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Separation of azeotropic mixture isopropyl alcohol + ethyl acetate by extractive distillation: Vapor-liquid equilibrium measurements and interaction exploration
Journal Pre-proof Separation of azeotropic mixture isopropyl alcohol + ethyl acetate by extractive distillation: Vapor-liquid equilibrium measurements and interaction exploration Yi Zhang, Xin Xu, Hui Yang, Jun Gao, Dongmei Xu, Lianzheng Zhang, Yinglong Wang PII:
S0378-3812(19)30490-X
DOI:
https://doi.org/10.1016/j.fluid.2019.112428
Reference:
FLUID 112428
To appear in:
Fluid Phase Equilibria
Received Date: 6 August 2019 Revised Date:
25 November 2019
Accepted Date: 29 November 2019
Please cite this article as: Y. Zhang, X. Xu, H. Yang, J. Gao, D. Xu, L. Zhang, Y. Wang, Separation of azeotropic mixture isopropyl alcohol + ethyl acetate by extractive distillation: Vapor-liquid equilibrium measurements and interaction exploration, Fluid Phase Equilibria (2020), doi: https://doi.org/10.1016/ j.fluid.2019.112428. This is a PDF file of an article that has undergone enhancements after acceptance, such as the addition of a cover page and metadata, and formatting for readability, but it is not yet the definitive version of record. This version will undergo additional copyediting, typesetting and review before it is published in its final form, but we are providing this version to give early visibility of the article. Please note that, during the production process, errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain. © 2019 Published by Elsevier B.V.
Author Contribution Statement Separation of azeotropic mixture isopropyl alcohol + ethyl acetate by extractive distillation: Vapor-liquid equilibrium measurements and interaction exploration Yi Zhang a, Xin Xu a, Hui Yang b, Jun Gao a,*, Dongmei Xu a, Lianzheng Zhang a, Yinglong Wang c a
College of Chemical and Environmental Engineering, Shandong University of Science and Technology, Qingdao 266590, China b School of Biological and Chemical Engineering, Qingdao Technical College, Qingdao 266555, China c College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao 266042, China * Corresponding author
Yi Zhang: Data curation and Writing- Original draft preparation. Xin Xu: Data curation and Correlation. Hui Yang: Validation. Jun Gao: Conceptualization and Methodology Dongmei Xu: Formal analysis. Lianzheng Zhang: Visualization, Investigation. Yinglong Wang: Writing - Review & Editing.
Separation of azeotropic mixture isopropyl alcohol + ethyl acetate by extractive distillation: Vapor-liquid equilibrium measurements and interaction exploration Yi Zhang a, Xin Xu a, Hui Yang b, Jun Gao a,*, Dongmei Xu a, Lianzheng Zhang a, Yinglong Wang c a
College of Chemical and Environmental Engineering, Shandong University of Science and Technology, Qingdao 266590, China b School of Biological and Chemical Engineering, Qingdao Technical College, Qingdao 266555, China c College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao 266042, China * Corresponding author E-mail address: [email protected]
Abstract: Isopropyl alcohol and ethyl acetate can be used to produce degradable and renewable fuel. Since isopropyl alcohol + ethyl acetate can form an azeotropic mixture, it is a tough task to separate the binary mixture by general distillation. In this work, extractive distillation process with N, N-dimethylformamide and dimethyl sulfoxide as entrainers was adopted to separate this azeotrope. The binary and ternary vapor-liquid equilibrium data for (isopropyl alcohol + N, N-dimethylformamide), (ethyl acetate + dimethyl sulfoxide), (isopropyl alcohol + ethyl acetate + N, N-dimethylformamide) and (isopropyl alcohol + ethyl acetate + dimethyl sulfoxide) were determined under 101.3 kPa. Meanwhile, the interaction energies between the molecules were calculated to provide the theoretical insight into the separation of the azeotrope of (EA + IPA) by the entrainers. In addition, the NRTL, UNIQUAC and Wilson models were used to fit the determined binary VLE data. The ternary VLE data for (isopropyl alcohol + ethyl acetate + N, N-dimethylformamide) and (isopropyl alcohol + ethyl acetate + dimethyl sulfoxide) were predicted using the NRTL, UNIQUAC and Wilson models with the parameters regressed from the experimental data. Keywords: Azeotrope; Separation; Vapor-liquid equilibrium; Isopropyl alcohol; Ethyl acetate
1
1
1. Introduction
2
At present, environment protection has attracted more attention from people and
3
governments [1]. Biodiesel is a degradable and renewable fuel, which is becoming
4
more and more attractive in environment and economy [2]. Isopropyl alcohol (IPA)
5
and ethyl acetate (EA), which are important chemicals [3-4], can be used to produce
6
biofuel [5-7]. Generally, IPA can be produced by catalytic hydrogenation of
7
isopropyl acetate [8] with ethyl acetate (EA) as byproduct. However, EA and IPA can
8
form a binary azeotrope with the azeotropic temperature of 348.10 K and azeotropic
9
composition of (EA : IPA= 0.707 : 0.293, mole fraction) at 101.3 kPa [9], which is
10
difficult to separate by conventional distillation. Hence, separation of EA and IPA is
11
important for the energy and chemical industry. To separate such binary azeotropic
12
mixtures, some special separation methods are adopted such as extractive distillation
13
(ED) [10-13], extraction [14-17] and membrane separation [18-22].
14
In this work, ED was applied to separate the binary azeotrope of EA and IPA with
15
N, N-dimethylformamide (DMF) [23] and dimethyl sulfoxide (DMSO) [24] as
16
entrainers. For design and simulation of the ED process for separating IPA and EA,
17
the VLE data contained IPA, EA and entrainers are required. In previous work, the
18
vapor-liquid equilibrium (VLE) for the binary system (EA + IPA) at pressure of
19
101.3 kPa have been determined by Murti and Hernández et al. [9, 25]. Yan et al. [26]
20
and Hong et al. [27] reported the VLE data for (EA + IPA) under the pressures of
21
38.8-229.9 kPa and 101.3-902.0 kPa, respectively. However, the binary and ternary
22
VLE data for (IPA + DMF), (IPA + EA + DMF) and (IPA + EA + DMSO) at 101.3
23
kPa have not been reported in the literatures.
24
Therefore, the binary and ternary VLE data for (IPA + DMF), (EA + DMSO),
25
(IPA + EA + DMF) and (IPA + EA + DMSO) were determined in this work. The
26
Dmol3 module from Materials Studio was adopted to calculate the interaction
27
energies between the components and entrainers to explore the interaction
28
mechanism at molecular level. Meanwhile, the experimental VLE data were fitted by
29
the thermodynamic models of NRTL [28], UNIQUAC [29] and Wilson [30]. In 2
1
addition, the thermodynamic consistency test of van Ness test [31] and Herington test
2
[32] were adopted to check the binary VLE data (IPA + DMF) and (EA + DMSO).
3
2. Experimental
4
2.1 Chemicals
5
The experimental chemicals used in this work were IPA, EA, DMF and DMSO,
6
which were supplied commercially and the purities were checked with gas
7
chromatography (GC). All the chemicals were used directly. The detailed
8
information about the chemicals is presented in Table 1. From Table 1, the boiling
9
temperatures of the pure components at 101.3 kPa agree well with the reference
10
values.
11
Table 1
12
Detailed information of the chemicals. Component
CAS
Suppliers
fraction 67-63-0
Chengdu Kelong Chemical Co., Ltd.
EA
141-78-6
Chengdu Kelong Chemical Co., Ltd
DMF
68-12-2
DMSO
67-68-5
IPA
Tb/Kb
Mass exp
Analysis lit
method [33]
355.34
355.33 355.35 [34]
GCa
0.990
349.90
350.25 [35] 350.15 [36]
GCa
Tian in Fuyu Fine Chemical Co., Ltd.
0.995
424.29
423.91 [37] 425.63 [38]
GCa
Tian in Fuchen Chemical Co., Ltd.
0.990
463.34
463.38 [39] 463.89 [40]
GCa
0.997
13
a
Gas chromatograph.
14
b
The experimental pressure for the measurement of boiling temperature is 101.3 kPa,
15
the standard uncertainties u of p and T are u(p)=0.35 kPa, u(T)=0.35 K.
16
2.2 Apparatus and procedures
17
A Rose-Williams still was used to measure the binary and ternary VLE data for the
18
systems (IPA + DMF), (EA + DMSO), (IPA + EA + DMF) and (IPA + EA +
19
DMSO).
20
The schematic diagram of Rose-Williams still is presented in Fig. 1. As shown in Fig.
21
1, the condensed vapor phase and the liquid phase are continuously recirculated in 3
1
the recirculating still to make the vapor and liquid phases contact fully and establish
2
the equilibrium state as soon as possible. When the fluctuation of temperature was
3
within 0.1 K over 50 min [41], the vapor-liquid equilibrium state was reached. Then,
4
all the samples of vapor and liquid phases were withdrawn at the same time, and
5
were analyzed by GC. The detailed measurements were described in the previous
6
literatures [42-45].
7 8
Fig. 1. Schematic diagram of the equilibrium still: (1) thermometer, (2) glass VLE
9
chamber, (3) liquid sample connection, (4) heating rod, (5) vapor sample connection,
10
(6) condenser.
11
2.3 Analysis
12
Gas chromatography (Lunan SP6890) was employed to analyze the sample
13
compositions of vapor and liquid phases. Table 2 summarized the GC operating
14
conditions. The method of area correction normalization and the analysis details can
15
be found in our previous literatures [46-47].
16
Table 2
17
The analysis conditions for the gas chromatography. Name
Characteristic
Description
Column
Type Specification
Packing column Porapak Q (3 mm × 2 m)
Carrier gas
Type Pressure Temperature
Hydrogen (25 mL/min) 0.09 MPa 483.15 K
Injector
4
Column Detector
Temperature Type Temperature
453.15 K Thermal conductivity detector (TCD) 483.15
1
3. Experimental results and interaction energy calculation
2
3.1 Experimental results
3
The binary and ternary VLE data for the systems (IPA + DMF), (EA + DMSO), (IPA
4
+ EA + DMF) and (IPA + EA + DMSO) were measured at 101.3 kPa and are
5
presented in Tables 3-6 and plotted in Figs. 2-3. For comparison, the predicted VLE
6
values by the UNIFAC model and the VLE data reported by Wang et al. [48] for the
7
system of EA (1) + DMSO (2) were plotted in Fig. 2. As seen from Fig. 2, the VLE
8
data reported by Wang et al. show deviations at the mole fraction less than 0.2, which
9
may be due to the difference the purity of reagents and analysis conditions.
10 11 12
13 14
Table 3 Experimental isobaric VLE data for the binary system EA (1) + DMSO (2) and activity coefficient (γ) at 101.3 kPa. a T/K
x1
y1
γ1
γ2
353.67 359.48 365.90 371.91 377.19 386.07 391.59 395.19 401.43 407.67 413.19 419.67 425.91 431.43 437.43 443.19 449.67 457.35 460.47
0.891 0.690 0.477 0.349 0.271 0.196 0.148 0.128 0.104 0.086 0.071 0.057 0.045 0.036 0.025 0.018 0.012 0.005 0.003
0.994 0.983 0.972 0.959 0.947 0.923 0.901 0.886 0.855 0.822 0.783 0.731 0.667 0.602 0.518 0.426 0.319 0.158 0.084
1.001 1.063 1.253 1.418 1.552 1.652 1.850 1.928 1.966 1.992 1.999 2.025 2.066 2.093 2.314 2.329 2.330 2.348 2.398
2.962 2.080 1.501 1.346 1.265 1.162 1.130 1.116 1.096 1.057 1.048 1.032 1.030 1.029 1.028 1.027 1.010 1.005 1.004
a
Standard uncertainties u of T, p, x and y are u(T)=0.35 K, u(p)=0.35 kPa, u(x1)=u(y1)=0.008. 5
1 2 3
4 5
Table 4 Experimental isobaric VLE data for the binary system IPA (1) + DMF (2) and activity coefficient (γ) at 101.3 kPa. a T/K
x1
y1
γ1
γ2
356.30 360.35 365.00 368.61 372.49 376.70 380.76 384.65 389.00 392.31 396.50 400.85 404.77 408.50 411.53 416.28 421.32 424.29
0.970 0.835 0.713 0.630 0.558 0.481 0.415 0.357 0.302 0.262 0.217 0.173 0.139 0.108 0.087 0.054 0.028 0.006
0.998 0.985 0.967 0.949 0.926 0.893 0.856 0.816 0.767 0.724 0.661 0.589 0.517 0.440 0.363 0.246 0.139 0.032
0.985 0.964 0.930 0.904 0.868 0.839 0.812 0.791 0.766 0.752 0.730 0.717 0.700 0.689 0.644 0.614 0.605 0.596
0.707 0.747 0.804 0.841 0.872 0.927 0.958 0.973 0.980 0.983 0.996 0.996 0.998 0.998 0.999 0.999 1.000 1.001
a
Standard uncertainties u of T, p, x and y are u(T)=0.35 K, u(p)=0.35 kPa, u(x1)=u(y1)=0.004.
6
Table 5
7
Experimental isobaric VLE data for the ternary system EA (1) + IPA (2) + DMSO (3) and the values of RMSD for NRTL, UNIQUAC and Wilson at 101.3 kPa. a
8
T/K
Liquid phase x1
380.39 379.04 377.70 376.43 375.18 374.12 373.08 372.23 371.44 9 10
0.090 0.198 0.301 0.400 0.503 0.599 0.701 0.799 0.909
x2 0.910 0.802 0.699 0.600 0.497 0.401 0.299 0.201 0.091
Vapor phase y1 0.165 0.320 0.448 0.550 0.646 0.729 0.807 0.874 0.945
y2
NRTL
UNIQUAC
Wilson
RMSD
RMSD
RMSD
(y1)
(T)
(y1)
Solvent to feed ratio=2:1 0.835 0.010 0.01 0.054 0.680 0.017 0.05 0.077 0.552 0.018 0.15 0.099 0.450 0.012 0.28 0.107 0.354 0.011 0.44 0.108 0.271 0.009 0.55 0.091 0.193 0.007 0.69 0.071 0.126 0.003 0.77 0.039 0.055 0.002 0.89 0.005
a
(T)
(y1)
0.66 0.11 0.64 0.28 0.17 0.35 0.06 0.24 0.52
0.013 0.020 0.024 0.026 0.027 0.028 0.033 0.030 0.032
α12
(T) 0.46 0.83 0.39 0.12 0.03 0.02 0.13 0.33 0.59
2.01 1.90 1.88 1.84 1.80 1.79 1.78 1.75 1.73
Standard uncertainties u of T, p, x and y are u(T)=0.35 K, u(p)=0.35 kPa, u(x1)=0.008, u(y1)=0.006. 6
1
Table 6
2
Experimental isobaric VLE data for the ternary system EA (1) + IPA (2) + DMF (3)
3
and the values of RMSD for NRTL, UNIQUAC and Wilson at 101.3 kPa. a T/K
Liquid phase x1
379.18 378.51 377.88 377.31 376.79 376.34 375.96 375.66 375.44
0.100 0.199 0.301 0.395 0.501 0.600 0.702 0.803 0.901
x2 0.900 0.801 0.699 0.605 0.499 0.400 0.298 0.197 0.099
Vapor phase y1 0.136 0.249 0.365 0.454 0.559 0.650 0.743 0.833 0.917
4
a
5
u(x1)=0.008, u(y1)=0.006.
y2
NRTL
UNIQUAC
Wilson
RMSD
RMSD
RMSD
(y1)
(T)
(y1)
Solvent to feed ratio=2.5:1 0.864 0.008 0.76 0.043 0.751 0.010 0.07 0.076 0.635 0.015 0.02 0.095 0.546 0.006 0.16 0.112 0.441 0.012 0.32 0.107 0.350 0.008 0.42 0.101 0.257 0.010 0.46 0.083 0.167 0.010 0.43 0.060 0.083 0.006 0.34 0.033
(T)
(y1)
0.39 0.28 0.18 0.81 0.28 0.17 0.73 0.53 0.41
0.042 0.074 0.093 0.109 0.104 0.099 0.082 0.059 0.033
α12
(T) 0.34 0.15 0.86 0.77 0.59 0.41 0.26 0.13 0.04
1.41 1.34 1.33 1.27 1.26 1.23 1.22 1.22 1.21
Standard uncertainties u of T, p, x and y are u(T)=0.35 K, u(p)=0.35 kPa,
6
7 8
Fig. 2. T-x-y diagram for the binary system EA (1) + DMSO (2) at 101.3 kPa; ■, T-x,
9
experimental values; ●, T-y, experimental values; ▲, T-x, reference values [48]; ▼,
10
T-y, reference values [48]; —, calculated by the NRTL model; —, predicted by the
11
UNIFAC model.
7
1 2
Fig. 3. T-x-y diagram for the binary system IPA (1) + DMF (2) at 101.3 kPa; ■, T-x,
3
experimental values; ●, T-y, experimental values; —, calculated by the NRTL model.
4
3.2 Effects of entrainers
5 6
To compare the suitability of the entrainers, the relative volatility (α12) of EA (1) to IPA (2) was used in this work and is defined as follows: y1
α12=y2
7
x1
(1)
x2
8
The diagram of the relative volatility for the system (EA + IPA), which are
9
presented in Fig. 4. For comparison of the relative volatility, the reference value [25]
10
of EA (1) + IPA (2) was also plotted in Fig. 4. As can be seen from Fig. 3, the α12
11
values with the added entrainers are all greater than 1. Furthermore, the values of α12
12
for the system EA (1) + IPA (2) with DMSO as entrainer are greater than those with
13
DMF, which is consistent with the calculation results of interaction energy in section
14
3.3.
8
1 2
5
Fig. 4. Relative volatility vs. x1 plot: ■, the reference value [23] of EA (1) + IPA (2); ●, experimental value for the system EA (1) + IPA (2) with the entrainer DMSO; ▲, experimental value for the system EA (1) + IPA (2) with the entrainer DMF; —, fitted line.
6
Also, the x-y plot for the system EA (1) + IPA (2) with two different entrainers
7
(DMF and DMSO) are presented in Fig. 5. As shown in Fig. 5, the azeotropic point is
8
eliminated with adding the entainers. With DMSO as the entrainer, there is a bigger
9
deviation from the x-y diagonal. Therefore, DMSO was selected as a suitable
3 4
10
11 12 13 14 15
entrainer to design the ED process for separating the binary mixture of EA and IPA.
Fig. 5. Influence on VLE for the system EA (1) + IPA (2) with different entrainers. ●, the reference value [9]; ■, the reference value [25]; —, calculated value by Aspen plus; ▲, experimental value (DMF as the extractive solvent); ■, experimental value (DMSO as the extractive solvent); —, fitted lines. 9
1
3.3 Interaction energy calculation
2
For exploring the interactions between molecules, the interaction energy, which is
3
the energy difference between the different entrainers (DMF and DMSO) and the
4
components (IPA and EA) was calculated using the DMol 3 module incorporated the
5
density functional theory. The detailed calculation was provided in our previous work
6
[49-50]. The optimized geometries and calculated interaction energies for the
7
entrainers and the components are shown in Fig. 6 and are summarized in Table 7,
8
respectively.
9
As shown in Table 7, the interaction energy values of the systems (IPA + DMF)
10
and (IPA + DMSO) are greater than that of the system (EA + IPA), which indicates
11
that the entrainers of DMF and DMSO can enhance the relative volatility of EA to
12
IPA. Compared with IPA + DMF, EA + DMF, EA + DMSO and IPA + EA, the
13
interaction energy for IPA + DMSO is the largest value. In the meantime, the
14
interaction energies between IPA and the entrainers are higher than those between
15
EA and the entrainers, which indicates that EA can be separated from the top of the
16
extractive distillation column.
17
18 19
EA+DMF
EA+DMSO
20
10
IPA+DMF
1
IPA+DMSO
2
EA + IPA
3
Fig. 6. Optimized geometries for the systems (EA + DMF), (EA + DMSO), (IPA +
4
DMF), (IPA + DMSO) and (IPA + EA).
5
Table 7
6
Interaction energies and bond lengths for the systems (EA + DMF), (EA + DMSO)
7
(IPA + DMF), (IPA + DMSO), and (IPA + EA). System
Bond lengths (nm)
Interaction energies (kJ·mol-1)
IPA+DMF EA+DMF IPA+DMSO EA+DMSO IPA+EA
0.1976 0.3141 0.1921 0.3404 0.1965
-21.5942 -1.5728 -22.9960 -3.3084 -18.5746
8
4. VLE calculation and regression
9
4.1 VLE calculation
10 11
The VLE thermodynamic relationship can be expressed by eq (2) [51]: Pyiφi =
φsi Psi exp and
VLi P-Psi RT
(2)
refer to liquid composition, vapor composition,
12
where xi, yi, T, P,
13
temperature, vapor pressure, fugacity coefficients at system pressure and saturation
14
pressure. The exponential part is Poynting factor, which is close to unity at 101.3 kPa. 11
1
Also, the fugacity coefficients are approximately close to unity at 101.3 kPa. So, eq
2
(2) can be simplified and expressed as:
3
γi=
Pyi
(3)
Psi xi
4
where Psi represents the pure compound vapor pressure and can be calculated by the
5
extended Antoine equation: ln(Psi /
6
)=C1i+
C2i
T/K+C3i
+C4i(T/K)+C5iln(T/K)+C6i( / )
C8i ≤T/K≤C9i
(4)
7
The parameter values of all the pure components are listed in Table 8, which were
8
taken from the Aspen databank [52]. The values of activity coefficient for the binary
9
mixtures were calculated and are listed in Tables 3-4.
10
Table 8
11
Parameters of the extended Antoine equation a Component EA DMF IPA DMSO
C 2i
C1i 59.92 75.85 103.81 49.37
C3i
-6227.60 -7955.50 -9040.00 -7620.60
C4i
0 0 0 0
C5i
0 0 0 0
12
a
13
4.2 Thermodynamic consistency test
-6.41 -8.80 -12.68 -4.63
C6i(×10-6) 1.79×10 4.24 5.54 0.44
-11
C7i
C8i /K
C9i /K
6 2 2 2
189.60 212.72 185.26 291.67
523.30 649.60 508.30 729.00
Taken from Aspen property databank. [52]
14
The van Ness test [31] and Herington test [32] were adopted to check the
15
thermodynamic consistency of the experimental VLE data. The van Ness test is
16
expressed as follows: !
17
∆y= ∑ & 100
18
∆P= ∑ & 100 '
(
)*+
(
−
#
%$,(-./
)*+
'
(5) (6)
19
where N represents the total number of experimental points; the superscripts “exp”
20
and “cal” are experimental and calculated values, respectively. In this test, the criteria
21
are the values of ∆y and ∆P less than 1.
22 23
The Herington method can be expressed as follows: (1,2)
D=1000(132)0
(7) 12
45.* ,45 6 0 45 6
J=1500
1
(8)
2
where A and B are the total area of ln(γ1/γ2 ) vs. x above the zero line and below the
3
zero line, respectively. Tmax and Tmin represent the maximum and minimum boiling
4
temperatures, respectively. If the value of |8 − 9| < 10, the measured VLE data
5
pass the thermodynamic consistency test.
6 7
The results of ∆y, ∆P and |8 − 9| are summarized in Table 9. As listed in Table 9,
all the values of ∆y and ∆P are less than unity. The values of |8 − 9| are less than
8
10. All the results indicate that the measured VLE data for the mixtures passed the
9
thermodynamic consistency test.
10 11
12
Table 9 Results of the thermodynamic consistency System
D
J
|8 − 9|
∆P
∆y
EAT+ DMSO
45.15
48.40
3.25
0.03
0.80
IPA+ DMF
39.29
29.97
9.32
0.06
0.35
4.3 VLE data regression
13
Based on the maximum likelihood principle [53-54], the determined binary VLE
14
data for the systems (EA + DMSO) and (IPA + DMF) were correlated by the NRTL,
15
UNIQUAC and Wilson models. The objective function is expressed as:
16
17 18 19 20 21
22
(
F=∑ FG
)*+
,(-./
HI
K
J +G
4
)*+
,4 -./
HM
K
J +G
)*+
K , -./
H*
J +G
N
)*+
,N -./
HO
K
J P
(9)
where T, P and Q refer to temperature, pressure and standard deviation at the equilibrium state. The values of standard deviation are as follows: Q( , 0.35 kPa; Q4 , 0.35 K; Q , 0.008; QN , 0.006.
For the Wilson model, the molar volumes of the pure components were calculated by the Rackett equation [55-56]. The Rackett equation is expressed as follows: R ∗,% =
U4- VW
_ ) ∗,XY [\([]M^
Z
(10)
(-
#
23
where Tr = T/Tci, T refers to temperature, Tci denotes critical temperature,
24
critical pressure, ` ∗,U1 is critical compressibility factor. All the parameters of the
25
Rackett equation were taken from Aspen plus and are listed in Table 10. 13
is
1
Table 10
2
Parameters of the Rackett equation for the components of EA, IPA, DMF and
3
DMSO. Component
Tci / K
EA IPA DMF DMSO
523.3 508.3 649.6 729.0
#
Zi*RA
/ kPa
3880.0 4765.0 4420.0 5650.0
0.26 0.25 0.21 0.21
4
The values of the binary interaction parameters of the NRTL, UNIQUAC and
5
Wilson models and the root-mean-square deviation (RMSD) are summarized in Table
6
11. From Table 11, the maximum RMSD values in liquid phase mole fraction (x1),
7
vapor phase mole fraction (y1), pressure (P) and temperature (T) are 0.0003, 0.0095,
8
0.54 K and 0.0027 kPa, respectively. It is concluded that three models can correlate
9
well the experimental VLE data. Also, for the ternary systems (IPA + EA + DMF)
10
and (IPA + EA + DMSO), the NRTL model with the regressed parameters was used
11
to predict the VLE data for the ternary mixtures. As shown in Tables 5-6, the values
12
of root-mean-square deviation of vapor mole fraction for the ternary mixtures are less
13
than 0.018, which indicates that the NRTL model with the regressed parameters can
14
be applied for the extractive distillation process design and optimization.
15 16
Table 11 Regressed parameters of the NRTL models and root-mean-square deviation (RMSD) of the systems.
17
Model
Parameters aij
e
18
aji
NRTL UNIQUAC f Wilson g
-2.3808 20.1761 13.7374
2.1266 -1.2389 6.3205
NRTL UNIQUAC Wilson
1.3430 0.4269 -13.0633
-6.4791 -7.1437 5.2320
a
RMSD(x1)=a∑
RMSD
bij/K
bji/K
EA+ DMSO 1002.95 -502.06 -1536.07 295.07 3009.14 -8042.17 IPA+ DMF -646.96 2652.42 -1218.49 2520.70 4081.92 -1714.87
_ )*+ , [-./ [
14
x1
a
y1
b
T/K c
P/ kPa d
0.0002 0.0003 0.0005
0.0095 0.0087 0.0034
0.54 0.52 0.37
0.0014 0.0026 0.0042
0.0003 0.0007 0.0006
0.0046 0.0081 0.0072
0.24 0.18 0.20
0.0027 0.0008 0.0008
)*+
RMSD(y1)=a∑
1
b
2
c
3
d
4
e
5
f
6
g
7
5. Conclusions
N[
RMSD(T)=a∑
RMSD(P)=a∑
NRTL,b c =
4
)*+
(
c
,4 -./
)*+
_
,N[-./
,(-./
_ _
+ e c / , the value of alpha ij was set at 0.3.
UNIQUAC, b c = f g Wilson, lnh c =
c
c
+ e c/
+ e c/
8
For separation of the binary azeotrope (EA + IPA) by extractive distillation, DMF
9
and DMSO were selected as the entrainers. The binary and ternary VLE data for (IPA
10
+ DMF), (EA + DMSO), (IPA + EA + DMF) and (IPA + EA + DMSO) were
11
determined. The experimental binary VLE data were fitted by the thermodynamic
12
model of NRTL. The thermodynamic consistency was validated by the van Ness test
13
and Herington test for the binary VLE data (IPA + DMF) and (EA + DMSO).
14
Moreover, the interaction energies for the different entrainers (DMF and DMSO) and
15
the components (IPA and EA) were calculated. The results indicate that the
16
interaction energies for (IPA + DMF) and (IPA + DMSO) are greater than that of (EA
17
+ IPA), which means DMSO is the suitable entrainer for separation of the azeotrope
18
of (EA +IPA) by extractive distillation. Furthermore, the ternary VLE data for (IPA +
19
EA + DMF) and (IPA + EA + DMSO) were predicted using the NRTL, UNIQUAC
20
and Wilson models with the regressed parameters and compared with those measured
21
in this work. The values of RMSD of vapor mole fraction are less than 0.018, 0.112
22
and 0.109 for the NRTL, UNIQUAC and Wilson models, which indicates that the
23
regressed parameters for the NRTL model can be applied for the extractive
24
distillation design and optimization.
15
1
Acknowledgement
2
The authors are grateful for the financial support of the National Natural Science
3
Foundation of China (21878178), Shandong Provincial Key Research &
4
Development Project (2018GGX107001), Project of Shandong Province Higher
5
Educational Science and Technology Program (J18KA072).
6 7
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Declaration of interests ☒ The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper. ☐The authors declare the following financial interests/personal relationships which may be considered as potential competing interests:
S0378-3812(19)30490-X
DOI:
https://doi.org/10.1016/j.fluid.2019.112428
Reference:
FLUID 112428
To appear in:
Fluid Phase Equilibria
Received Date: 6 August 2019 Revised Date:
25 November 2019
Accepted Date: 29 November 2019
Please cite this article as: Y. Zhang, X. Xu, H. Yang, J. Gao, D. Xu, L. Zhang, Y. Wang, Separation of azeotropic mixture isopropyl alcohol + ethyl acetate by extractive distillation: Vapor-liquid equilibrium measurements and interaction exploration, Fluid Phase Equilibria (2020), doi: https://doi.org/10.1016/ j.fluid.2019.112428. This is a PDF file of an article that has undergone enhancements after acceptance, such as the addition of a cover page and metadata, and formatting for readability, but it is not yet the definitive version of record. This version will undergo additional copyediting, typesetting and review before it is published in its final form, but we are providing this version to give early visibility of the article. Please note that, during the production process, errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain. © 2019 Published by Elsevier B.V.
Author Contribution Statement Separation of azeotropic mixture isopropyl alcohol + ethyl acetate by extractive distillation: Vapor-liquid equilibrium measurements and interaction exploration Yi Zhang a, Xin Xu a, Hui Yang b, Jun Gao a,*, Dongmei Xu a, Lianzheng Zhang a, Yinglong Wang c a
College of Chemical and Environmental Engineering, Shandong University of Science and Technology, Qingdao 266590, China b School of Biological and Chemical Engineering, Qingdao Technical College, Qingdao 266555, China c College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao 266042, China * Corresponding author
Yi Zhang: Data curation and Writing- Original draft preparation. Xin Xu: Data curation and Correlation. Hui Yang: Validation. Jun Gao: Conceptualization and Methodology Dongmei Xu: Formal analysis. Lianzheng Zhang: Visualization, Investigation. Yinglong Wang: Writing - Review & Editing.
Separation of azeotropic mixture isopropyl alcohol + ethyl acetate by extractive distillation: Vapor-liquid equilibrium measurements and interaction exploration Yi Zhang a, Xin Xu a, Hui Yang b, Jun Gao a,*, Dongmei Xu a, Lianzheng Zhang a, Yinglong Wang c a
College of Chemical and Environmental Engineering, Shandong University of Science and Technology, Qingdao 266590, China b School of Biological and Chemical Engineering, Qingdao Technical College, Qingdao 266555, China c College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao 266042, China * Corresponding author E-mail address: [email protected]
Abstract: Isopropyl alcohol and ethyl acetate can be used to produce degradable and renewable fuel. Since isopropyl alcohol + ethyl acetate can form an azeotropic mixture, it is a tough task to separate the binary mixture by general distillation. In this work, extractive distillation process with N, N-dimethylformamide and dimethyl sulfoxide as entrainers was adopted to separate this azeotrope. The binary and ternary vapor-liquid equilibrium data for (isopropyl alcohol + N, N-dimethylformamide), (ethyl acetate + dimethyl sulfoxide), (isopropyl alcohol + ethyl acetate + N, N-dimethylformamide) and (isopropyl alcohol + ethyl acetate + dimethyl sulfoxide) were determined under 101.3 kPa. Meanwhile, the interaction energies between the molecules were calculated to provide the theoretical insight into the separation of the azeotrope of (EA + IPA) by the entrainers. In addition, the NRTL, UNIQUAC and Wilson models were used to fit the determined binary VLE data. The ternary VLE data for (isopropyl alcohol + ethyl acetate + N, N-dimethylformamide) and (isopropyl alcohol + ethyl acetate + dimethyl sulfoxide) were predicted using the NRTL, UNIQUAC and Wilson models with the parameters regressed from the experimental data. Keywords: Azeotrope; Separation; Vapor-liquid equilibrium; Isopropyl alcohol; Ethyl acetate
1
1
1. Introduction
2
At present, environment protection has attracted more attention from people and
3
governments [1]. Biodiesel is a degradable and renewable fuel, which is becoming
4
more and more attractive in environment and economy [2]. Isopropyl alcohol (IPA)
5
and ethyl acetate (EA), which are important chemicals [3-4], can be used to produce
6
biofuel [5-7]. Generally, IPA can be produced by catalytic hydrogenation of
7
isopropyl acetate [8] with ethyl acetate (EA) as byproduct. However, EA and IPA can
8
form a binary azeotrope with the azeotropic temperature of 348.10 K and azeotropic
9
composition of (EA : IPA= 0.707 : 0.293, mole fraction) at 101.3 kPa [9], which is
10
difficult to separate by conventional distillation. Hence, separation of EA and IPA is
11
important for the energy and chemical industry. To separate such binary azeotropic
12
mixtures, some special separation methods are adopted such as extractive distillation
13
(ED) [10-13], extraction [14-17] and membrane separation [18-22].
14
In this work, ED was applied to separate the binary azeotrope of EA and IPA with
15
N, N-dimethylformamide (DMF) [23] and dimethyl sulfoxide (DMSO) [24] as
16
entrainers. For design and simulation of the ED process for separating IPA and EA,
17
the VLE data contained IPA, EA and entrainers are required. In previous work, the
18
vapor-liquid equilibrium (VLE) for the binary system (EA + IPA) at pressure of
19
101.3 kPa have been determined by Murti and Hernández et al. [9, 25]. Yan et al. [26]
20
and Hong et al. [27] reported the VLE data for (EA + IPA) under the pressures of
21
38.8-229.9 kPa and 101.3-902.0 kPa, respectively. However, the binary and ternary
22
VLE data for (IPA + DMF), (IPA + EA + DMF) and (IPA + EA + DMSO) at 101.3
23
kPa have not been reported in the literatures.
24
Therefore, the binary and ternary VLE data for (IPA + DMF), (EA + DMSO),
25
(IPA + EA + DMF) and (IPA + EA + DMSO) were determined in this work. The
26
Dmol3 module from Materials Studio was adopted to calculate the interaction
27
energies between the components and entrainers to explore the interaction
28
mechanism at molecular level. Meanwhile, the experimental VLE data were fitted by
29
the thermodynamic models of NRTL [28], UNIQUAC [29] and Wilson [30]. In 2
1
addition, the thermodynamic consistency test of van Ness test [31] and Herington test
2
[32] were adopted to check the binary VLE data (IPA + DMF) and (EA + DMSO).
3
2. Experimental
4
2.1 Chemicals
5
The experimental chemicals used in this work were IPA, EA, DMF and DMSO,
6
which were supplied commercially and the purities were checked with gas
7
chromatography (GC). All the chemicals were used directly. The detailed
8
information about the chemicals is presented in Table 1. From Table 1, the boiling
9
temperatures of the pure components at 101.3 kPa agree well with the reference
10
values.
11
Table 1
12
Detailed information of the chemicals. Component
CAS
Suppliers
fraction 67-63-0
Chengdu Kelong Chemical Co., Ltd.
EA
141-78-6
Chengdu Kelong Chemical Co., Ltd
DMF
68-12-2
DMSO
67-68-5
IPA
Tb/Kb
Mass exp
Analysis lit
method [33]
355.34
355.33 355.35 [34]
GCa
0.990
349.90
350.25 [35] 350.15 [36]
GCa
Tian in Fuyu Fine Chemical Co., Ltd.
0.995
424.29
423.91 [37] 425.63 [38]
GCa
Tian in Fuchen Chemical Co., Ltd.
0.990
463.34
463.38 [39] 463.89 [40]
GCa
0.997
13
a
Gas chromatograph.
14
b
The experimental pressure for the measurement of boiling temperature is 101.3 kPa,
15
the standard uncertainties u of p and T are u(p)=0.35 kPa, u(T)=0.35 K.
16
2.2 Apparatus and procedures
17
A Rose-Williams still was used to measure the binary and ternary VLE data for the
18
systems (IPA + DMF), (EA + DMSO), (IPA + EA + DMF) and (IPA + EA +
19
DMSO).
20
The schematic diagram of Rose-Williams still is presented in Fig. 1. As shown in Fig.
21
1, the condensed vapor phase and the liquid phase are continuously recirculated in 3
1
the recirculating still to make the vapor and liquid phases contact fully and establish
2
the equilibrium state as soon as possible. When the fluctuation of temperature was
3
within 0.1 K over 50 min [41], the vapor-liquid equilibrium state was reached. Then,
4
all the samples of vapor and liquid phases were withdrawn at the same time, and
5
were analyzed by GC. The detailed measurements were described in the previous
6
literatures [42-45].
7 8
Fig. 1. Schematic diagram of the equilibrium still: (1) thermometer, (2) glass VLE
9
chamber, (3) liquid sample connection, (4) heating rod, (5) vapor sample connection,
10
(6) condenser.
11
2.3 Analysis
12
Gas chromatography (Lunan SP6890) was employed to analyze the sample
13
compositions of vapor and liquid phases. Table 2 summarized the GC operating
14
conditions. The method of area correction normalization and the analysis details can
15
be found in our previous literatures [46-47].
16
Table 2
17
The analysis conditions for the gas chromatography. Name
Characteristic
Description
Column
Type Specification
Packing column Porapak Q (3 mm × 2 m)
Carrier gas
Type Pressure Temperature
Hydrogen (25 mL/min) 0.09 MPa 483.15 K
Injector
4
Column Detector
Temperature Type Temperature
453.15 K Thermal conductivity detector (TCD) 483.15
1
3. Experimental results and interaction energy calculation
2
3.1 Experimental results
3
The binary and ternary VLE data for the systems (IPA + DMF), (EA + DMSO), (IPA
4
+ EA + DMF) and (IPA + EA + DMSO) were measured at 101.3 kPa and are
5
presented in Tables 3-6 and plotted in Figs. 2-3. For comparison, the predicted VLE
6
values by the UNIFAC model and the VLE data reported by Wang et al. [48] for the
7
system of EA (1) + DMSO (2) were plotted in Fig. 2. As seen from Fig. 2, the VLE
8
data reported by Wang et al. show deviations at the mole fraction less than 0.2, which
9
may be due to the difference the purity of reagents and analysis conditions.
10 11 12
13 14
Table 3 Experimental isobaric VLE data for the binary system EA (1) + DMSO (2) and activity coefficient (γ) at 101.3 kPa. a T/K
x1
y1
γ1
γ2
353.67 359.48 365.90 371.91 377.19 386.07 391.59 395.19 401.43 407.67 413.19 419.67 425.91 431.43 437.43 443.19 449.67 457.35 460.47
0.891 0.690 0.477 0.349 0.271 0.196 0.148 0.128 0.104 0.086 0.071 0.057 0.045 0.036 0.025 0.018 0.012 0.005 0.003
0.994 0.983 0.972 0.959 0.947 0.923 0.901 0.886 0.855 0.822 0.783 0.731 0.667 0.602 0.518 0.426 0.319 0.158 0.084
1.001 1.063 1.253 1.418 1.552 1.652 1.850 1.928 1.966 1.992 1.999 2.025 2.066 2.093 2.314 2.329 2.330 2.348 2.398
2.962 2.080 1.501 1.346 1.265 1.162 1.130 1.116 1.096 1.057 1.048 1.032 1.030 1.029 1.028 1.027 1.010 1.005 1.004
a
Standard uncertainties u of T, p, x and y are u(T)=0.35 K, u(p)=0.35 kPa, u(x1)=u(y1)=0.008. 5
1 2 3
4 5
Table 4 Experimental isobaric VLE data for the binary system IPA (1) + DMF (2) and activity coefficient (γ) at 101.3 kPa. a T/K
x1
y1
γ1
γ2
356.30 360.35 365.00 368.61 372.49 376.70 380.76 384.65 389.00 392.31 396.50 400.85 404.77 408.50 411.53 416.28 421.32 424.29
0.970 0.835 0.713 0.630 0.558 0.481 0.415 0.357 0.302 0.262 0.217 0.173 0.139 0.108 0.087 0.054 0.028 0.006
0.998 0.985 0.967 0.949 0.926 0.893 0.856 0.816 0.767 0.724 0.661 0.589 0.517 0.440 0.363 0.246 0.139 0.032
0.985 0.964 0.930 0.904 0.868 0.839 0.812 0.791 0.766 0.752 0.730 0.717 0.700 0.689 0.644 0.614 0.605 0.596
0.707 0.747 0.804 0.841 0.872 0.927 0.958 0.973 0.980 0.983 0.996 0.996 0.998 0.998 0.999 0.999 1.000 1.001
a
Standard uncertainties u of T, p, x and y are u(T)=0.35 K, u(p)=0.35 kPa, u(x1)=u(y1)=0.004.
6
Table 5
7
Experimental isobaric VLE data for the ternary system EA (1) + IPA (2) + DMSO (3) and the values of RMSD for NRTL, UNIQUAC and Wilson at 101.3 kPa. a
8
T/K
Liquid phase x1
380.39 379.04 377.70 376.43 375.18 374.12 373.08 372.23 371.44 9 10
0.090 0.198 0.301 0.400 0.503 0.599 0.701 0.799 0.909
x2 0.910 0.802 0.699 0.600 0.497 0.401 0.299 0.201 0.091
Vapor phase y1 0.165 0.320 0.448 0.550 0.646 0.729 0.807 0.874 0.945
y2
NRTL
UNIQUAC
Wilson
RMSD
RMSD
RMSD
(y1)
(T)
(y1)
Solvent to feed ratio=2:1 0.835 0.010 0.01 0.054 0.680 0.017 0.05 0.077 0.552 0.018 0.15 0.099 0.450 0.012 0.28 0.107 0.354 0.011 0.44 0.108 0.271 0.009 0.55 0.091 0.193 0.007 0.69 0.071 0.126 0.003 0.77 0.039 0.055 0.002 0.89 0.005
a
(T)
(y1)
0.66 0.11 0.64 0.28 0.17 0.35 0.06 0.24 0.52
0.013 0.020 0.024 0.026 0.027 0.028 0.033 0.030 0.032
α12
(T) 0.46 0.83 0.39 0.12 0.03 0.02 0.13 0.33 0.59
2.01 1.90 1.88 1.84 1.80 1.79 1.78 1.75 1.73
Standard uncertainties u of T, p, x and y are u(T)=0.35 K, u(p)=0.35 kPa, u(x1)=0.008, u(y1)=0.006. 6
1
Table 6
2
Experimental isobaric VLE data for the ternary system EA (1) + IPA (2) + DMF (3)
3
and the values of RMSD for NRTL, UNIQUAC and Wilson at 101.3 kPa. a T/K
Liquid phase x1
379.18 378.51 377.88 377.31 376.79 376.34 375.96 375.66 375.44
0.100 0.199 0.301 0.395 0.501 0.600 0.702 0.803 0.901
x2 0.900 0.801 0.699 0.605 0.499 0.400 0.298 0.197 0.099
Vapor phase y1 0.136 0.249 0.365 0.454 0.559 0.650 0.743 0.833 0.917
4
a
5
u(x1)=0.008, u(y1)=0.006.
y2
NRTL
UNIQUAC
Wilson
RMSD
RMSD
RMSD
(y1)
(T)
(y1)
Solvent to feed ratio=2.5:1 0.864 0.008 0.76 0.043 0.751 0.010 0.07 0.076 0.635 0.015 0.02 0.095 0.546 0.006 0.16 0.112 0.441 0.012 0.32 0.107 0.350 0.008 0.42 0.101 0.257 0.010 0.46 0.083 0.167 0.010 0.43 0.060 0.083 0.006 0.34 0.033
(T)
(y1)
0.39 0.28 0.18 0.81 0.28 0.17 0.73 0.53 0.41
0.042 0.074 0.093 0.109 0.104 0.099 0.082 0.059 0.033
α12
(T) 0.34 0.15 0.86 0.77 0.59 0.41 0.26 0.13 0.04
1.41 1.34 1.33 1.27 1.26 1.23 1.22 1.22 1.21
Standard uncertainties u of T, p, x and y are u(T)=0.35 K, u(p)=0.35 kPa,
6
7 8
Fig. 2. T-x-y diagram for the binary system EA (1) + DMSO (2) at 101.3 kPa; ■, T-x,
9
experimental values; ●, T-y, experimental values; ▲, T-x, reference values [48]; ▼,
10
T-y, reference values [48]; —, calculated by the NRTL model; —, predicted by the
11
UNIFAC model.
7
1 2
Fig. 3. T-x-y diagram for the binary system IPA (1) + DMF (2) at 101.3 kPa; ■, T-x,
3
experimental values; ●, T-y, experimental values; —, calculated by the NRTL model.
4
3.2 Effects of entrainers
5 6
To compare the suitability of the entrainers, the relative volatility (α12) of EA (1) to IPA (2) was used in this work and is defined as follows: y1
α12=y2
7
x1
(1)
x2
8
The diagram of the relative volatility for the system (EA + IPA), which are
9
presented in Fig. 4. For comparison of the relative volatility, the reference value [25]
10
of EA (1) + IPA (2) was also plotted in Fig. 4. As can be seen from Fig. 3, the α12
11
values with the added entrainers are all greater than 1. Furthermore, the values of α12
12
for the system EA (1) + IPA (2) with DMSO as entrainer are greater than those with
13
DMF, which is consistent with the calculation results of interaction energy in section
14
3.3.
8
1 2
5
Fig. 4. Relative volatility vs. x1 plot: ■, the reference value [23] of EA (1) + IPA (2); ●, experimental value for the system EA (1) + IPA (2) with the entrainer DMSO; ▲, experimental value for the system EA (1) + IPA (2) with the entrainer DMF; —, fitted line.
6
Also, the x-y plot for the system EA (1) + IPA (2) with two different entrainers
7
(DMF and DMSO) are presented in Fig. 5. As shown in Fig. 5, the azeotropic point is
8
eliminated with adding the entainers. With DMSO as the entrainer, there is a bigger
9
deviation from the x-y diagonal. Therefore, DMSO was selected as a suitable
3 4
10
11 12 13 14 15
entrainer to design the ED process for separating the binary mixture of EA and IPA.
Fig. 5. Influence on VLE for the system EA (1) + IPA (2) with different entrainers. ●, the reference value [9]; ■, the reference value [25]; —, calculated value by Aspen plus; ▲, experimental value (DMF as the extractive solvent); ■, experimental value (DMSO as the extractive solvent); —, fitted lines. 9
1
3.3 Interaction energy calculation
2
For exploring the interactions between molecules, the interaction energy, which is
3
the energy difference between the different entrainers (DMF and DMSO) and the
4
components (IPA and EA) was calculated using the DMol 3 module incorporated the
5
density functional theory. The detailed calculation was provided in our previous work
6
[49-50]. The optimized geometries and calculated interaction energies for the
7
entrainers and the components are shown in Fig. 6 and are summarized in Table 7,
8
respectively.
9
As shown in Table 7, the interaction energy values of the systems (IPA + DMF)
10
and (IPA + DMSO) are greater than that of the system (EA + IPA), which indicates
11
that the entrainers of DMF and DMSO can enhance the relative volatility of EA to
12
IPA. Compared with IPA + DMF, EA + DMF, EA + DMSO and IPA + EA, the
13
interaction energy for IPA + DMSO is the largest value. In the meantime, the
14
interaction energies between IPA and the entrainers are higher than those between
15
EA and the entrainers, which indicates that EA can be separated from the top of the
16
extractive distillation column.
17
18 19
EA+DMF
EA+DMSO
20
10
IPA+DMF
1
IPA+DMSO
2
EA + IPA
3
Fig. 6. Optimized geometries for the systems (EA + DMF), (EA + DMSO), (IPA +
4
DMF), (IPA + DMSO) and (IPA + EA).
5
Table 7
6
Interaction energies and bond lengths for the systems (EA + DMF), (EA + DMSO)
7
(IPA + DMF), (IPA + DMSO), and (IPA + EA). System
Bond lengths (nm)
Interaction energies (kJ·mol-1)
IPA+DMF EA+DMF IPA+DMSO EA+DMSO IPA+EA
0.1976 0.3141 0.1921 0.3404 0.1965
-21.5942 -1.5728 -22.9960 -3.3084 -18.5746
8
4. VLE calculation and regression
9
4.1 VLE calculation
10 11
The VLE thermodynamic relationship can be expressed by eq (2) [51]: Pyiφi =
φsi Psi exp and
VLi P-Psi RT
(2)
refer to liquid composition, vapor composition,
12
where xi, yi, T, P,
13
temperature, vapor pressure, fugacity coefficients at system pressure and saturation
14
pressure. The exponential part is Poynting factor, which is close to unity at 101.3 kPa. 11
1
Also, the fugacity coefficients are approximately close to unity at 101.3 kPa. So, eq
2
(2) can be simplified and expressed as:
3
γi=
Pyi
(3)
Psi xi
4
where Psi represents the pure compound vapor pressure and can be calculated by the
5
extended Antoine equation: ln(Psi /
6
)=C1i+
C2i
T/K+C3i
+C4i(T/K)+C5iln(T/K)+C6i( / )
C8i ≤T/K≤C9i
(4)
7
The parameter values of all the pure components are listed in Table 8, which were
8
taken from the Aspen databank [52]. The values of activity coefficient for the binary
9
mixtures were calculated and are listed in Tables 3-4.
10
Table 8
11
Parameters of the extended Antoine equation a Component EA DMF IPA DMSO
C 2i
C1i 59.92 75.85 103.81 49.37
C3i
-6227.60 -7955.50 -9040.00 -7620.60
C4i
0 0 0 0
C5i
0 0 0 0
12
a
13
4.2 Thermodynamic consistency test
-6.41 -8.80 -12.68 -4.63
C6i(×10-6) 1.79×10 4.24 5.54 0.44
-11
C7i
C8i /K
C9i /K
6 2 2 2
189.60 212.72 185.26 291.67
523.30 649.60 508.30 729.00
Taken from Aspen property databank. [52]
14
The van Ness test [31] and Herington test [32] were adopted to check the
15
thermodynamic consistency of the experimental VLE data. The van Ness test is
16
expressed as follows: !
17
∆y= ∑ & 100
18
∆P= ∑ & 100 '
(
)*+
(
−
#
%$,(-./
)*+
'
(5) (6)
19
where N represents the total number of experimental points; the superscripts “exp”
20
and “cal” are experimental and calculated values, respectively. In this test, the criteria
21
are the values of ∆y and ∆P less than 1.
22 23
The Herington method can be expressed as follows: (1,2)
D=1000(132)0
(7) 12
45.* ,45 6 0 45 6
J=1500
1
(8)
2
where A and B are the total area of ln(γ1/γ2 ) vs. x above the zero line and below the
3
zero line, respectively. Tmax and Tmin represent the maximum and minimum boiling
4
temperatures, respectively. If the value of |8 − 9| < 10, the measured VLE data
5
pass the thermodynamic consistency test.
6 7
The results of ∆y, ∆P and |8 − 9| are summarized in Table 9. As listed in Table 9,
all the values of ∆y and ∆P are less than unity. The values of |8 − 9| are less than
8
10. All the results indicate that the measured VLE data for the mixtures passed the
9
thermodynamic consistency test.
10 11
12
Table 9 Results of the thermodynamic consistency System
D
J
|8 − 9|
∆P
∆y
EAT+ DMSO
45.15
48.40
3.25
0.03
0.80
IPA+ DMF
39.29
29.97
9.32
0.06
0.35
4.3 VLE data regression
13
Based on the maximum likelihood principle [53-54], the determined binary VLE
14
data for the systems (EA + DMSO) and (IPA + DMF) were correlated by the NRTL,
15
UNIQUAC and Wilson models. The objective function is expressed as:
16
17 18 19 20 21
22
(
F=∑ FG
)*+
,(-./
HI
K
J +G
4
)*+
,4 -./
HM
K
J +G
)*+
K , -./
H*
J +G
N
)*+
,N -./
HO
K
J P
(9)
where T, P and Q refer to temperature, pressure and standard deviation at the equilibrium state. The values of standard deviation are as follows: Q( , 0.35 kPa; Q4 , 0.35 K; Q , 0.008; QN , 0.006.
For the Wilson model, the molar volumes of the pure components were calculated by the Rackett equation [55-56]. The Rackett equation is expressed as follows: R ∗,% =
U4- VW
_ ) ∗,XY [\([]M^
Z
(10)
(-
#
23
where Tr = T/Tci, T refers to temperature, Tci denotes critical temperature,
24
critical pressure, ` ∗,U1 is critical compressibility factor. All the parameters of the
25
Rackett equation were taken from Aspen plus and are listed in Table 10. 13
is
1
Table 10
2
Parameters of the Rackett equation for the components of EA, IPA, DMF and
3
DMSO. Component
Tci / K
EA IPA DMF DMSO
523.3 508.3 649.6 729.0
#
Zi*RA
/ kPa
3880.0 4765.0 4420.0 5650.0
0.26 0.25 0.21 0.21
4
The values of the binary interaction parameters of the NRTL, UNIQUAC and
5
Wilson models and the root-mean-square deviation (RMSD) are summarized in Table
6
11. From Table 11, the maximum RMSD values in liquid phase mole fraction (x1),
7
vapor phase mole fraction (y1), pressure (P) and temperature (T) are 0.0003, 0.0095,
8
0.54 K and 0.0027 kPa, respectively. It is concluded that three models can correlate
9
well the experimental VLE data. Also, for the ternary systems (IPA + EA + DMF)
10
and (IPA + EA + DMSO), the NRTL model with the regressed parameters was used
11
to predict the VLE data for the ternary mixtures. As shown in Tables 5-6, the values
12
of root-mean-square deviation of vapor mole fraction for the ternary mixtures are less
13
than 0.018, which indicates that the NRTL model with the regressed parameters can
14
be applied for the extractive distillation process design and optimization.
15 16
Table 11 Regressed parameters of the NRTL models and root-mean-square deviation (RMSD) of the systems.
17
Model
Parameters aij
e
18
aji
NRTL UNIQUAC f Wilson g
-2.3808 20.1761 13.7374
2.1266 -1.2389 6.3205
NRTL UNIQUAC Wilson
1.3430 0.4269 -13.0633
-6.4791 -7.1437 5.2320
a
RMSD(x1)=a∑
RMSD
bij/K
bji/K
EA+ DMSO 1002.95 -502.06 -1536.07 295.07 3009.14 -8042.17 IPA+ DMF -646.96 2652.42 -1218.49 2520.70 4081.92 -1714.87
_ )*+ , [-./ [
14
x1
a
y1
b
T/K c
P/ kPa d
0.0002 0.0003 0.0005
0.0095 0.0087 0.0034
0.54 0.52 0.37
0.0014 0.0026 0.0042
0.0003 0.0007 0.0006
0.0046 0.0081 0.0072
0.24 0.18 0.20
0.0027 0.0008 0.0008
)*+
RMSD(y1)=a∑
1
b
2
c
3
d
4
e
5
f
6
g
7
5. Conclusions
N[
RMSD(T)=a∑
RMSD(P)=a∑
NRTL,b c =
4
)*+
(
c
,4 -./
)*+
_
,N[-./
,(-./
_ _
+ e c / , the value of alpha ij was set at 0.3.
UNIQUAC, b c = f g Wilson, lnh c =
c
c
+ e c/
+ e c/
8
For separation of the binary azeotrope (EA + IPA) by extractive distillation, DMF
9
and DMSO were selected as the entrainers. The binary and ternary VLE data for (IPA
10
+ DMF), (EA + DMSO), (IPA + EA + DMF) and (IPA + EA + DMSO) were
11
determined. The experimental binary VLE data were fitted by the thermodynamic
12
model of NRTL. The thermodynamic consistency was validated by the van Ness test
13
and Herington test for the binary VLE data (IPA + DMF) and (EA + DMSO).
14
Moreover, the interaction energies for the different entrainers (DMF and DMSO) and
15
the components (IPA and EA) were calculated. The results indicate that the
16
interaction energies for (IPA + DMF) and (IPA + DMSO) are greater than that of (EA
17
+ IPA), which means DMSO is the suitable entrainer for separation of the azeotrope
18
of (EA +IPA) by extractive distillation. Furthermore, the ternary VLE data for (IPA +
19
EA + DMF) and (IPA + EA + DMSO) were predicted using the NRTL, UNIQUAC
20
and Wilson models with the regressed parameters and compared with those measured
21
in this work. The values of RMSD of vapor mole fraction are less than 0.018, 0.112
22
and 0.109 for the NRTL, UNIQUAC and Wilson models, which indicates that the
23
regressed parameters for the NRTL model can be applied for the extractive
24
distillation design and optimization.
15
1
Acknowledgement
2
The authors are grateful for the financial support of the National Natural Science
3
Foundation of China (21878178), Shandong Provincial Key Research &
4
Development Project (2018GGX107001), Project of Shandong Province Higher
5
Educational Science and Technology Program (J18KA072).
6 7
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Declaration of interests ☒ The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper. ☐The authors declare the following financial interests/personal relationships which may be considered as potential competing interests: