rights reserved.

et al. explored the possibility pressure-swing distillation and extractive distillation using heavy entrainer 2-methoxyethanol [15]. They further compared the cost and dynamic control of the two schemes and found that the heat-integrated pressure-swing distillation system has 5.75% reduction in the TAC and 7.97% in energy cost. You et al. further explored the effects of pressure on the extractive distillation scheme and analyzed it from a thermodynamic perspective. They found that low pressure can reduce the amount of entrainers and significantly reduce the energy consumption. A competitive comparison was found: energy cost and TAC are reduced by 13.4% and 6.3%, respectively compared to Luo's design [16]. Recently, Zhang et al. proposed an extractive distillation scheme using butyl acetate as an entrainer and verified its feasibility by phase equilibrium experiments and process simulation [17]. As environmental pollution problems worsen, the emission of volatile organic compounds produced by using conventional entrainers has attracted increasing attention, while ionic liquids (ILs) are promising as substitute for traditional entrainers due to their low vapor pressure, high boiling point, environmental friendliness and high selectivity [18–19]. Zhang et al. researched the effect of [Bmim] [BF4], [Emim] [BF4], [Bmim] [N(CN)2], [Emim] [N(CN)2], [Bmim] [Cl], [Emim] [Cl], [Bmim] [OAC], and [Emim] [OAC] on the phase equilibrium of ethanolwater system [20–21]. In addition, many research works have reported the use of ionic liquids for the separation of various azeotropes composed of alcohols, ethers and esters, such as methanol-methyl ethyl

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ketone, ethanol-ethyl acetate, and isopropanol-acetone [22–26]. However, there is no literature on the separation of DIPE/IPA azeotrope by extractive distillation using ionic liquids. [Mmim] [DMP] was screened from ILs with different anions (such as − − − − BF− 4 , PF6 , N(CN)2 , DMP , NTf2 ) by preliminary experiments because of its ability to enhance the relative volatility of DIPE and IPA. The main contributions of this paper include the following four sections: (1) The isobaric VLE experiment of DIPE + IPA + [Mmim] [DMP] and DIPE + IPA + EG was completed through an improved Othmer. (2) The correlation of experimental data was completed based on NRTL model and the corresponding interaction parameters containing IL were obtained. (3) After embedding the model parameters and IL physical properties into Aspen Plus, the process simulation of two extractive distillation schemes was completed for separating DIPE/IPA azeotrope. (4) Based on the total annual cost (TAC), the economic analysis was completed for the two schemes, and the optimal process design was obtained. 2. Experimental 2.1. Chemicals The conventional chemical materials such as IPA and DIPE used in this paper were provided by Beijing Chemical Reagents Company. The selected ionic liquid [Mmim] [DMP], which was supplied by Shanghai Chengjie Chemical Co. LTD with a purity of 99 mass%, was added to the rotary evaporator for 48 h under negative pressure and 373.15 K to remove as much water as possible before the experiment. After that, the ionic liquid was placed in a Karl Fischer instrument to measure moisture content. Detailed information on the purity and source of all chemicals is shown in Table 1. 2.2. Apparatus and procedure An analytical balance with standard uncertainty of 0.1 mg was selected as a weighing instrument for weighing the required raw materials, which was to be formulated into 60 ml experimental samples by predetermined ratios in 100 ml Erlenmeyer flasks. The main instrument of this experiment was an improved Othmer used to measure isobaric VLE data for binary and ternary systems. The standard uncertainty of the thermometer used to measure the equilibrium temperature was 0.1 K. The composition of the vapor phase and the liquid phase at equilibrium was measured by gas chromatography (GC 9790II) of Zhejiang Fuli Company, in which the detector was FID and the column was KB5 column (30 m × 0.53 mm). The operating conditions were as follows: oven temperature at 353.15 K, injector temperature at 413.15 K and detector temperature at 423.15 K. N2 was used as a carrier gas whose flow rate was 35 cm3·min−1. First, the vapor-liquid equilibrium data of the DIPE/IPA azeotrope was measured to verify the accuracy of the experimental apparatus. Then, a ternary sample containing entrainer in a certain molar ratio was added to Othmer for heating. When the condensed droplets of the vapor phase port were maintained at a rate of 2 drops/s and the temperature was constant, heating was continued for 30 min to ensure

that the vapor-liquid phase reached equilibrium. When the vapor-liquid phase composition was measured using gas chromatography, the operation was repeated three times to avoid random errors. 3. Results and discussion 3.1. Experimental data The VLE data for DIPE (1) + IPA (2) are shown in Table 2, where T represents equilibrium temperature, x1 and y1 represent the liquid and vapor phase mole fractions of DIPE. By comparing the literature data with the correlation results based on the NRTL model (as displayed in Fig. 1), it was found that the binary VLE experimental data is credible, and the phase equilibrium behavior of the DIPE/IPA binary system could be predicted accurately by the NRTL model. The effect of entrainers on the phase equilibrium of the DIPE/IPA mixture was determined experimentally and the corresponding experimental data are listed in Tables 3 and 4, where x3 represents the mole fraction of the entrainer, x′1 represents the liquid mole fraction of DIPE (without entrainer), γ1 and γ2 refer to the activity coefficients of this two components, and the expression is shown in Eq. (1).

γi ¼

yi P xi Psi

ð1Þ

where P is the system operating pressure, xi and yi represent the mole fraction of the liquid and vapor phase containing the entrainer, and Psi refers to the saturated vapor pressure of pure component i. Relative volatility (α12) is an important indicator to reveal the effect of entrainer on the phase equilibrium behavior of azeotropes, calculated by the following Eq. (2). It is well known that the relative volatility of the azeotropes is 1, which is the reason why conventional distillation cannot achieve effective separation. And the greater the relative volatility of the system, the lower the energy consumption of the separation process.

α 12 ¼

y1 =x1 y2 =x2

ð2Þ

3.2. Correlation of model Suitable thermodynamic models are important for predicting the phase equilibrium behavior, PVT properties, etc. of fluid mixtures. The NRTL model chosen in this paper is suitable for strong non-ideal systems, especially for systems with limited miscibility. It introduces a non-random parameter α12 that reflects the characteristics of the system in the Boltzmann equation, which makes it more accurately predict the phase equilibrium behavior of the azeotropes containing ionic liquid

Table 1 Detailed information of experimental materials. Name

CAS

Source

DIPE IPA EG [Mmim] [DMP]a

108-20-3 67-63-0 107-21-1 654058-04-5

Beijing Chemical Reagents Company. Beijing Chemical Reagents Company. Beijing Chemical Reagents Company. Shanghai Chengjie Chemical Co. Ltd.

a b c d

[Mmim] [DMP] = 1,3-dimethylimidazolium dimethylphosphate. GC = Gas chromatography. KF = Karl Fischer instrument. The values of the boiling point are derived from the Aspen plus database.

Mass fraction

Boiling point/Kd

Analysis method

≥99.9% ≥99.9% ≥99.9% ≥99%

341.77 355.27 470.45 –

GCb GCb GCb KFc

J. Qi et al. / Journal of Molecular Liquids 293 (2019) 111406 Table 2 The VLE data for DIPE (1) + IPA (2) based on experiment (at 101.3 kPa).

Table 3 Ternary vapor-liquid phase equilibrium data of DIPE (1) + IPA (2) + [Mmim] [DMP] (3), x ′1 represents the mole fraction of DIPE in the liquid phase without IL.

T/K

x1

y1

351.3 348.0 345.4 343.3 341.7 340.5 339.7 339.2 339.2 339.8

0.061 0.128 0.201 0.282 0.370 0.469 0.578 0.702 0.841 0.918

0.197 0.342 0.450 0.533 0.597 0.649 0.695 0.742 0.812 0.876

Standard uncertainties u of T, P, x and y are u(T) = 0.3 K, u(P) = 0.3 kPa, u(x) = 0.005, u (y) = 0.008.

[28–29]. The activity coefficient equation is as follows: X



X 2 3 τ ij Gij xi X 7 Gij x j 6 j i 6τ − X ln γi ¼ X þ  7 Gki xk 4 ij ðGki xk Þ Gkj xk 5 j τ ji Gji x j

k

τji ¼

ð3Þ

k

g ji −g jj RT

ð4Þ

  Gji ¼ exp −α ji τji

ð5Þ

where gji is the interaction parameters between molecules of i and j. The parameter αji is related to the non-randomness in the mixture, usually from 0.2 to 0.47. In the previous part, the binary VLE data of DIPE/IPA azeotrope has been correlated and verified by NRTL model. After that, the experimentally obtained VLE data containing IL was embedded into the Matlab program for correlation, and the interaction parameter g of the corresponding NRTL model was obtained. The objective function F (as shown in Eq. (6)) is introduced to obtain the deviation between the experimental data and the model correlation results [30].



X n

2 4 1− γ 1cal γ1 exp

!2 þ

γ 1− 2cal γ 2 exp

!2 32 5

ð6Þ

where γiexp and γical are the activity coefficients obtained by experiment and correlation. It was found that the NRTL model could accurately 360

355

T/K

350

345

340

335 0.0

0.2

0.4

0.6

0.8

3

1.0

x1, y1 Fig. 1. T-x-y diagram of DIPE (1) + IPA (2) mixture at 101.3 kPa; ○, literature value [27]; ●, experimental value; solid line, NRTL model value.

T/K

100 x3

x′1

y1

γ1

γ2

α12

351.0 346.1 342.7 341.0 340.1 340.0 340.4 340.8 341.3 351.6 346.4 343.1 341.1 340.2 340.3 340.6 341.1 341.5 352.6 346.3 343.3 340.7 340.0 340.0 340.1 340.6 341.0

3.011 3.034 3.051 3.002 3.011 3.100 3.003 3.021 3.014 6.002 6.043 6.006 6.021 6.005 6.072 6.054 6.124 6.002 9.995 9.992 10.001 9.999 9.998 9.989 9.991 10.006 10.003

0.085 0.189 0.310 0.421 0.590 0.678 0.796 0.870 0.923 0.088 0.189 0.292 0.418 0.586 0.641 0.735 0.838 0.892 0.083 0.197 0.311 0.482 0.542 0.758 0.816 0.853 0.914

0.278 0.484 0.611 0.690 0.769 0.800 0.836 0.877 0.919 0.318 0.536 0.657 0.740 0.799 0.815 0.841 0.883 0.908 0.339 0.584 0.706 0.801 0.823 0.889 0.900 0.916 0.939

2.523 2.302 1.976 1.734 1.420 1.291 1.133 1.073 1.044 2.808 2.598 2.288 1.919 1.523 1.418 1.264 1.145 1.091 3.226 2.820 2.383 1.897 1.771 1.368 1.283 1.230 1.160

0.985 0.975 0.999 1.023 1.123 1.242 1.582 1.842 1.987 0.940 0.893 0.872 0.881 0.997 1.053 1.211 1.447 1.688 0.909 0.850 0.798 0.811 0.841 1.017 1.216 1.273 1.649

4.143 4.026 3.504 3.063 2.309 1.902 1.306 1.060 0.955 4.817 4.960 4.644 3.950 2.803 2.474 1.919 1.459 1.198 5.702 5.707 5.328 4.318 3.926 2.560 2.033 1.879 1.437

Standard uncertainties u of T, P, x and y are u(T) = 0.3 K, u(P) = 0.3 kPa, u(x) = 0.006, u (y) = 0.006.

correlate and predict the vapor-liquid equilibrium data of ternary systems containing IL or EG, with a relative error of only 2.37% and 3.65%. The relevant parameters of the NRTL model are listed in Table 5. The y-x diagrams and T-x-y diagrams of DIPE (1) + IPA (2) + IL (3) and DIPE (1) + IPA (2) + EG (3) are shown in Figs. 2 and 3, respectively.

Table 4 Ternary VLE data of DIPE (1) + IPA (2) + EG (3), mole fraction of DIPE in the liquid phase without EG x′1, activity coefficient γ, relative volatility of DIPE and IPA α12. T/K

100 x3

x′1

y1

γ1

γ2

α12

350.7 346.8 344.3 342.2 341.4 340.9 340.4 340.3 340.0 351.2 347.6 344.7 342.6 341.6 341.4 341.3 341.3 341.0 353.1 348.3 345.0 343.5 342.4 341.9 341.7 341.6 341.8

4.998 5.020 5.015 5.023 4.997 4.981 5.001 5.026 5.015 9.996 10.034 10.006 10.016 9.994 10.003 10.014 10.021 10.021 15.016 14.997 15.029 15.017 15.007 14.997 14.999 14.998 15.018

0.090 0.187 0.269 0.375 0.469 0.589 0.695 0.795 0.903 0.095 0.181 0.272 0.388 0.497 0.594 0.717 0.793 0.903 0.091 0.193 0.316 0.369 0.448 0.585 0.723 0.809 0.900

0.221 0.380 0.473 0.567 0.625 0.692 0.748 0.811 0.877 0.258 0.393 0.505 0.610 0.679 0.731 0.786 0.830 0.893 0.261 0.437 0.582 0.621 0.680 0.753 0.812 0.857 0.909

1.952 1.817 1.703 1.571 1.419 1.270 1.183 1.126 1.083 2.228 1.991 1.863 1.688 1.513 1.372 1.227 1.171 1.117 2.325 2.119 1.909 1.832 1.714 1.476 1.296 1.227 1.163

1.104 1.159 1.220 1.282 1.356 1.475 1.667 1.872 2.636 1.095 1.150 1.195 1.230 1.292 1.361 1.574 1.735 2.505 1.065 1.117 1.134 1.191 1.209 1.289 1.528 1.760 2.408

2.874 2.658 2.434 2.190 1.888 1.564 1.299 1.107 0.769 3.318 2.937 2.735 2.472 2.141 1.856 1.453 1.276 0.898 3.539 3.246 3.003 2.797 2.623 2.161 1.652 1.416 1.115

Standard uncertainties u of T, P, x and y are u(T) = 0.3 K, u(P) = 0.3 kPa, u(x) = 0.007, u (y) = 0.008.

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J. Qi et al. / Journal of Molecular Liquids 293 (2019) 111406

produces different molecular forces for DIPE and IPA [31,32]. Therefore, IPA is more retained in the liquid phase during vapor-liquid contact.

Table 5 Interaction parameters of DIPE, IPA and entrainers of NRTL model. Component i

Component j

αij

DIPE DIPE IPA DIPE IPA

IPA EG EG [Mmim] [DMP] [Mmim] [DMP]

0.30 0.20 0.47 0.47 0.47

Δgij/J·mol−1

Δgji/J·mol−1

3492.96 4721.51 14,440.01 34,404.68 7435.40

389.36 5846.62 −529.12 6669.70 −7592.84

4. Process design and simulation 4.1. Design basis The simulation basis for extractive distillation with ionic liquid was to determine the binary interaction parameters between IL and azeotrope, which can be obtained by correlating VLE data based on the NRTL model, as shown in Table 5. It should be noted that DIPE and IPA are common organic components and have detailed physical data in

α12 can directly reflect the effect of entrainer on the phase equilibrium behavior of azeotrope. Fig. 4 provides the relative volatility of

1.0

1.0

(b)

(a)

0.8

0.6

0.6

y1

y1

0.8

0.4

0.4

0.2

0.2

0.0 0.0

0.2

0.4

0.6

0.8

0.0 0.0

1.0

0.2

0.4

0.6

0.8

1.0

x’1

x’1

Fig. 2. y-x diagram of DIPE (1) + IPA (2) + entrainer (3) at 101.3 kPa; (a) [Mmim] [DMP]: □, x3 ≈ 0.03, ○, x3 ≈ 0.06, △, x3 ≈ 0.10; (b) EG: □, x3 ≈ 0.05, ○, x3 ≈ 0.10, △, x3 ≈ 0.15; solid line, correlation results of NRTL model.

DIPE and IPA with addition of [Mmim] [DMP] or EG, indicating that [Mmim] [DMP] has prominent effect in increasing α12 of DIPE to IPA compared to EG. When [Mmim] [DMP] was added in an amount of only 0.05 (mole fraction), the azeotropic point disappeared. An acceptable explanation from many literatures is that [Mmim] [DMP] forms a hydrogen bond with the hydroxyl group in IPA so that [Mmim] [DMP]

(a)

the Aspen Plus database. However, for ionic liquids, due to their unusual structural and physical properties, the relevant physical data has not been provided yet in the database. Therefore, the group contribution method proposed by Jose'O et al. was chosen to estimate the critical properties of ionic liquids [33]. The extended Antoine saturated vapor pressure equation parameters were calculated using the method

(b)

360

360

355

T/K

T/K

355

350

345

345

340 0.0

350

340 0.2

0.4

0.6

x1', y1

0.8

1.0

0.0

0.2

0.4

0.6

0.8

1.0

x’1, y1

Fig. 3. T-y-x diagram for DIPE (1) + IPA (2) + entrainer (3) at 101.3 kPa; (a) [Mmim] [DMP]: ○, x3 ≈ 0.03, △, x3 ≈ 0.06, □, x3 ≈ 0.10; (b) EG: ○, x3 ≈ 0.05, △, x3 ≈ 0.10, □, x3 ≈ 0.15; solid line, correlation result of NRTL model.

J. Qi et al. / Journal of Molecular Liquids 293 (2019) 111406

(a)

8

(b)

4

3

α12

α12

6

4

2

0 0.0

5

2

1

0.2

0.4

0.6

0.8

1.0

0 0.0

x1'

0.2

0.4

0.6

0.8

1.0

x1'

Fig. 4. α12 ~ x′1 diagram for DIPE (1) + IPA (2) + entrainer (3); (a) [Mmim] [DMP] (□, x3 ≈ 0.03, ○, x3 ≈ 0.06, △, x3 ≈ 0.10); (b) EG (□, x3 ≈ 0.05, ○, x3 ≈ 0.10, △, x3 ≈ 0.15); solid line, correlation result of NRTL model.

proposed by Rudkin. The parameters of the isobaric molar heat capacity equation were estimated by the method proposed by Gardas et al. [34] The corresponding expressions are displayed in Eq. (7) and Eq. (8), and detailed information on the physical properties of [Mmim] [DMP] is listed in the Tables S1–S3. ln P si ¼ D1 þ

D2 þ D4 T þ D5 ln T þ D6 T D7 T þ D3

ð7Þ

CP ¼ C1 þ C2 T þ C3 T 2 þ C4 T 3 þ C5 T 4 þ C6 T 5

ð8Þ

where psi is the saturated vapor pressure (kPa), the unit of T is K, and Cp is the isobaric molar heat capacity (J·mol−1·K−1).

4.2. Process simulation with [Mmim] [DMP] as entrainer The physical properties of [Mmim] [DMP] and the binary interaction parameters of the mixture were embedded in the Aspen Plus software, then the simulation of the separation process was implemented. The main equipment consists of an extractive distillation column (EDC), a entrainer recovery column (ERC) and the matching heat exchangers. The mixture with a flow rate of 100 kmol/h (50% mol% DIPE + 50% mol% IPA, random setting) was fed to the middle stage of the EDC. [Mmim] [DMP] entered from the top of the EDC. The EDC was divided into three parts: the rectifying section, the extraction section, and the stripping section [35]. When [Mmim] [DMP] and the azeotrope were countercurrently contacted in the extraction section, the strong interaction between [Mmim] [DMP] and IPA increased the relative

Fix P1=1atm P2=1atm Give NT2 Give S1 Give NF2 Give NT1 Vary RR2,D2 to meet two design specifications of ERC Give NF1, NFE1 No Is QR2 minimal with NT2 fixed? Vary D1, RR1 to meet one design specifications of EDC Yes No

Is QR1 minimal with NT1 fixed?

No Is TAC minimal ? Yes

Yes No Is TAC minimal ?

Get the optimal NT2 ,NF2

Yes Get optimal NF1 NFE1 NT1 S1

Over

Fig. 5. Sequential iterative optimization procedure for process simulation with IL as an entrainer.

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J. Qi et al. / Journal of Molecular Liquids 293 (2019) 111406

1150

290000 289000

QR1

TAC($/y)

1100

1050

288000 287000 286000

1000

285000 950

8

10

12

14

284000

16

NF1 300000

20

NT1

22

24

1.0

DIPE IPA [Mmim][DMP]

0.8

Mole fraction of liquid

296000

TAC($/y)

18

292000

288000

0.6

0.4

0.2

284000

0.0 14

16

18

20

S1

22

0

5

10

15

20

NT1 369000

1430 368000

QR2

TAC($/y)

1420

367000

1410 366000

1400

2

4

6

8

365000

6

NF2

8

10

12

NT2 Fig. 6. Optimization results for operating variables.

volatility of DIPE and IPA and eliminated azeotropic point, so the light component DIPE (99.9 mol%) was obtained from the top of the EDC. For further separation, [Mmim] [DMP] and IPA mixture was sent to the ERC in which the distillate is 99.9 mol% of IPA and the bottom IL was fed back to the EDC. Product purity settings are derived from industry requirements. For the optimization of operating variables, the minimum TAC was chosen as the objective function and the sequential iterative

procedure was selected as optimization method as shown in Fig. 5. Variables that need to be determined contain total stages (NT1 , NT2), feed stage (NF1 , NFE1 , NF2 ), entrainer flow rate (S 1 ). As the most critical variable, S1 not only affects the separation of azeotrope in the EDC, but also affects the energy consumption of the ERC. The optimization results of the operating variables are presented in Fig. 6, which are: NT1 = 15, NT2 = 16, NF1 = 20, NFE1 = 2, NF2 = 5, S1 = 19 kmol/h. Fig. 7 provides the optimal flowsheet of extractive

J. Qi et al. / Journal of Molecular Liquids 293 (2019) 111406

7

Fig. 7. Optimal flowsheet of DIPE/IPA separation with IL as entrainer.

distillation with [Mmim] [DMP] as entrainer. The total energy consumption of this separation process is 2650 kW. 4.3. Improved scheme with [Mmim] [DMP] as entrainer The problem to be aware of is that IL has physical properties of low vapor pressure and high boiling point, which inevitably causes the temperature of the ERC to be much higher than the temperature of the high pressure steam. Therefore, it is necessary to introduce an alkyl biphenyl conduction oil to heat the reboiler. In addition, the use of the ERC also causes an increase in energy consumption. A feasible improvement is to replace the ERC with a flash tank, which has also been reported in many literatures. However, preliminary simulation indicated that it is difficult to separate [Mmim] [DMP] from IPA by single-stage flash. The possible reasons for the analysis are:

(1) there is a strong interaction between the two components; (2) [Mmim] [DMP] is more volatile than other ILs. Therefore, the separation scheme needs to be further improved as follows. First, the pressure and temperature of the first flash tank (FL1) were designed to ensure that the vapor phase stream was IPA with a purity of 99.9 mol%. The IPA vapor is further condensed into a liquid phase product by cooling water at 25 °C. Considering that the vacuum degree of FL1 was 10 kPa, the water ring pump could meet the vacuum requirements. The liquid phase stream containing a small amount of IPA and all [Mmim] [DMP] was sent to the second flash tank (FL2) after heating, which required further temperature increase and pressure reduction. Operating temperature was set to 383.15 K to avoid [Mmim] [DMP] decomposition, and operating pressure was set to 350 Pa based on the purity requirement of [Mmim] [DMP]. The top vapor phase stream of the FL2 was sent

Fig. 8. Optimal flowsheet of the improved scheme.

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J. Qi et al. / Journal of Molecular Liquids 293 (2019) 111406

Table 6 Economic comparison of the two schemes. Scheme 1

Diameter (ID)/m Height (H)/m Shell capital cost/k$ Condenser capital cost/k$ Reboiler capital cost/k$ Vacuum capital cost/k$ Condenser duty (QC)/kW Reboiler duty (QR)/kW Total reboiler duty/kW Energy saving/% Other capital cost/k$ Total capital cost/k$ Vacuum operating cost/k$ Utility operating cost/k$ Total operating cost/k$ TAC/k$ ΔTAC/%

Scheme 2

EDC

ERC

EDC

FL1

FL2

0.767 13.176 105.137 55.994 50.859 – −463.667 963.109

1.103 4.392 64.165 57.479 129.570 – −673.080 1404.410

0.767 13.176 105.137 55.994 50.859 – −463.667 963.109

1.437 2.156 12.308 54.314 27.101 2.100 −716.620 441.413 1627.877 31.241 33.528 424.271 15.130 340.956 356.086 497.510 24.582

2.423 3.635 38.382 – 26.338 18.310 – 223.355

2367.519 – 28.194 491.398 – 495.873 495.873 659.672 –

back to the FL1. Flowsheet of the improved scheme and corresponding stream data, operating conditions, energy consumption are displayed in Fig. 8. 4.4. Economic analysis of the two schemes As mentioned above, TAC has been chosen as the objective function of the optimization, whose calculation method is derived from Luyben's literature [36,37]. Corresponding calculation formulas and details (including capital cost calculations and prices and temperatures of utilities) are provided in Table S4. The capital costs only include the columns, heat exchangers, flash tanks and vacuum water ring pump. Since no expensive chilled water was used in this scheme (FL2 steam did not need condensation and FL1 steam only needed cooling water), so only the steam consumption and electricity of vacuum system were taken into account in the operating cost. The annual operating time was set to 8000 h. The results of the economic analysis are listed in Table 6, from which it can be found that compared with scheme 1, scheme 2 reduces energy consumption by 31.241% and TAC by 24.582%. Therefore scheme 2 using flash tanks demonstrates greater application potential. 5. Conclusions Ternary VLE data of DIPE + IPA + EG and DIPE + IPA + [Mmim] [DMP] were measured using a modified Othmer. The NRTL model was introduced to correlate the experimental data and predict the phase behavior of system containing entrainers. y-x diagrams indicate that [Mmim] [DMP] performs stronger molecular force on IPA, which is significantly better than EG in enhancing the relative volatility of DIPE and IPA. After embedding model parameters of NRTL model and physical properties of IL into Aspen Plus, the process simulation of the two extractive distillation alternatives for separating DIPE/IPA was completed. The economic analysis of the two design schemes shows that the scheme 2 using the flash tank has greater application prospect, its energy consumption and TAC are reduced by 31.241% and 24.582%, respectively, compared to scheme 1. Appendix A. Supplementary data Supplementary data to this article can be found online at https://doi. org/10.1016/j.molliq.2019.111406.

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