Article pubs.acs.org/jced
Solvent Effects on Vapor−Liquid Equilibria of the Binary System 1‑Hexene + n‑Hexane Beatriz Marrufo,† Ben Rigby,‡ Jordi Pla-Franco,‡ and Sonia Loras*,‡ †
Departamento de Ingeniería Química Básica, Universidad del Zulia, 4011, Maracaibo, Venezuela Departamento de Ingeniería Química, Escuela Técnica Superior de Ingeniería, Universitat de València, 46100 Burjassot, Valencia, Spain
‡
ABSTRACT: In order to study the separation of 1-hexene and n-hexane, two solvents, 2-pentanol and ethyl-butyrate, are tested as possible entrainers for an extractive distillation. In this way, isobaric vapor−liquid equilibrium (VLE) data at 100 kPa have been measured for the two ternary systems formed by the initial mixture and one of the mentioned solvents: 1-hexene + n-hexane + ethyl butyrate and 1-hexene + n-hexane + 2-pentanol. VLE data for the four constituent binary systems have also been measured. These systems are 1-hexene + ethyl butyrate, n-hexane + ethyl butyrate, 1-hexene + 2-pentanol, and finally n-hexane + 2-pentanol. All binary systems show moderate positive deviations from the ideal behavior and do not form an azeotrope. The well-known local composition models Wilson, UNIQUAC, and NRTL have been used for correlating VLE data. Prediction with the UNIFAC method has been also obtained.
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INTRODUCTION Olefins and paraffins have very similar boiling points and are only separated by a few degrees in temperature. When this occurs, conventional distillation is not a recommended technique due its expensive cost, as a consequence of the high reflux ratios and large number of stages required for carrying out the separation. In these cases, one alternative is extractive distillation, which uses a third component called an entrainer to alter the relative volatility of the components to be separated. However, only a few specific solvents allow an efficient and economical extractive distillation process; thus, entrainer selection is a crucial step during the design stage. The choice of entrainer can be carried out correctly with complete and accurate vapor−liquid equilibrium (VLE) data of the systems formed by the components to be separated and the possible solvent. The present work tackles a part of thermodynamic research on the separation of paraffin and olefin mixtures using different solvents. In this work, the initial mixture is formed by 1-hexene and n-hexane, and ethyl butyrate and 2-pentanol are the entrainer candidates due to certain alcohols and esters being recommended as entrainers for separation of hydrocarbons.1 An ideal entrainer will increase the relative volatility considerably between the 1-hexane and hexane and it will be very selective. Another important characteristic that makes an entrainer more suitable is that it is one of the so-called “ecofriendly” solvents. In the present paper, both proposed compounds can be considered green solvents. In this work, we measured isobaric VLE data for the two ternary systems 1-hexene (1) + n-hexane (2) + ethyl butyrate (3) and 1hexene (1) + n-hexane (2) + 2-pentanol (3) and four constituent binary systems 1-hexene (1) + ethyl butyrate (3), n-hexane (2) + ethyl butyrate (3), 1-hexene (1) + 2-pentanol (3), and n-hexane (2) + 2-pentanol (3) at 100 kPa. In a previous paper,2 we reported © 2012 American Chemical Society
VLE data for the binary system 1-hexene (1) + n-hexane (2). Isobaric VLE data of binary hexane + 2-pentanol system was obtained by Linek et al.3 at different pressures but none was recorded at 100 kPa. However, for the other binary systems and for the ternary systems, no VLE data have been previously published.
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EXPERIMENTAL SECTION Chemicals. The chemicals 1-hexene (99+ % mass), ethyl butyrate (100 w ≥ 99.5, puriss. p.a.) and 2-pentanol (100 w ≥ 99) were supplied by Sigma-Aldrich and the chemical n-hexane (99.5+ % mass) was supplied by Fluka. Purity grade of 2-pentanol was increased to 99.99 % after being treated in a batch multistage rectification process in a Fischer SPALTROHR HMS-500 column. The other reagents were used without further purification since impurities are smaller than the detection limit of the analytical method used. The specifications of the used chemicals are summarized in Table 1.
Table 1. Specifications of Chemical Samples chemical name n-hexane 1-hexene ethyl butyrate 2-pentanol a
source
initial mass fraction purity
purification method
Fluka Aldrich Aldrich Aldrich
0.9950 0.9900 0.9900 0.9900
none none none distillation
final mass fraction purity
analysis method
0.9999
GCa GCa GCa GCa
Gas−liquid chromatography.
Received: September 3, 2012 Accepted: October 25, 2012 Published: November 12, 2012 3721
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Table 2. Experimental Vapor Pressure (P0i ) of Ethyl Butyratea
Apparatus and Procedure. The VLE data and the vapor pressure of the pure compounds were determined using a dynamic-recirculating still (Pilodist VLE 100 D) equipped with a Cottrell circulation pump. This still is capable of handling pressures from 0.25 to 400 kPa, and temperatures up to 523 K. The equilibrium temperature was measured with a digital Hart Scientific thermometer, model 1502A, and a Hart Scientific Pt 100 probe, model AlB0888, calibrated at the Spanish Instituto Nacional de Técnica Aeroespacial. The uncertainty is estimated to be 0.02 K. A Pilodist M101 pressure control system was used to measure and control the pressure and the heating power. The pressure is kept constant by means of a vacuum pump with and an electrovalve. The manometer was calibrated using the vapor pressure of ultrapure water. The uncertainty is estimated to be 0.05 kPa. In each VLE experiment, first step was to fix and hold the pressure using a vacuum pump, and then turn on the heating and stirring systems of the liquid mixture. The experiment was running at constant pressure until equilibrium was reached. Equilibrium conditions were assumed when no changes in temperature and pressure were recorded. Then, samples of liquid and condensate vapor were taken for analysis. To verify equilibrium conditions, the vapor and liquid were continuously analyzed until a value less than 0.001 of the variation of the mole fraction of both liquid and vapor phase was reached. The sample extractions were carried out with special syringes which allowed the withdrawal of small volume samples. Analysis. Compositions of the liquid and condensed phases were determined using a Varian CP-3800 gas chromatograph (GC), after calibration with gravimetrically prepared standard solutions. A flame ionization detector was used together with a 30 m, 0.25 mm i.d. capillary column CP-Wax 52 CB. The GC response peaks were treated with Varian Star No. 1 for MS Windows. The column, injector, and detector temperatures were (353.15, 473.15, and 493.15) K, respectively, for all systems. Very good peak separation was achieved under these conditions, and calibration analyses were carried out to convert the peak area ratio to the mass composition of the sample. The average absolute deviation in the mole fraction was usually less than 0.001.
a
RESULTS AND DISCUSSION Pure Component Vapor Pressures. The pure component vapor pressures for ethyl butyrate and the pure component vapor pressures for 2-pentanol, P0i , were determined experimentally using the same equipment as that used to obtain the VLE data. The pertinent results appear in Table 2 for ethyl butyrate and in Table 3 for 2-pentanol. The measured vapor pressures were correlated using the Antoine equation: Bi T /K + Ci
P/kPa
T/K
P/kPa
22.50 25.00 27.50 30.03 32.51 35.02 37.52 40.04 42.51 45.03 47.51 50.01 52.51 55.03 57.52 60.00
379.48 380.68 381.82 382.96 384.07 385.13 386.19 387.22 388.21 389.17 390.12 391.06 391.95 392.83 393.71 394.13
65.03 67.53 70.00 72.52 75.01 77.51 79.99 82.52 85.02 87.54 89.98 92.53 94.92 97.45 100.01 101.31
u(T) = 0.02 K, and u(p) = 0.05 kPa.
Table 3. Experimental Vapor Pressure (P0i ) of 2-Pentanola
■
ln Pi0/kPa = Ai −
T/K 348.97 351.83 354.39 356.76 359.00 361.08 363.04 364.86 366.66 368.24 369.87 371.44 372.86 374.31 375.67 376.95
a
T/K
P/kPa
T/K
P/kPa
354.47 356.60 358.86 360.81 362.76 364.38 366.00 367.49 369.03 370.52 371.77 373.05 374.33 375.51 376.68 377.81 378.95
22.55 25.05 27.47 30.03 32.55 35.08 37.56 40.03 42.51 45.07 47.54 50.02 52.58 55.03 57.48 60.03 62.45
379.97 381.01 381.96 382.90 383.88 384.79 385.64 386.56 387.35 388.19 389.00 389.76 390.55 391.31 392.07 392.26
65.05 67.58 70.03 72.56 75.05 77.59 80.03 82.58 85.01 87.55 90.09 92.52 95.02 97.55 100.03 101.38
u(T) = 0.02 K, and u(p) = 0.05 kPa.
Table 4. Antoine Coefficients, eq 1
a
(1)
whose parameters Ai, Bi, and Ci are reported in Table 4 together with the Antoine parameters for n-hexane and 1-hexene obtained in a previous work2 and were fitted by a nonlinear optimization method to minimize the average relative deviation in pressure (ARDP). The vapor pressures of ethyl butyrate were correlated with an ARDP of 0.07 %. A value for the ARDP equal to 0.18 % was obtained for the case of 2-pentanol. For both compounds experimental data are in good agreement to the equation reported by Reid et al.4 since this equation gives a correlation of the experimental vapor pressures reported in this work with a ARDP = 0.39 % for the ethyl butyrate and a ARDP = 1.01 % for the 2-pentanol.
compound
Ai
Bi
Ci
1-hexene (1)a n-hexane (2)a ethyl butyrate (3) 2-pentanol (3)
15.1210 12.8239 14.2846 14.1880
3433.84 2178.19 3234.24 2684.71
−9.47 −76.39 −59.59 −111.86
Parameters obtained in ref 2.
Binary Systems. The temperature T and the liquid-phase xi, and vapor-phase yi mole fractions at 100.0 kPa for the systems 1-hexene (1) + ethyl butyrate (3), n-hexane (2) + ethyl butyrate (3), 1-hexene (1) + 2-pentanol (3), and n-hexane (2) + 2-pentanol (3) are reported in Tables 5−8 and plotted in Figures 1−4. The activity coefficients (γi) were calculated from the following equation5 assuming nonideality of both liquid and vapor phases: ln γi = ln
yP i xiPi0
+
(Bii − ViL)(P − Pi0) P + RT 2RT
∑ ∑ yyi k (2δji − δjk) (2)
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Table 5. Experimental Vapor−Liquid Equilibrium Data for the Binary System 1-Hexene (1) + Ethyl Butyrate (3) at 100.0 kPaa
a
T/K
x1
y1
γ1
γ3
389.91 384.16 378.87 374.16 370.93 366.17 362.24 359.78 356.05 353.79 351.38 349.22 347.45 345.55 343.80 342.17 340.51 339.04
0.027 0.069 0.112 0.160 0.198 0.256 0.318 0.351 0.426 0.468 0.525 0.573 0.630 0.679 0.735 0.804 0.851 0.905
0.133 0.300 0.433 0.535 0.603 0.686 0.749 0.779 0.825 0.850 0.876 0.896 0.916 0.931 0.943 0.960 0.970 0.984
1.224 1.225 1.227 1.185 1.167 1.155 1.123 1.128 1.087 1.084 1.064 1.060 1.036 1.032 1.016 0.992 0.995 0.992
0.991 0.988 0.983 0.987 0.979 0.977 0.973 0.981 1.002 1.006 1.017 1.029 1.026 1.045 1.119 1.132 1.193 1.059
Table 7. Experimental Vapor−Liquid Equilibrium Data for the Binary System 1-Hexene (1) + 2-Pentanol (3) at 100.0 kPaa
u(T) = 0.02 K, u(p) = 0.1 kPa, and u(x2) = u(y2) = 0.001. a
Table 6. Experimental Vapor−Liquid Equilibrium Data for the Binary System n-Hexane (2) + Ethyl Butyrate (3) at 100.0 kPaa
a
T/K
x2
y2
γ2
γ3
391.29 388.09 380.32 375.04 371.15 366.84 364.07 359.85 357.90 355.64 353.83 352.05 350.37 349.07 347.72 346.41 345.27 344.13 343.07 342.44
0.015 0.038 0.098 0.148 0.195 0.251 0.294 0.373 0.418 0.472 0.520 0.573 0.633 0.682 0.730 0.783 0.833 0.878 0.925 0.954
0.085 0.192 0.398 0.517 0.596 0.671 0.709 0.778 0.804 0.832 0.852 0.873 0.891 0.905 0.921 0.935 0.947 0.963 0.978 0.987
1.677 1.594 1.509 1.459 1.396 1.353 1.306 1.256 1.218 1.184 1.155 1.127 1.090 1.065 1.052 1.033 1.016 1.013 1.007 1.004
0.994 0.984 0.983 0.983 0.985 0.994 1.023 1.017 1.036 1.061 1.098 1.13 1.201 1.269 1.308 1.408 1.559 1.558 1.571 1.552
x1
y1
γ1
γ3
0.017 0.032 0.056 0.100 0.125 0.150 0.215 0.282 0.344 0.395 0.456 0.515 0.577 0.628 0.681 0.734 0.789 0.841 0.895 0.949 0.983
0.140 0.268 0.395 0.568 0.626 0.678 0.778 0.834 0.873 0.892 0.910 0.923 0.933 0.940 0.947 0.953 0.958 0.961 0.971 0.983 0.994
2.136 2.323 2.179 2.058 2.006 1.939 1.854 1.731 1.628 1.559 1.468 1.395 1.314 1.255 1.206 1.155 1.104 1.062 1.030 1.005 0.992
0.999 0.963 0.969 0.939 0.983 0.974 0.973 0.992 0.968 1.010 1.041 1.099 1.181 1.267 1.383 1.536 1.796 2.300 2.686 3.368 3.639
u(T) = 0.02 K, u(p) = 0.1 kPa, and u(x2) = u(y2) = 0.001.
Table 8. Experimental Vapor−Liquid Equilibrium Data for the Binary System n-Hexane (2) + 2-Pentanol (3) at 100.0 kPaa
u(T) = 0.02 K, u(p) = 0.1 kPa, and u(x1) = u(y1) = 0.001.
a
where T and P are the equilibrium temperature and pressure, ViL is the molar liquid volume of component i, Bii and Bjj are the second virial coefficients of the pure gases, Poi is the purecomponent vapor pressure, Bij the cross second virial coefficient and
δij = 2Bij − Bjj − Bii
T/K 388.05 384.92 380.13 373.28 369.17 366.41 359.50 354.51 351.19 348.56 346.34 344.39 342.90 341.85 340.71 339.87 339.14 338.40 337.71 337.00 336.63
T/K
x2
y2
γ2
γ3
389.01 384.49 377.11 369.96 365.17 359.96 355.46 352.14 350.43 348.74 347.29 346.18 345.15 344.42 343.74 343.13 342.55 342.10 341.72 341.46
0.013 0.033 0.080 0.127 0.165 0.232 0.294 0.372 0.414 0.470 0.531 0.588 0.646 0.699 0.753 0.806 0.862 0.911 0.952 0.992
0.106 0.248 0.468 0.605 0.684 0.776 0.824 0.863 0.875 0.887 0.901 0.910 0.920 0.926 0.930 0.937 0.948 0.959 0.971 0.994
2.535 2.563 2.337 2.241 2.188 2.010 1.893 1.712 1.635 1.530 1.432 1.348 1.278 1.214 1.154 1.106 1.064 1.032 1.011 1.001
1.002 1.005 0.978 1.008 1.024 0.984 1.025 1.044 1.105 1.196 1.269 1.386 1.508 1.701 2.028 2.396 2.863 3.583 4.794 6.040
u(T) = 0.02 K, u(p) = 0.1 kPa, and u(x1) = u(y1) = 0.001.
coefficient is adequate to describe the vapor phase of the pure components and their mixtures, and liquid volumes of the pure components are incompressible over the pressure range under consideration. The molar virial coefficients Bii and Bij were estimated by the method of Hayden and O’Connell6 using the molecular parameters suggested by Prausnitz et al.7 The critical properties of all components were taken from DIPPR.8 According to the results, the binary systems show positive deviations from ideal behavior and do not form any azeotrope.
(3)
The standard state for the calculation of activity coefficients is the pure component at the pressure and temperature of the solution. Equation 2 is valid at low and moderate pressures when the virial equation of state truncated after the second 3723
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Figure 1. Experimental VLE data for the system 1-hexene (1) + ethyl butyrate (3) at 100.0 kPa: ●, experimental data; solid line, smoothed data using the NRTL model with the parameters given in Table 10; dotted−dashed line, predicted by UNIFAC method.
Figure 3. Experimental VLE data for the system 1-hexene (1) + 2-pentanol (3) at 100.0 kPa: ●, experimental data; solid line, smoothed data using the NRTL model with the parameters given in Table 10, dotted−dashed line, predicted by UNIFAC method.
Figure 2. Experimental VLE data for the system n-hexane (2) + ethyl butyrate (3) at 100.0 kPa: ● , experimental data; solid line, smoothed data using the NRTL model with the parameters given in Table 10; dotted−dashed line, predicted by UNIFAC method.
Figure 4. Experimental VLE data for the system n-hexane (2) + 2-pentanol (3) at 100.0 kPa: ●, experimental data; solid line, smoothed data using the NRTL model with the parameters given in Table 10, dotted−dashed line, predicted by UNIFAC method.
Table 9. Consistency Test Statistics for the Binary Systems 1-Hexene (1) + Solvent (3) and n-Hexane (2) + Solvent (3)
The thermodynamic consistency of the VLE data, for each binary system, has been verified with the Fredenslund test.9 Pertinent consistency details and statistics are presented in Table 9, and it can be seen that the consistency criteria (AADy < 0.01) was achieved using a two parameter Legendre polynomial for the binary systems with ethyl butyrate and three parameters for the binary sytems with 2-pentanol. The VLE data for each binary system have been correlated using local composition models (Wilson, NRTL, and UNIQUAC) and predicted by the UNIFAC contribution method.9,10 For the Wilson model, molar liquid volumes of pure components have been estimated with the Rackett equation,11 and for UNIQUAC model, volume and surface parameters were taken from DECHEMA.12
system i + solvent (j)
A1a
A2a
1 + ethyl butyrate (3) 2 + ethyl butyrate (3) 1 + 2-pentanol (3) 2 + 2-pentanol (3)
0.1643 0.4726 1.0054 1.2294
0.0120 0.0845 0.3577 0.4009
A3a
100·AAD yib
AAD Pc/kPa
0.0610 0.0941
0.625 0.529 0.872 0.270
0.635 0.176 0.539 0.837
a
Legendre polynomial parameters. bAverage absolute deviation in vapor-phase composition. cAverage absolute deviation in pressure.
The parameters of these models have been determined minimizing the following objective function (OF) 3724
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Table 10. Parameters and Correlation Statistics for Different GE Models for the Ternary Systems 1-Hexene (1) + n-Hexane (2) + Solvent (3) model Wilsonc
NRTL
UNIQUACf
UNIFACg
Wilsonc
Aij
Aji
system i + j
J·mol−1
J·mol−1
1 + 2d 1+3 2+3 1 + 2 + 3e 1 + 2d 1+3 2+3 1 + 2 + 3e 1 + 2d 1+3 2+3 1 + 2 + 3e 1 + 2d 1+3 2+3 1+2+3
171.69 200.90 30.68
1+3 2+3 1 + 2 + 3e 1+3 2+3 1 + 2 + 3e 1+3 2+3 1 + 2 + 3e 1+3 2+3 1+2+3
NRTL
UNIQUACf
UNIFAC
−154.16 549.21 1992.99 −78.87 −188.21 625.83
−343.44 15.18 4933.42 5239.21 2457.76 2644.33
bubble point αij
Ethylbutyrate (3) −131.00 502.90 1744.60 195.66 148.65 −237.06
0.2 0.3 0.3
88.58 380.77 −201.10
2-Pentanol (3) 4984.29 5618.24 −682.15 −260.36 −1042.08 −1024.92
0.3 0.3
ARDTa/%
100·AADy1b
0.056 0.191 0.054 0.218 0.056 0.190 0.054 0.217 0.056 0.189 0.054 0.225 0.062 0.230 0.518 0.290
0.115 0.651
0.138 0.136 0.164 0.151 0.270 0.258 0.132 0.242 0.225 1.395 0.970 1.051
0.885
0.473 0.116 0.654 0.473 0.116 0.653 0.486 0.072 0.752 0.487
0.338 0.951 0.341 0.897 0.368 1.141 0.879
100·AADy2b
0.571 0.254
0.578 0.254
0.582 0.259
1.002 0.357
0.295 0.585 0.331 0.599 0.292 0.597 0.484 0.841
a Average relative deviation in temperature. bAverage absolute in vapor phase composition. cMolar liquid volumes of pure components have been estimated with the Rackett equation.11 dReference 2. eTernary estimation from binary parameters. fVolume and surface parameters from DECHEMA.12 g Calculations based on UNIFAC.9,10
N
OF =
⎛ T expt − T calc ⎞ i i expt calc ⎟ y y + | − | ⎟ i i Tiexpt ⎝ ⎠
∑ 100 × ⎜⎜ i=1
(4)
and are reported in Table 10, together with the obtained average deviations of the correlation. An inspection of the results given in that table shows that the three composition models are adequate for the description of the VLE of the binary systems, without any significant difference between them. However, it must be pointed out that the deviations obtained for the prediction with UNIFAC are larger. Ternary Systems. VLE data for the ternary system 1-hexene (1) + n-hexane (2) + ethyl butyrate (3) are reported in Table 11 and Figure 5 and VLE data for the ternary system 1-hexene (1) + n-hexane (2) + 2-pentanol (3) are reported in Table 12 and Figure 6. The activity coefficients (γi) were calculated from eq 2, and the molar virial coefficients were estimated as well as for the binary systems. The ternary data were found to be thermodynamically consistent by the Wisniak and Tamir13 modification of the McDermott−Ellis test.14 The test requires that D < Dmax for every experimental point, where the local deviation D is given by
Figure 5. Diagram of VLE for the ternary system 1-hexene (1) + n-hexane (2) + ethyl butyrate (3) at 100.0 kPa: ●, liquid-phase mole fractions; Δ, vapor-phase mole fractions.
N
D=
∑ (xia + xib)(ln γia − ln γib) i=1
and N is the number of components. The maximum deviation Dmax is given by
(5) 3725
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Table 11. Experimental Vapor−Liquid Equilibrium Data for the System 1-Hexene (1) + n-Hexane (2) + Ehtyl Butyrate (3) at 100.0 kPaa T/K
x1
x2
y1
y2
γ1
γ2
γ3
382.56 365.54 353.97 345.33 337.54 339.34 342.39 345.34 349.26 352.96 357.80 363.28 373.80 364.85 357.52 352.65 349.17 345.36 340.59 338.70 340.72 343.66 346.23 350.02 353.71 361.02 353.67 349.98 346.66 343.31 339.31 341.56 344.66 347.70 352.46 347.60 344.86 340.33 342.67 347.64 342.52 342.37
0.046 0.224 0.424 0.643 0.912 0.807 0.702 0.600 0.486 0.392 0.294 0.199 0.087 0.104 0.203 0.305 0.398 0.506 0.676 0.739 0.585 0.490 0.411 0.301 0.200 0.056 0.108 0.210 0.306 0.420 0.553 0.395 0.293 0.213 0.057 0.105 0.198 0.357 0.195 0.050 0.102 0.051
0.045 0.043 0.050 0.051 0.051 0.097 0.096 0.094 0.097 0.102 0.097 0.104 0.090 0.198 0.208 0.209 0.206 0.204 0.199 0.206 0.305 0.295 0.301 0.289 0.296 0.298 0.397 0.403 0.396 0.399 0.401 0.499 0.490 0.478 0.495 0.609 0.601 0.600 0.701 0.674 0.821 0.895
0.175 0.563 0.762 0.865 0.945 0.884 0.851 0.813 0.755 0.692 0.611 0.473 0.261 0.246 0.407 0.523 0.613 0.680 0.767 0.793 0.664 0.616 0.556 0.471 0.349 0.126 0.184 0.312 0.415 0.511 0.604 0.460 0.377 0.287 0.098 0.145 0.252 0.402 0.237 0.068 0.123 0.061
0.174 0.116 0.091 0.068 0.047 0.096 0.107 0.119 0.143 0.171 0.197 0.248 0.279 0.475 0.409 0.340 0.292 0.251 0.203 0.194 0.309 0.333 0.378 0.423 0.506 0.640 0.662 0.576 0.505 0.442 0.383 0.512 0.567 0.630 0.770 0.777 0.694 0.584 0.732 0.856 0.854 0.921
1.126 1.109 1.071 1.020 0.989 0.990 1.001 1.028 1.054 1.082 1.118 1.110 1.084 1.062 1.087 1.060 1.048 1.019 0.988 0.989 0.985 1.001 1.000 1.040 1.048 1.114 1.024 0.988 0.991 0.979 0.989 0.986 0.995 0.956 1.068 0.982 0.979 0.989 0.996 0.966 0.993 0.989
1.384 1.417 1.282 1.191 1.035 1.053 1.084 1.130 1.179 1.213 1.294 1.323 1.342 1.281 1.262 1.187 1.137 1.098 1.046 1.021 1.034 1.058 1.093 1.146 1.212 1.261 1.183 1.121 1.097 1.048 1.017 1.022 1.053 1.101 1.140 1.069 1.045 1.005 1.007 1.062 1.007 1.001
0.985 1.031 0.986 1.068 1.438 1.288 1.138 1.084 1.027 0.991 0.970 1.017 1.006 0.964 0.970 1.044 1.011 1.159 1.411 1.499 1.435 1.235 1.080 1.055 1.025 0.995 1.109 1.183 1.244 1.370 1.749 1.494 1.292 1.196 1.098 1.219 1.335 1.934 1.613 1.229 1.627 1.826
Figure 6. Diagram of VLE for the ternary system 1-hexene (1) + n-hexane (2) + 2-pentanol (3) at 100.0 kPa: ●, liquid-phase mole fractions; Δ, vapor-phase mole fractions.
in eq 6. For each experimental point reported in this work the value of D was always smaller than the value of Dmax. VLE data for both ternary systems have been estimated by using the Wilson, NRTL, and UNIQUAC models with the binary interaction parameters obtained from the regression of binary data and also are predicted by the UNIFAC method. Table 10 lists the average deviations between experimental and calculated temperatures and vapor phase mole fractions of the components. The three models represent the data successfully. Thus, these models can be used to calculate boiling points from liquid phase compositions at the system pressure. As an example, boiling isotherms calculated for both ternary systems with the aid of the NRTL model are presented in Figure 7 when ethyl butyrate is the solvent and Figure 8 for the case of 2-pentanol.
a
u(T) = 0.02 K, u(p) = 0.1 kPa, and u(x1) = u(x2) = u(y1) = u(y2) = 0.001.
⎛
N
Dmax =
∑ (xia + xib)⎜⎜ 1
+
⎝ xia
i=1
N
+
∑ (xia + xib) ΔP i=1 N
+
p
1 1 1⎞ + + ⎟⎟Δx yia xib yib ⎠ N
+ 2 ∑ |ln γb − ln γia|Δx i=1
∑ (xia + xib)Bj{(Ta + Cj)−2 + (Tb + Cj)−2 }ΔT i=1
(6)
Figure 7. Boiling isotherms (K) for the ternary system 1-hexene (1) + n-hexane (2) + ethyl butyrate (3) at 100.0 kPa calculated with NRTL model with the parameters given in Table 10.
The errors in the measurements Δx, ΔP, and ΔT were as previously indicated. The first and fourth are the dominant terms 3726
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Table 12. Experimental Vapor−Liquid Equilibrium Data for the System 1-Hexene (1) + n-Hexane (2) + 2-Pentanol (3) at 100.0 kPaa
Figure 8. Boiling isotherms (K) for the ternary system 1-hexene (1) + n-hexane (2) + 2-pentanol (3) at 100.0 kPa calculated with NRTL model with the parameters given in Table 10.
Solvent Effects. The solvent effects in the VLE of the mixture 1-hexene + n-hexane will be discussed taking into account two aspects: analysis of the residue curve map,15 and changes in relative volatility. Furthermore, both solvents will be compared with each other in order to recommend the most suitable entrainer for the separation of the mixture 1-hexene + n-hexane by extractive distillation. In Figure 9 and Figure 10, residue curves simulated by Aspen properties v7.3 using the NRTL model with the experimental
VT/K
x1
x2
y1
y2
γ1
γ2
γ3
377.02 353.23 344.42 340.56 337.17 338.06 339.45 340.99 342.96 345.32 348.90 352.35 353.65 358.48 351.95 346.86 343.90 341.54 340.04 338.36 338.03 339.63 341.43 343.70 345.72 349.55 350.51 345.42 342.75 341.03 338.80 339.96 341.63 343.07 346.69 343.53 341.97 339.86 340.95 343.40 341.54 341.40
0.044 0.252 0.454 0.646 0.900 0.794 0.696 0.590 0.487 0.392 0.305 0.225 0.209 0.111 0.127 0.247 0.364 0.493 0.580 0.711 0.727 0.540 0.434 0.318 0.235 0.138 0.098 0.176 0.295 0.382 0.546 0.399 0.293 0.210 0.051 0.103 0.203 0.344 0.208 0.057 0.104 0.056
0.038 0.052 0.064 0.059 0.054 0.106 0.102 0.100 0.096 0.099 0.100 0.095 0.092 0.127 0.240 0.226 0.207 0.205 0.210 0.200 0.211 0.316 0.297 0.289 0.291 0.281 0.288 0.383 0.387 0.411 0.405 0.498 0.491 0.491 0.488 0.613 0.600 0.599 0.696 0.698 0.798 0.895
0.267 0.691 0.804 0.868 0.930 0.866 0.842 0.812 0.774 0.730 0.671 0.591 0.569 0.368 0.303 0.474 0.601 0.677 0.719 0.775 0.775 0.627 0.577 0.496 0.417 0.289 0.228 0.301 0.423 0.483 0.593 0.460 0.376 0.299 0.093 0.146 0.265 0.385 0.246 0.079 0.128 0.066
0.250 0.160 0.118 0.079 0.051 0.104 0.114 0.131 0.149 0.183 0.220 0.258 0.256 0.433 0.569 0.420 0.322 0.264 0.234 0.193 0.199 0.330 0.363 0.424 0.493 0.584 0.657 0.615 0.509 0.466 0.384 0.503 0.566 0.634 0.809 0.787 0.683 0.585 0.714 0.854 0.834 0.906
2.035 1.668 1.379 1.172 0.998 1.025 1.090 1.185 1.291 1.413 1.508 1.637 1.638 1.753 1.503 1.394 1.305 1.163 1.098 1.015 1.003 1.041 1.130 1.240 1.331 1.410 1.525 1.294 1.172 1.087 0.998 1.023 1.084 1.153 1.331 1.133 1.092 0.997 1.020 1.112 1.043 1.003
2.654 2.211 1.691 1.374 1.073 1.085 1.185 1.327 1.484 1.652 1.778 1.998 1.977 2.135 1.763 1.590 1.448 1.284 1.161 1.057 1.044 1.101 1.222 1.373 1.497 1.650 1.763 1.430 1.265 1.147 1.025 1.055 1.146 1.231 1.425 1.207 1.120 1.023 1.040 1.155 1.041 1.013
0.892 0.970 1.114 1.501 4.118 2.853 1.926 1.502 1.366 1.125 1.016 1.047 1.112 0.937 0.973 1.229 1.267 1.553 1.920 3.366 3.996 2.617 1.783 1.451 1.226 1.175 0.963 1.249 1.599 2.011 4.300 3.098 2.125 1.650 1.310 1.698 2.055 4.565 3.418 1.982 3.089 4.588
a
u(T)= 0.02 K, u(p) = 0.1 kPa, and u(x1) = u(x2) = u(y1) = u(y2) = 0.001.
node. Thus, 1-hexene could be obtained as an overhead product and any of the two solvents will be the bottom product. This behavior is consistent with the boiling points of the components employed here. In this case the residue curve maps give little valuable information. A more useful tool to evaluate the influence of adding a solvent in a binary mixture is to calculate the change in the relative volatility of the initial mixture. In this occasion, the value of the relative volatility of 1-hexene to n-hexane is very close to unity (α12 = 1.177).2 In this sense, Table 13 provides values of mean relative volatilities in the presence of the solvent (αS12) for two different solvents compositions (for x3 = 0.7 and 0.8). NRTL model has been used in order to calculate the relative volatilities.
Figure 9. Residue curve map for the ternary system 1-hexene (1) + n-hexane (2) + ethyl butyrate (3) at 100.0 kPa simulated by Aspen split using the NRTL model with the parameters given in Table 10.
parameters reported in Table 10 are shown. As can be seen in these figures, in both ternary systems there is only a one distillation region with three singular points (nodes and saddles) which correspond with the three pure component vertices. 1-Hexene is an unstable node; n-hexane is a saddle, and ethyl butyrate in Figure 5 and 2-pentanol in Figure 6 are both a stable 3727
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effect of adding solvent can be more appreciated. Data from both the Table 13 and the Figure 11 imply that the two proposed solvents show a poor behavior as extractive distillation entrainers because the two values of mean relative volatility in the presence of the solvents are closer to unity than the original solvent-free value. Additionally, a reversal in the initial mixture volatility is produced when ethyl butyrate is employed as entrainer. This behavior implies the refusal of the candidate entrainers for extractive distillation to separate 1-hexene from n-hexane.
■
CONCLUSIONS In order to study the 1-hexene + n-hexane separation process, two solvents, ethyl butyrate and 2-pentanol have been proposed as possible entrainers in an extractive distillation process. Hence, consistent VLE data at 100.0 kPa have been determined for the binary systems formed by 1-hexene (1) or n-hexane (2) with each proposed solvent. Also the two ternary systems constituted by the initial binary mixture and the solvent have been measured. Wilson, NRTL, and UNIQUAC models correlated well the binary systems and yielded good estimation for the ternary system, without any remarkable difference. From the results, it is observed that both solvents decrease the relative volatility of 1-hexene to n-hexane, ethyl butyrate even inverts this relative volatility for high compositions of this solvent. So, these solvents are not good entrainers for extractive distillation to separate 1-hexene from n-hexane.
Figure 10. Residue curve map for the ternary system 1-hexene (1) + n-hexane (2) + ethyl butyrate (3) at 100.0 kPa simulated by Aspen split using the NRTL model with the parameters given in Table 10.
Table 13. Mean Relative Volatility in the Presence of the Solvent (αS12) Calculated Using the NRTL Model for the System 1-Hexene (1) + n-Hexane (2) + Solvent (3) at 100.0 kPa
■
α12S solvent
x3 = 0.0
x3 = 0.7
x3 = 0.8
ethyl butyrate 2-pentanol
1.177 1.177
0.995 1.045
0.986 1.038
AUTHOR INFORMATION
Corresponding Author
*Tel.: +34 963544317. Fax: +34 963544898. E-mail: sonia.loras@ uv.es. Funding
Financial support from the Ministerio de Ciencia e Innovación of Spain, through Project No. CTQ2010-18848, is gratefully acknowledged. B.M. and J.P.-F. have been funded by a grant from La Universidad del Zulia of Venezuela and from the Ministerio de Economiá y Competitividad of Spain (BES-201104636 6), respectively. Notes
The authors declare no competing financial interest.
■
REFERENCES
(1) Brix-Berg. http://www.brix-berg.com/ruleofthumbgraph.htm (accessed July 27, 2012). (2) Marrufo, B.; Aucejo, A.; Loras, S.; Sanchotello, M. Isobaric VaporLiquid Equilibrium for Binary Mixtures of 1-Hexene + n-Hexane and Cyclohexane + Cyclohexene at 30, 60 and 101.3 kPa. Fluid Phase Equilib. 2009, 279, 11−16. (3) Linek, J.; Wichterle, I. ELDATA. Int. Electron. J. Phys-Chem. Data 1995, 1 (4), 265−274. (4) Reid, R. C., Prausnitz, J. M., Sherwood, T. K. The Properties of Gases and Liquids; McGraw-Hill: New York, 1977. (5) Van Ness, H. C.; Abbott, M. M. Classical Thermodynamics of Nonelectrolyte Solutions; McGraw-Hill: New York, 1982. (6) Hayden, J.; O’Connell, J. A Generalized Method for Predicting Second Virial Coefficients. Ind. Eng. Chem. Process Des. Dev. 1975, 14, 209−216. (7) Prausnitz, J.; Anderson, T.; Grens, E.; Eckert, C.; Hsieh, R.; O’Connell, J. Computer Calculation for Multicomponent Vapor-Liquid and Liquid-Liquid Equilibria; Prentice Hall: Englewood Cliffs, NJ, 1980. (8) Daubert, T. E.; Danner, R. P. Physical and Thermodynamic Properties of Pure Chemicals. Data Compilation; Taylor & Francis: Bristol, PA, 1989.
Figure 11. VLE data plotted on a solvent-free basis for the system 1-hexene (1) + n-hexane (2) + solvent (3) at 100.0 kPa. Continuous line4 for x3 = 0.00: Dashed line for x3 = 0.70, calculated using the NRTL model with the parameters given in Table 10: ···, ethyl butyrate; --, 2-pentanol.
Moreover, liquid molar fractions of each ternary system versus the differences between molar fractions of the vapor phase and the liquid phase of above systems are plotted in Figure 11. These representations are made on a solvent-free basis, so the 3728
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(9) Fredenslund, A.; Gmehling, J.; Rasmussen, P. Vapor-Liquid Equilibria Using UNIFAC. A Group Contribution Method; Elsevier: Amsterdam, 1977. (10) Hansen, H. K.; Rasmussen, P.; Fredenslund, A.; Schiller, M.; Gmehling, J. Vapor-Liquid Equilibria by UNIFAC Group Contribution. 5. Revision and Extension. Ind. Eng. Chem. Res. 1991, 30, 2352−2355. (11) Rackett, H. G. Equation of State for Saturated Liquids. J. Chem. Eng. Data 1970, 15, 514−517. (12) Gmehling, J.; Onken, U. Vapor-Liquid Equilibrium Data Collection; DECHEMA: Frankfurt, 1977. (13) Wisniak, J.; Tamir, A. Vapor-Liquid Equilibria in the Ternary System Water-Formic Acid-Acetic Acid and Water-Acetic AcidPropionic Acid. J. Chem. Eng. Data 1977, 22, 253−260. (14) McDermott, C; Ellis, S. R. M. A Multicomponent Consistency Test. Chem. Eng. Sci. 1965, 20, 293−296. (15) Doherty, M. F.; Malone, M. F. Conceptual Design of Distillation Systems; McGraw-Hill: New York, 2001.
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