Isobaric Vapor–Liquid Equilibrium for the Binary Systems of Diethyl

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Isobaric Vapor−Liquid Equilibrium for the Binary Systems of Diethyl Carbonate with Xylene Isomers and Ethylbenzene at 101.33 kPa Jiahai Ding, Wei Guan, Ping Wan, Lei Wang, Hui Wan,* and Guofeng Guan* State Key Laboratory of Materials-Oriented Chemical Engineering, College of Chemical Engineering, Jiangsu National Synergetic Innovation Center for Advanced Materials, Nanjing Tech University, Nanjing 210009, P. R. China S Supporting Information *

ABSTRACT: Isobaric binary vapor−liquid equilibrium (VLE) data for diethyl carbonate with ethylbenzene and xylene isomers are measured at 101.33 kPa by using a modified Rose still. The binary VLE data are tested to be thermodynamically consistent by the Herington method and the point-to-point test of the Fredenslund method. Taking account of the nonideality of the vapor phase, the activity coefficients of the components are calculated. All systems present a positive deviation from ideality. The experimental VLE data are well correlated by the nonrandom two-liquid (NRTL), universal quasichemical activity coefficient (UNIQUAC), and Wilson equations. The calculated vapor-phase compositions and temperature agree well with the experimental values. These experimental data can provide basic thermodynamic data for practical application in developing the distillation simulation of diethyl carbonate with ethylbenzene and xylene isomers.

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INTRODUCTION The term “C8 aromatic mixtures” describes a mixture of oxylene, m-xylene, p-xylene, and ethylbenzene, which are normally derived from petroleum fractions. Ethylbenzene and xylenes are of great importance to the petrochemical industry for their use as a main base for the synthesis of many organic compounds. However, mixed xylenes without separation are generally used as low-value solvents and gasoline additives, and high-purity aromatics are the main body of the market demand, which is used for synthesizing rubber, resins, and fine chemicals. Therefore, the separation of each isomer is of great significance. Although ethylbenzene, m-xylene, and p-xylene do not form an azeotrope, using ordinary distillation to separate these components requires many equilibrium stages and the result is impractical. The reason is that the system has a small boilingpoint range (409 to 412 K) and other similar physical properties.1 It is more habitual to use a solvent via an azetropic distillation system for this separation if high-purity ethylbenzene must be produced. A suitable solvent for this system will be selected from several candidate solvents. In general, one way to select the solvent is through the literature report and experimental research. For example, Berg2 has reported that methyl formate, 1-butanol, and cyclopentanol can be used as agents for separating ethylbenzene C8 aromatic isomer compounds. Li and Gmehling3 have reported that 2-methyl1-butanol can form the binary azeotrope with ethylbenzene or o-, m-, or p-xylene at different pressures. In our previous research, Yang4,5 and Xue6 have measured isobaric VLE data for binary systems (2-methyl-1-butanol, 1-butanol, or methyl formate) with ethylbenzene, o-, m-, or p-xylene at 101.33 kPa to reconfirm the reliability of literature reports. When there is no experimental data in the open published literature, another © XXXX American Chemical Society

main method to select the azeotrope solvent is to use the group contribution model, similar to the UNIFAC model,7 to predict the liquid activity coefficient for this binary pair or the entire component system. For instance, Kuramochi et al.8 have reported that several UNIFAC models are selected as prediction models to represent phase equilibrium used in process design for biodiesel fuel (BDF) production. Nagata et al.9 have reported that the UNIFAC method can successfully predict a better fit of vapor−liquid equilibria for binary alcohol−hydrocarbon systems. Liu et al.10 have reported that the UNIFAC model can extend its application field and can be used to calculate the VLE data of the binary carbonate systems with good accuracy. It is a good way to obtain the vapor−liquid equilibria data quickly and conveniently without too many experiments in many cases. However, compared to the experimental data, the prediction of the UNIFAC method is not always accurate. Therefore, it is necessary to measure the vapor−liquid equilibria data and provide more accurate thermodynamic data in the simulation and design of the distillation process. According to the reported literature,10 the UNIFAC model can be used to calculate the VLE data of the binary carbonate systems with good accuracy, such as methanol−dimethyl carbonate (DMC), methanol−dimethyl oxalate (DMO), DMC−DMO, and DMC−phenol. In this work, the specific function “ternary plot” with the UNIFAC model is employed to predict the vapor−liquid equilibrium of diethyl carbonate + o-, m-, p-xylene, or ethylbenzene. The analysis results in Figures S1−S3 show that diethyl carbonate may be an appropriate Received: January 27, 2016 Accepted: September 27, 2016

A

DOI: 10.1021/acs.jced.6b00085 J. Chem. Eng. Data XXXX, XXX, XXX−XXX

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Table 1. Suppliers and Purity (Mass Fraction) of Chemical Reagents

a

chemical name

source

initial mass fraction purity

diethyl carbonate o-xylene m-xylene p-xylene ethylbenzene

Aladdin Aladdin Aladdin Aladdin Aladdin

0.990 0.990 0.990 0.990 0.990

purification method adsorption adsorption adsorption adsorption adsorption

and and and and and

distillation distillation distillation distillation distillation

final mass fraction purity

analysis method

0.995 0.995 0.995 0.995 0.995

GCa GCa GCa GCa GCa

Gas chromatography.

Table 2. Measured Physical Properties of Pure Compoundsa properties Tb/K (101.3 kPa) ρ/kg·m−3 (298.15K) Antoine constants

exptl lit exptl lit A B C

diethyl carbonate

o-xylene

m-xylene

p-xylene

ethylbenzene

399.84 399.48b 967.83 967.04b 5.8830f 1223.77f −84.304f

417.55 417.55c 875.81 875.57d 6.12644g 1476.39g −59.278g

412.28 412.22c 860.06 860.02d 6.13399g 1462.270g −58.039g

411.43 411.38c 856.54 856.82d 6.11543g 1453.43g −57.840g

409.31 409.31c 862.57 862.63e 6.09070g 1429.55g −59.383g

a Boiling points and liquid densities in comparison with literature data (standard uncertainties u are u(T) = 0.2 K, u(p) = 0.04 kPa, and u(ρ) = 0.4 kg· m−3) and Antoine parameters A, B and C. bReference 16. cReference 17. dReference 18 eReference 19. Antoine equation: log PS/kPa = A − B/(C + T/K). Antoine parameters A, B, and C. fReference 20. gReference 21. Superscripts lit and exptl indicate literature and experimental values, respectively.

both liquid and vapor phases. The experimental procedure and modified equilibrium still are described in detail in the previous work.22 The temperature is measured by a standard mercury thermometer with an uncertainty of 0.2 K. The pressure is kept at 101.33 kPa with an uncertainty of 0.04 kPa by using a pressure control system, which consists of one electromagnetic relay, one vacuum pump, two manometers, two reservoirs, and three triple valves. It takes approximately 3 h to achieve equilibrium when the system temperature is constant, and then the samples of liquid and condensed vapor are taken out and analyzed by GC. Sample Analysis. An SP6800A gas chromatograph with a flame ionization detector (FID) is used to analyze the samples composition. An SE-54 capillary column (20 m × 0.32 mm × 1 μm) and a B-34 capillary column (30 m × 0.25 mm × 0.3 μm) are used, and the FID responses are treated with a Zhejiang Zhida chromatography station. The carrier gas is high-purity nitrogen with a purity of 99.9% and a flow rate of 50 mL/min. The temperatures of the column, injector, and detector of the GC are kept at 383.15, 413.15, and 413.15 K, respectively, and the detector current is set at 120 mA. The required factors determining the compositions from the recorded peak area ratios are obtained using prepared test mixtures of accurately known composition. In each experiment, samples are measured more than three times, and the mean value is recorded. The uncertainty of the measured mole fraction is 0.001.

solvent in separating ethylbenzene from the mixtures containing ethylbenzene, p-xylene, and m-xylene by azeotropic distillation. However, the literature data on this system have been never reported. Therefore, the isobaric VLE data of the four binary systems including diethyl carbonate + o-, m-, pxylene, or ethylbenzene are measured at a pressure of 101.33 kPa for the purpose of reconfirming the validity of predicted experimental VLE data in the UNIFAC method. The binary VLE data are tested to be thermodynamically consistent by the Herington method11 and the point-to-point test of the Fredenslund method12 individually. The experimental data of each binary mixture are correlated by the nonrandom twoliquid (NRTL),13 universal quasichemical activity coefficient (UNIQUAC),14 and Wilson15 equations. Meanwhile, these accurate experimental data may provide basic thermodynamic data for practical application in developing the distillation simulation of diethyl carbonate with ethylbenzene and xylene isomers.

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EXPERIMENTAL SECTION Materials. Diethyl carbonate, ethylbenzene, and xylene isomers purchased from Aladdin Chemistry Co., Ltd. are all analytical reagent (AR) grade materials with a minimum mass fraction purity of 0.990. All chemicals are dried over 0.4 nm molecular sieves for the purpose of removing any free water and then distilled and degassed as described by Fischer and Gmehling for further purification. All reagents are analyzed by gas chromatography (GC), and the minimum purity by mass is found to be 0.995. The suppliers and purity (mass fraction) of chemical reagents are shown in Table 1. The densities of the pure components are measured using an Anton Paar DMA 58 densimeter with the uncertainty of 0.4 kg· m−3. The uncertainty in temperature is 0.2 K. The measured physical properties of pure compounds such as boiling points and liquid densities in comparison with their literature values,17,18 and the Antoine parameters are given in Table 2. Apparatus and Procedure. The isobaric VLE data are obtained by using a modified Rose still with the recirculation of

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RESULTS AND DISCUSSION Experimental Data. To observe the suitability of the experimental procedure, the VLE data of p-xylene + m-xylene are measured at 101.33 kPa in comparison with the data reported in the literature,23 just as shown in Table S1. The results show that the experimental data of this work are in good agreement with the literature, and the apparatus and method for the measurement of vapor−liquid equilibrium data are reliable. The VLE data and calculated activity coefficients for binary systems of diethyl carbonate + ethylbenzene, diethyl carbonate B

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Table 3. Experimental Vapor−Liquid Equilibrium Data for the System Diethyl Carbonate (1) + Ethylbenzene (2) at Temperature T, Liquid Mole Fraction x, Gas Mole Fraction y, and Pressure 101.33 kPaa T/K

x1

y1

γ1

399.84 399.86 399.88 400.02 400.10 400.15 400.55 400.89 401.20 401.50 402.10 402.43 402.99 403.72 404.47 405.28 406.29 407.02 407.73 408.35 409.10 409.31

1.0000 0.9901 0.9712 0.9411 0.8733 0.8321 0.7335 0.6816 0.6457 0.5826 0.5379 0.4859 0.4266 0.3654 0.3001 0.2360 0.1802 0.1327 0.0897 0.0572 0.0183 0.0000

1.0000 0.9907 0.9723 0.9455 0.8824 0.8466 0.7651 0.7198 0.6896 0.6311 0.5987 0.5522 0.4989 0.4361 0.3757 0.3095 0.2446 0.1840 0.1335 0.0868 0.0297 0.0000

1.000 1.004 1.004 1.003 1.007 1.012 1.026 1.029 1.032 1.038 1.049 1.062 1.076 1.079 1.106 1.133 1.152 1.164 1.190 1.208 1.237

Table 5. Experimental Vapor−Liquid Equilibrium Data for the System Diethyl Carbonate (1) + m-Xylene (2) at Temperature T, Liquid Mole Fraction x, Gas Mole Fraction y, and Pressure 101.33 kPaa

γ2

T/K

x1

y1

γ1

1.231 1.228 1.196 1.181 1.166 1.122 1.110 1.095 1.085 1.058 1.052 1.039 1.035 1.018 1.009 1.000 1.001 0.994 0.994 0.994 1.000

399.84 399.87 399.89 399.91 400.25 400.70 400.75 401.05 401.35 401.75 402.26 402.67 403.38 403.79 404.30 405.03 406.56 407.37 407.97 409.41 410.65 412.28

1.0000 0.9582 0.9201 0.8930 0.8540 0.8006 0.7758 0.7399 0.7029 0.6563 0.6015 0.5645 0.5001 0.4764 0.4322 0.3875 0.2848 0.2325 0.1994 0.1273 0.0711 0.0000

1.0000 0.9622 0.9235 0.9052 0.8770 0.8353 0.8145 0.7829 0.7532 0.7229 0.6748 0.6387 0.5879 0.5630 0.5256 0.4854 0.3842 0.3281 0.2898 0.1983 0.1208 0.0000

1.000 1.001 1.003 1.015 1.019 1.022 1.027 1.027 1.031 1.048 1.053 1.050 1.070 1.064 1.079 1.090 1.126 1.153 1.169 1.206 1.223

γ2 1.309 1.269 1.248 1.176 1.138 1.138 1.138 1.123 1.078 1.075 1.080 1.052 1.053 1.039 1.024 1.006 1.000 0.996 0.992 0.988 1.000

a Standard uncertainties u are u(T) = 0.2 K, u(p) = 0.04 kPa, and u(x) = u(y) = 0.001.

a Standard uncertainties u are u(T) = 0.2 K, u(p) = 0.04 kPa, and u(x) = u(y) = 0.001.

Table 4. Experimental Vapor−Liquid Equilibrium Data for the System Diethyl Carbonate (1) + o-Xylene (2) at Temperature T, Liquid Mole Fraction x, Gas Mole Fraction y, and Pressure 101.33 kPaa

Table 6. Experimental Vapor−Liquid Equilibrium Data for the System Diethyl Carbonate (1) + p-Xylene (2) at Temperature T, Liquid Mole Fraction x, Gas Mole Fraction y, and Pressure 101.33 kPaa

T/K

x1

y1

γ1

399.84 399.88 399.99 400.50 400.62 401.03 401.74 403.07 403.58 405.11 405.93 406.34 408.18 409.30 410.22 411.25 412.27 413.19 413.70 414.63 415.54 417.55

1.0000 0.9601 0.9155 0.8625 0.7959 0.7422 0.6851 0.6029 0.5649 0.4700 0.4191 0.3699 0.3113 0.2692 0.2305 0.1932 0.1582 0.1248 0.1095 0.0822 0.0599 0.0000

1.0000 0.9633 0.9352 0.8973 0.8422 0.8014 0.7565 0.7035 0.6642 0.5976 0.5496 0.5028 0.4459 0.4030 0.3581 0.3152 0.2718 0.2254 0.2100 0.1474 0.1077 0.0000

1.000 1.006 1.011 1.015 1.030 1.041 1.051 1.071 1.064 1.104 1.114 1.142 1.145 1.162 1.177 1.203 1.234 1.247 1.251 1.222 1.187

γ2 1.500 1.327 1.246 1.234 1.205 1.175 1.110 1.131 1.065 1.062 1.051 1.035 1.019 1.014 1.003 0.994 0.992 0.981 1.001 0.999 1.000

T/K

x1

y1

γ1

399.84 399.90 400.11 400.54 400.88 401.32 402.91 403.22 403.57 404.09 404.56 405.11 405.52 406.15 406.66 407.17 407.57 408.29 409.32 410.13 411.43

1.0000 0.9682 0.9021 0.8133 0.7525 0.6832 0.5130 0.4770 0.4409 0.4041 0.3671 0.3338 0.3004 0.2665 0.2301 0.2039 0.1801 0.1511 0.0991 0.0633 0.0000

1.0000 0.9701 0.9112 0.8355 0.7837 0.7270 0.5950 0.5640 0.5376 0.5040 0.4622 0.4291 0.4066 0.3617 0.3196 0.2911 0.2609 0.2229 0.1529 0.1026 0.0000

1.000 1.004 1.006 1.011 1.016 1.025 1.069 1.081 1.104 1.113 1.110 1.116 1.162 1.146 1.157 1.173 1.177 1.176 1.197 1.231

γ2 1.296 1.243 1.193 1.172 1.141 1.053 1.046 1.028 1.020 1.027 1.020 0.999 1.007 1.008 1.002 1.003 0.999 0.998 0.995 1.000

a Standard uncertainties u are u(T) = 0.2 K, u(p) = 0.04 kPa, and u(x) = u(y) = 0.001.

a

Standard uncertainties u are u(T) = 0.2 K, u(p) = 0.04 kPa and u(x) = u(y) = 0.001.

+ o-xylene, diethyl carbonate + m-xylene, and diethyl carbonate + p-xylene have been obtained at 101.33 kPa, and the results C

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Table 7. Physical Properties of the Experimental Materialsa,b properties

diethyl carbonate

o-xylene

m-xylene

p-xylene

ethylbenzene

Tc/K Pc/MPa Vc/cm3·mol−1 Vi/cm3·mol−1 ZC ω r q

687.1b 3.59b 406.0b 120.9b 0.255b 0.416b 4.843b 3.812b

630.2c 3.729c 369.0c 121.9c 0.263c 0.314c 4.66c 3.54c

617.0c 3.546c 376.0c 121.9c 0.260c 0.331c 4.66c 3.54c

616.2c 3.516c 379.0c 121.9c 0.260c 0.324c 4.66c 3.54c

617.1c 3.607c 374.0c 45.7c 0.263c 0.301c 4.60c 3.51c

a

TC, critical temperature; PC, critical presure; VC, critical volume; Vi, molar volume parameter of components for the Wilson equation; r, structure volume parameter for the UNIQUAC equation; q, structure area parameter for the UNIQUAC equation; ZC, critical compressibility factor; and ω, acentric factor. bReference 26. cReference 27.

Figure 1. T−x1−y1 diagram for diethyl carbonate (1) + ethylbenzene (2) at 101.33 kPa: ●, experimental liquid-phase mole fractions, x1; ○, experimental vapor-phase mole fractions, y1; , NRTL equation.

Figure 3. T−x1−y1 diagram for diethyl carbonate (1) + m-xylene (2) at 101.33 kPa: ●, experimental liquid-phase mole fractions, x1; ○, experimental vapor-phase mole fractions, y1; , NRTL equation.

Figure 2. T−x1−y1 diagram for diethyl carbonate (1) + o-xylene (2) at 101.33 kPa: ●, experimental liquid-phase mole fractions, x1; ○, experimental vapor-phase mole fractions, y1; , NRTL equation.

Figure 4. T−x1−y1 diagram for diethyl carbonate (1) + p-xylene (2) at 101.33 kPa: ●, experimental liquid-phase mole fractions, x1; ○, experimental vapor-phase mole fractions, y1; , NRTL equation.

are reported in Tables 3 to 6, respectively. The experiment results show that diethyl carbonate, ethylbenzene, and p-xylene cannot form a binary azeotrope either, which reveals that the use of diethyl carbonate as a solvent for the separation of the aromatic compounds by azeotropic distillation is impractical. The Herington method11 is used to check the thermodynamic consistency of experimental data for all binary systems, and the values of |D − J| for the diethyl carbonate +

ethylbenzene, diethyl carbonate + o-xylene, diethyl carbonate + m-xylene, and diethyl carbonate + p-xylene system are 3.41, 2.68, 5.75, and 3.96, respectively, and less than 10, which shows that the binary VLE data of this work are thermodynamically consistent. Furthermore, the well-known point-to-point test of Van Ness,24 which is modified by Fredenslund, is applied to reconfirm the validity of the experimental data. The D

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⎞ ⎛ N P ⎜ ̂ ln ϕi = 2 ∑ y Bij − Bm ⎟⎟ RT ⎜⎝ j = 1 j ⎠

Table 8. Correlation Parameters and the Root-Mean-Square Deviations from the NRTL, UNIQUAC, and Wilson Equations at 101.33 kPa root-mean-squared deviations

model parameters model NRTL UNIQUAC Wilson NRTL UNIQUAC Wilson NRTL UNIQUAC Wilson NRTL UNIQUAC Wilson

A12/J·mol−1a

A21/J·mol−1a

α12

σy1b

Diethyl Carbonate (1) + Ethylbenzene (2) 593.62 151.82 0.3 0.0014 −982.94 840.51 0.0014 −240.70 −511.60 0.0013 Diethyl Carbonate (1) + o-Xylene (2) 3455.42 −1745.07 0.3 0.0062 −2615.32 1808.58 0.0065 1015.94 −2639.42 0.0061 Diethyl Carbonate (1) + m-Xylene (2) 2822.14 −1566.26 0.3 0.0036 −2322.73 1705.16 0.0036 950.29 −2143.30 0.0036 Diethyl Carbonate (1) + p-Xylene (2) 2015.90 −965.67 0.3 0.0042 −1857.75 1431.99 0.0025 447.23 −1442.77 0.0025

where N

σTc/K

Bm =

(4)

Bii and Bjj are the second virial coefficients of the pure components, and Bij is the cross second virial coefficient. The Pitzer and Curl equations as modified by Tsonopoulos25 are used for the calculation of the second virial coefficients. Meanwhile, the pure component vapor pressure (PiS) is calculated with the Antoine equation

0.16 0.18 0.17

log ps /kPa = A −

0.11 0.10 0.099

B C + T /K

(5)

where A, B, and C are expressed as the Antoine constants. The physical properties of the experimental materials required in this calculation are summarized in Table 7. The γi values calculated and the relationships between activity coefficients γ1 and γ2 and mole fractions x1 for diethyl carbonate + ethylbenzene, diethyl carbonate + o-xylene, diethyl carbonate + m-xylene, and diethyl carbonate + p-xylene are shown in Figures S8−S11, respectively. All binary systems present positive deviations from ideality. Correlation of Binary Systems. In the process of the simulation of distillation separation for the system of C8 aromatic hydrocarbon isomers + diethyl carbonate, it is crucial to obtain the model parameters of the pertinent binary system. Using a least-squares procedure developed by Fredenslund et al., the VLE data are correlated together with the NRTL, UNIQUAC, and Wilson models to acquire the binary interaction parameters and the calculated and experimental activity coefficient values of diethyl carbonate + ethylbenzene, diethyl carbonate + o-xylene, diethyl carbonate + m-xylene, and diethyl carbonate + p-xylene by minimizing the following objective function F

0.13 0.065 0.058

The interaction parameters for various models are as follows: NRTL: A12 = (g12 − g22); A21 = (g21 − g11). UNIQUAC: A12 = (U12 − U22), A21 = (U21 − U11). Wilson: A12 = (λ12 − λ22), A21 = (λ21 − λ11). bσy = [∑ni=1(ycal − yexp)/n]1/2. cσT = [∑ni=1(Tcal − Texp)2/n]1/2.

mathematical procedure for the Fredenslund method13 is used to calculate absolute deviations of total pressure p and mole fraction in vapor phase y1 for each point. The test results for binary system diethyl carbonate + ethylbenzene, diethyl carbonate + o-xylene, diethyl carbonate + m-xylene, and diethyl carbonate + p-xylene are reported in Figures S4−S7, respectively. It can be seen that the absolute deviation between y1,exptl and y1,calcd is less than 0.01 and the absolute deviation between pexptl and pcalcd is less than 1.33 kPa (10 mmHg), indicating that the VLE results for the four systems are reliable. Vapor−Liquid Equilibrium Model. Taking account of the nonideality of the vapor phase, the liquid-phase activity coefficients of the components are calculated using eq 1

N

F=

∑ {(γ1exp − γ1cal)k k=1

2

+ (γ2exp − γ2cal) k

2

}

(6)

where γ and N correspond to the activity coefficient and the number of experimental points, respectively. The experimental and calculated vapor-phase and liquid-phase mole composition and equilibrium temperature for each data point with the NRTL model can be seen in Figures 1 to 4. Interaction parameters A12 and A21 for the NRTL, UNIQUAC, and Wilson equations along with the root-mean-squared deviations between the experimental and calculated values of equilibrium temperature and vapor-phase composition are listed in Table 8. It can be observed that the experimental data agree well with data calculated by using NRTL, UNIQUAC, and Wilson models.

(1)

where xi and yi correspond to the liquid and vapor mole fractions of component i in equilibrium; p and pSi refer to the total pressure and the saturated pure component i vapor pressure at system temperature T, respectively; ϕ̂ Vi and ϕSi correspond to the fugacity coefficient of component i in the mixture vapor phase and pure vapor at saturation; γi is the activity coefficient of component i; VLi is the liquid molar volume; R is the universal gas constant; and T is the system temperature. The numerical results of exp[V iL (p − piS )/RT] are approximately equal to 1 at low or moderate pressure, so eq 2 can be expressed as

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CONCLUSIONS The isobaric VLE data for binary systems of diethyl carbonate + o-xylene, diethyl carbonate + m-xylene, diethyl carbonate + pxylene, and diethyl carbonate + ethylbenzene are measured using a modified Rose still at 101.33 kPa. The VLE data are considered to be thermodynamically consistent according to the Herington method and the point-to-point test, individually.

V

ϕî pyi = γixipiS ϕiS

N

∑ ∑ yyi j Bij i=1 j=1

0.13 0.12 0.13

a

⎡ V L(p − pS ) ⎤ V i i ⎥ pyi ϕî = piS ϕiSγixi exp⎢ ⎢⎣ ⎥⎦ RT

(3)

(2)

The vapor-phase fugacity coefficients are calculated using the following virial equation with the second virial coefficient (eq 3) E

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The experimental VLE data are successfully correlated by the NRTL, UNIQUAC, and Wilson equations. All deviations are small, which satisfies the engineering separation design requirement. No azeotropes are observed. The four sets of accurate vapor−liquid equilibrium data at atmospheric pressure for the binary systems can provide basic thermodynamic data for practical application in developing the distillation simulation and optimization of separating diethyl carbonate with ethylbenzene and xylene isomers.

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.jced.6b00085. Residue curve maps in Aspen Plus with the UNIFAC model, results of thermodynamic consistency tests for VLE using the Fredenslund method, and plots of activity coefficients and molar fraction in the liquid phase for different systems (PDF)

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AUTHOR INFORMATION

Corresponding Authors

*(H.W.) E-mail: [email protected]. Tel: +86-25-8358 7198. *(G.G.) E-mail: [email protected]. Tel: +86-25-8358 7198. Funding

We gratefully acknowledge the National Key Technology R&D Program of China (no. 2011BAE05B03), a Project Funded by the Priority Academic Program Development of Jiangsu Higher Education Institutions (PAPD), and the Sinopec Yangzi Petrochemical Company Ltd. Notes

The authors declare no competing financial interest.

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REFERENCES

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DOI: 10.1021/acs.jced.6b00085 J. Chem. Eng. Data XXXX, XXX, XXX−XXX