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Vapour–liquid–liquid and vapour–liquid equilibrium of the system water + ethanol + heptane at 101.3 kPa
Fluid Phase Equilibria 248 (2006) 206–210
Vapour–liquid–liquid and vapour–liquid equilibrium of the system water + ethanol + heptane at 101.3 kPa Vicente Gomis ∗ , Alicia Font, Maria Dolores Saquete Departamento de Ingenier´ıa Qu´ımica, Universidad de Alicante, Ap. 99, E-03080 Alicante, Spain Received 22 June 2006; received in revised form 4 August 2006; accepted 7 August 2006 Available online 18 August 2006
Abstract Consistent vapour–liquid (VLE) and vapour–liquid–liquid (VLLE) data for the ternary system water (1) + ethanol (2) + n-heptane (3) are reported at 101.3 kPa at temperatures in the range of 341–352 K. An all-glass, dynamic recirculating still equipped with an ultrasonic homogenizer has been used in the determination of the VLLE. The experimental data indicates that a ternary heterogeneous azeotrope is present at 341.83 K with a composition of 0.205, 0.432, and 0.363 mole fractions of water, ethanol, and heptane, respectively. The experimental results have been used to test the accuracy of the UNIFAC, UNIQUAC, and NRTL models in the calculation and prediction of the equilibrium data. None of these methods properly correlate and approximate the top zone of the non-isothermal solubility curve, which is important for the simulation of the azeotropic distillation. © 2006 Elsevier B.V. All rights reserved. Keywords: Water; Ethanol; n-Heptane; Vapour–liquid–liquid equilibrium; Distillation
1. Introduction Systems composed of hydrocarbons, water, and ethanol are important in the fuel industry where ethanol + gasoline blends are widely used. Alcohol increases octane levels and also promotes a more complete fuel burning that reduces harmful exhaust pipe emissions. However, small amounts of water in this blend can lead to a phase split of the mixture, causing the motor malfunction. For this reason, ethanol must be dehydrated prior to blending with gasoline. This process is currently carried out by different techniques, adsorption and azeotropic distillation being the most commonly used. Heterogeneous azeotropic distillation works by adding a third component, which causes liquid–liquid phase separation over a wide range of compositions in the ternary phase diagram. This liquid–liquid phase split provides a cheap and efficient method for moving across distillation boundaries caused by the presence of azeotropes in the mixture. Design of a column sequence suitable for the dehydration of ethanol by azeotropic distillation requires knowledge of
∗
Corresponding author. Tel.: +34 965 90 38 67; fax: +34 965 90 38 26. E-mail address: [email protected] (V. Gomis).
0378-3812/$ – see front matter © 2006 Elsevier B.V. All rights reserved. doi:10.1016/j.fluid.2006.08.012
vapour–liquid (VLE) and vapour–liquid–liquid (VLLE) equilibrium data of the water + ethanol + hydrocarbon systems. However, despite the extensive use of this technique and the fact that n-heptane can be a good entrainer to dehydrate ethanol, complete experimental study of the isobaric VLLE and VLE of the water + ethanol + n-heptane is not available in the literature. The advantage of this entrainer is that the remaining quantity of heptane in the ethanol will not be a problem for its subsequent use as a fuel since it is one of the common compounds in gasoline. Lack of experimental VLE and VLLE data is usually substituted by their prediction using activity coefficient group contribution models like UNIFAC or by the UNIQUAC or NRTL models with parameters based on the correlation of VLE and LLE data. In order to correctly simulate azeotropic distillation columns, a set of fitted parameters that predict the equilibrium well is very important, since small deviations in the azeotropic compositions or in the non-isothermal binodal curve can lead to important deviations between the simulations and reality. The objective of the present paper is to report experimentally determined isobaric VLE and the VLLE of the system water + ethanol + n-heptane at normal atmospheric pressure. The results allow analyzing the reliability of the use of models such as UNIFAC, UNIQUAC and NRTL in the calculation and prediction of the equilibrium data.
V. Gomis et al. / Fluid Phase Equilibria 248 (2006) 206–210
2. Experimental 2.1. Chemicals All chemicals used were “for analysis” grade and obtained from Merck. No further purification was needed, since ethanol and n-heptane have nominal purities above 99.8 mass%, and 1propanol above 99.5 mass%. The Karl–Fischer titration method was used to determine the water content in mass% for n-heptane, ethanol and 1-propanol that was around 0.005, 0.08 and 0.07, respectively. The water used was purified using a Milli Q-Plus system.
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controlled by a potentiometer so that the vapour becomes unsaturated and condensation is avoided. (b) For the sampling of the liquid phase in the heterogeneous region, a small amount of the liquid coming from the separation chamber of the instrument was diverted to a tube using a solenoid valve. In this tube the dispersed liquid phases enter and separate into two layers at their bubble point. A sample of each layer was taken and placed in a vial with a small amount of 1-propanol as an internal standard. (c) In the homogeneous region, liquid samples were withdrawn from the liquid coming from the separator chamber with a syringe and put into a vial with a small amount of internal standard.
2.2. Apparatus and procedures An all-glass dynamic recirculating still with an ultrasonic homogenizer (Braun Labsonic P) coupled to the boiling flask was used to determine the VLLE data. This is a commercial unit (Labodest model 602) built in Germany by Fischer Labor und Verfahrenstechnik and modified by Gomis et al. [1]. The use of ultrasonic sound on the boiling flask causes the emulsification of the two liquid phases throughout the still and thus prevents the oscillations in temperature and flow rate of systems with two liquid phases. Visual, observation indicates that during operation the emulsion is maintained throughout the apparatus. Evaporation is carried on with an electrical immersion heater located concentrically in a flow heater. Before entering the separation chamber the vapour liquid mixture passes an extended contact line (Cottrell pump), which guarantees an intense phase exchange. The phases merge together on the thermometer, which measures its boiling point. The separated gas and liquid phases are condensed and returned to a mixing chamber, where they are stirred by a magnetic stirrer and returned again to the immersion heater, where the ultrasound homogenizer disperses the liquid phases. The construction of the separation chamber prevents an entrainment of liquid drops and partial condensation of the vapour phase. Regular circulation of both liquid and vapour phases and the simultaneous mixing of the back flowing circulation streams in the mixer chamber, allows a quick establishment of the equilibrium. For VLE determinations, the apparatus was used without further modification since it allows good mixing and separation of the vapour and liquid phases once they reach equilibrium. A Pt-100 sensor was employed to measure the equilibrium temperatures. The probe was connected to a Cropico thermometer (model 3002) with an uncertainty of 0.006 K according to the calibration certificate (scale ITS 90 [2]). A Fischer M101 control system was used to measure and control the pressure and the heating power. The pressure in the still was 101.3 kPa, measured and controlled with an accuracy of 0.1 kPa. In order to guarantee the correct operation of the equipment, the boiling point of water was measured and checked with that in the bibliography. Sampling was carried out using three different methods: (a) Gaseous samples were injected into the GC through an UW Type, 6-port valve from Valco Instruments Co. The connecting tube walls were superheated with a resistance tape
All analytical work was carried out by gas chromatography on a Shimadzu GC-14B coupled with a personal computer by using the Shimadzu CLASS-VP Chromatography Data System. Separation of the components was obtained on a 2 m × 3 mm column packed with Porapak Q 80/100. The oven temperature was 453 K and the helium flow rate was 40 mL/min. Detection was carried out by different techniques depending on the composition of the samples: TCD –Thermal Conductivity detector – for organic and aqueous samples (analyzing water, ethanol and heptane) and FID – Flame Ionization Detector – for aqueous samples (analyzing ethanol and heptane). The temperature of the detector was 473 K and the current for the TCD was 100 mA. Water in the organic phase was also determined using the Karl–Fischer titration method. An internal standard was used to obtain quantitative results in the analysis of the liquid phases. For this reason, 1-propanol, which is completely miscible in water, ethanol and heptane, was also added to the sample vials. Furthermore, the addition of the standard prevents phase split when changing the temperature after the separation of the phases. The accuracy of the mole fraction measurements was estimated at ±0.002 for all the compounds except for the water in the liquid organic phase and heptane in the aqueous phase, where the accuracy was approximately ±0.005. 3. Results and discussion The experimental VLLE data of the studied ternary system are reported in Table 1 and the corresponding activity coefficients are reported in Table 2. VLE data and the activity coefficients for the homogeneous region are shown in Table 3. In these tables, the bubble point (Tb (K)), the composition (mol fraction) of the liquid phases (xi ) and the vapour phase (yi ), as well as the activity coefficients (γ i ) are presented. Table 1 also reports the data for the binary heterogeneous azeotrope water–n-heptane (BIN). Fig. 1 represents the graphical ternary diagram that includes the vapour line and the non-isothermal binodal curve. This figure represents the projection on the ternary composition diagram of the intersection of the single liquid–liquid envelope with the VLE surface. It looks like an isothermal liquid–liquid envelope but it is not. This figure also shows the presence of a ternary azeotrope in the heterogeneous region. The points
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V. Gomis et al. / Fluid Phase Equilibria 248 (2006) 206–210
Fig. 1. Equilibrium diagram (mol%) for the ternary system water + ethanol + nheptane at 101.3 kPa.
of the vapour line corresponding to the liquid–liquid equilibrium lines 1–10 are above the tie line. However, the points corresponding to lines 11, 12 and 13 are under the tie line. Therefore, there is a point between lines 10 and 11 where the composition of the vapour coincides with that of a liquid heterogeneous mixture. The composition and temperature of the ternary azeotrope was determined by numerical interpolation as x1 = 0.205, x2 = 0.432, and x3 = 0.363, and 341.83 K, respectively. The composition of the two liquid phases in equilibrium is x1 = 0.030, x2 = 0.195, and x3 = 0.775 in the organic-rich phase, and x1 = 0.341, x2 = 0.614, and x3 = 0.045 in the aqueous-rich phase. Although the azeotropic temperature determined in this work is similar to that reported by Fritzweiler et al. in 1933 [3], the compositions differ significantly (azeotropic composition: x1 = 0.3625, x2 = 0.3035, x3 = 0.3340, T = 341.15 K). The ternary VLE and VLLE experimental data were tested by the point-to-point L–W Wisniak [4] method and found to be thermodynamically consistent. Calculations of the test and the activity coefficients shown in Tables 2 and 3 were done using the program PRO-VLE 2.0 [5]. All the values of L/W are between
Table 1 Vapour–liquid–liquid equilibrium data (mole fraction) for the ternary system water (1) + ethanol (2) + n-heptane (3) at 101.3 kPa Organic phase
BIN 1 2 3 4 5 6 7 8 9 10 11 12 13
Aqueous phase
Vapour phase
Tb (K)
x1
x2
x3
x1
x2
x3
y1
y2
y3
0.001 0.002 0.002 0.003 0.003 0.006 0.007 0.008 0.013 0.018 0.025 0.032 0.036 0.049
0.000 0.019 0.026 0.038 0.048 0.060 0.075 0.084 0.100 0.123 0.164 0.209 0.248 0.294
0.999 0.979 0.971 0.959 0.949 0.934 0.919 0.908 0.887 0.859 0.812 0.759 0.716 0.657
0.9998 0.914 0.845 0.786 0.718 0.672 0.612 0.543 0.487 0.429 0.402 0.320 0.282 0.235
0.000 0.086 0.154 0.213 0.280 0.325 0.383 0.448 0.497 0.547 0.569 0.629 0.646 0.646
0.0002 0.0003 0.0006 0.0010 0.002 0.003 0.006 0.009 0.015 0.025 0.029 0.050 0.072 0.119
0.439 0.313 0.284 0.271 0.261 0.253 0.246 0.238 0.230 0.222 0.215 0.201 0.195 0.190
0 0.257 0.314 0.335 0.357 0.368 0.381 0.394 0.406 0.415 0.423 0.434 0.442 0.449
0.561 0.429 0.403 0.394 0.382 0.379 0.373 0.368 0.365 0.362 0.362 0.364 0.363 0.362
352.39 345.47 343.98 343.34 342.97 342.77 342.60 342.50 342.34 342.25 342.22 341.87 341.89 341.91
Table 2 Activity coefficients of the VLLE data for the ternary system water (1) + ethanol (2) + n-heptane (3) at 101.3 kPa Tb (K)
BIN 1 2 3 4 5 6 7 8 9 10 11 12 13
352.39 345.47 343.98 343.34 342.97 342.77 342.60 342.50 342.34 342.25 342.22 341.87 341.89 341.91
Organic phase
Aqueous phase
γ1
γ2
γ3
γ1
γ2
γ3
970.6 455.16 440.04 287.74 281.6 137.69 115.61 98.31 58.88 41.22 28.78 21.35 18.4 13.16
– 17.16 16.25 12.17 10.42 8.67 7.23 6.7 5.84 4.87 3.72 3.04 2.61 2.23
1.023 0.98 0.97 0.98 0.97 0.99 1 1 1.01 1.05 1.11 1.21 1.27 1.38
0.971 1 1.04 1.1 1.18 1.23 1.32 1.45 1.57 1.73 1.79 2.14 2.35 2.75
– 3.79 2.74 2.17 1.79 1.6 1.42 1.26 1.17 1.09 1.07 1.01 1 1.02
5110.8 956.18 939.8 941.22 462.05 307.69 182.75 100.51 56.23 37.5 30.98 17.97 12.65 7.61
V. Gomis et al. / Fluid Phase Equilibria 248 (2006) 206–210
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Table 3 Vapour–liquid equilibrium data (mole fraction) and activity coefficients (γ i ) for the ternary system water (1) + ethanol (2) + n-heptane (3) at 101.3 kPa Tb (K)
343.28 344.70 345.18 347.22 348.64 350.90 348.93 345.82 345.26 347.81 348.61 346.57 350.10 347.56 345.89 345.12 344.38 343.42 343.09 342.89 342.82 342.80 342.77 342.74 342.69 342.68 342.80 342.84 342.86 342.51
Liquid phase
Vapour phase
Activity coefficients
x1
x2
x3
y1
y2
y3
γ1
γ2
γ3
0.174 0.144 0.144 0.138 0.128 0.074 0.227 0.232 0.321 0.315 0.411 0.465 0.071 0.071 0.069 0.066 0.089 0.081 0.074 0.069 0.064 0.061 0.059 0.052 0.048 0.042 0.034 0.034 0.024 0.027
0.716 0.786 0.794 0.831 0.853 0.922 0.762 0.732 0.653 0.669 0.580 0.527 0.919 0.889 0.865 0.845 0.809 0.730 0.679 0.636 0.596 0.567 0.522 0.469 0.431 0.388 0.345 0.299 0.251 0.224
0.109 0.070 0.062 0.031 0.019 0.004 0.012 0.036 0.026 0.016 0.009 0.008 0.009 0.040 0.066 0.090 0.102 0.189 0.248 0.294 0.339 0.372 0.419 0.479 0.521 0.570 0.621 0.667 0.725 0.749
0.145 0.121 0.119 0.116 0.112 0.071 0.177 0.167 0.193 0.208 0.237 0.236 0.080 0.076 0.074 0.074 0.096 0.099 0.104 0.107 0.110 0.113 0.117 0.122 0.125 0.129 0.126 0.132 0.142 0.169
0.510 0.575 0.603 0.685 0.755 0.906 0.707 0.592 0.527 0.616 0.603 0.529 0.870 0.767 0.700 0.671 0.621 0.576 0.556 0.544 0.538 0.533 0.527 0.519 0.513 0.507 0.507 0.497 0.485 0.461
0.345 0.303 0.278 0.199 0.132 0.024 0.116 0.242 0.281 0.176 0.160 0.235 0.050 0.157 0.225 0.255 0.283 0.325 0.340 0.349 0.352 0.354 0.357 0.359 0.362 0.364 0.368 0.371 0.373 0.370
2.67 2.54 2.45 2.29 2.25 2.25 1.97 2.07 1.77 1.75 1.48 1.41 2.73 2.88 3.09 3.33 3.31 3.90 4.55 5.06 5.63 6.07 6.51 7.71 8.57 10.11 12.14 12.69 19.32 20.72
0.98 0.95 0.97 0.97 0.99 1.00 1.02 1.01 1.03 1.06 1.16 1.22 1.00 1.00 1.01 1.02 1.01 1.08 1.14 1.20 1.27 1.32 1.42 1.56 1.69 1.85 2.07 2.34 2.72 2.94
7.51 9.91 10.07 13.48 14.04 10.73 20.96 14.72 24.13 22.67 35.70 63.04 9.58 8.16 7.51 6.45 6.40 4.10 3.32 2.87 2.52 2.31 2.07 1.83 1.70 1.56 1.44 1.35 1.25 1.21
0.97 and 0.98. The test did not reveal any significant inconsistency in the data. Vapour pressures of the three components were calculated using the Antoine equation, with parameters Ai , Bi , and Ci taken from literature [6,7] and shown in Table 4. VLE data correlation. The 13 ternary experimental VLLE data obtained (treated as 26 VLE data), the 30 experimental VLE data of the homogeneous region, together with binary data from the bibliography [8,9] have been correlated using the UNIQUAC and NRTL models. Only some points of each binary have been used to do the correlation. All the calculations have been made using CHEMCAD V [10]. Fig. 2 shows the graphical representation of three different experimental tie lines compared with those obtained using UNIQUAC and NRTL. Results of the tie lines obtained by UNIFAC (original) are also shown. The slope of the tie lines calculated using NRTL and UNIFAC is similar to the experimental one in the bottom part of the heterogeneous
region. In the top part, the experimental tie line coincides better with those obtained with UNIQUAC. In all cases, it can be observed that the experimental heterogeneous region is smaller than the calculated one that even contains some homogeneous experimental points. It is important to notice that the calculation of the equilibrium compositions of this region close to the
Table 4 Antoine equation parametersa of the pure substances Compound
A
B
C
Temperature range (K)
Water Ethanol n-Heptane
7.19621 7.2371 6.0273
1730.63 1592.864 1268.115
−39.724 −46.966 −56.25
+274/+373 +293/+366 –
a
Antoine equation: log(P) = A − B/[T + C], with: P (Pa) and T (K).
Fig. 2. Experimental and calculated LLE at 101.3 kPa. Parameters of the models obtained from correlation of VLE data.
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simulate the behaviour of the decanter since the prediction of the experimental data is better. However, none of these methods properly correlate and approximate the top zone of the nonisothermal solubility curve, which is important for the simulation of the azeotropic distillation. Acknowledgements The authors thank the European Community (FEDER Funds) and the DGICYT of Spain for the financial support of project PPQ2003-03385. References
Fig. 3. Experimental and calculated LLE at 101.3 kPa. Parameters of the models obtained from correlation of LLE data.
azeotrope is important for the simulation of ethanol dehydration. As seen in Fig. 2, the use of these models may lead to inaccurate predictions of the equilibrium and important differences in the simulation results. LLE data correlation. In order to improve these calculations, model parameters have been determined regressing exclusively LLE data. The non-isothermal VLLE experimental data treated as 13 LLE experimental tie lines have been correlated using UNIQUAC and NRTL models. Fig. 3 shows the representation of three experimental tie lines and those obtained by means of UNIFAC-LLE (original), UNIQUAC and NRTL and the fact that there is not too much improvement correlating only the tie lines. The top part of the heterogeneous region is not correlated with sufficient precision although these parameters could be used to
[1] V. Gomis, F. Ruiz, J.C. Asensi, Fluid Phase Equilib. 172 (2000) 245–259. [2] B.W. Mangum, G.T. Furukawa, US Department of Commerce, National Institute of Standards and Technology, Springfield, 1990. [3] J. Gmehling, J. Menken, J. Krafczyk, K. Fischer, Azeotropic Data, VCH, Weinheim, 1994. [4] J. Wisniak, Ind. Eng. Chem. Res. 32 (1993) 1531–1533. [5] M. Elly, M. Landa, J. Wisniak, PRO-VLE 2.0. Vapor–Liquid Equilibria Computer Program, Ben-Gurion University of the Negev, Beer-Sheva, 2003. [6] C.L. Yaws, Chemical Properties Handbook, McGraw-Hill, New York, 1999. [7] R.M. Felder, R.W. Rousseau, Principios Elementales de los Procesos Qu´ımicos, 2nd ed., Addison-Wesley Iberoamericana, Wilmington, 1986. [8] J.D. Raal, R.K. Code, D.A. Best, J. Chem. Eng. Data. 17 (1972) 211–221, VLE data extracted from: J. Gmehling, U. Onken, Chemistry Data Series, DECHEMA, Dortmund, 1977. [9] V.N. Stabnikov, B.Z. Matyushev, T.B. Protsyuk, N.M. Yushchenko, Pishch. Prom. (Kiev) 15 (1972) 49, VLE data extracted from: J. Gmehling, U. Onken, Chemistry Data Series, DECHEMA, Dortmund, 1977. [10] CHEMCAD V, Process Flow Sheet Simulator, Chemstations Inc., Houston, 2002.
Vapour–liquid–liquid and vapour–liquid equilibrium of the system water + ethanol + heptane at 101.3 kPa Vicente Gomis ∗ , Alicia Font, Maria Dolores Saquete Departamento de Ingenier´ıa Qu´ımica, Universidad de Alicante, Ap. 99, E-03080 Alicante, Spain Received 22 June 2006; received in revised form 4 August 2006; accepted 7 August 2006 Available online 18 August 2006
Abstract Consistent vapour–liquid (VLE) and vapour–liquid–liquid (VLLE) data for the ternary system water (1) + ethanol (2) + n-heptane (3) are reported at 101.3 kPa at temperatures in the range of 341–352 K. An all-glass, dynamic recirculating still equipped with an ultrasonic homogenizer has been used in the determination of the VLLE. The experimental data indicates that a ternary heterogeneous azeotrope is present at 341.83 K with a composition of 0.205, 0.432, and 0.363 mole fractions of water, ethanol, and heptane, respectively. The experimental results have been used to test the accuracy of the UNIFAC, UNIQUAC, and NRTL models in the calculation and prediction of the equilibrium data. None of these methods properly correlate and approximate the top zone of the non-isothermal solubility curve, which is important for the simulation of the azeotropic distillation. © 2006 Elsevier B.V. All rights reserved. Keywords: Water; Ethanol; n-Heptane; Vapour–liquid–liquid equilibrium; Distillation
1. Introduction Systems composed of hydrocarbons, water, and ethanol are important in the fuel industry where ethanol + gasoline blends are widely used. Alcohol increases octane levels and also promotes a more complete fuel burning that reduces harmful exhaust pipe emissions. However, small amounts of water in this blend can lead to a phase split of the mixture, causing the motor malfunction. For this reason, ethanol must be dehydrated prior to blending with gasoline. This process is currently carried out by different techniques, adsorption and azeotropic distillation being the most commonly used. Heterogeneous azeotropic distillation works by adding a third component, which causes liquid–liquid phase separation over a wide range of compositions in the ternary phase diagram. This liquid–liquid phase split provides a cheap and efficient method for moving across distillation boundaries caused by the presence of azeotropes in the mixture. Design of a column sequence suitable for the dehydration of ethanol by azeotropic distillation requires knowledge of
∗
Corresponding author. Tel.: +34 965 90 38 67; fax: +34 965 90 38 26. E-mail address: [email protected] (V. Gomis).
0378-3812/$ – see front matter © 2006 Elsevier B.V. All rights reserved. doi:10.1016/j.fluid.2006.08.012
vapour–liquid (VLE) and vapour–liquid–liquid (VLLE) equilibrium data of the water + ethanol + hydrocarbon systems. However, despite the extensive use of this technique and the fact that n-heptane can be a good entrainer to dehydrate ethanol, complete experimental study of the isobaric VLLE and VLE of the water + ethanol + n-heptane is not available in the literature. The advantage of this entrainer is that the remaining quantity of heptane in the ethanol will not be a problem for its subsequent use as a fuel since it is one of the common compounds in gasoline. Lack of experimental VLE and VLLE data is usually substituted by their prediction using activity coefficient group contribution models like UNIFAC or by the UNIQUAC or NRTL models with parameters based on the correlation of VLE and LLE data. In order to correctly simulate azeotropic distillation columns, a set of fitted parameters that predict the equilibrium well is very important, since small deviations in the azeotropic compositions or in the non-isothermal binodal curve can lead to important deviations between the simulations and reality. The objective of the present paper is to report experimentally determined isobaric VLE and the VLLE of the system water + ethanol + n-heptane at normal atmospheric pressure. The results allow analyzing the reliability of the use of models such as UNIFAC, UNIQUAC and NRTL in the calculation and prediction of the equilibrium data.
V. Gomis et al. / Fluid Phase Equilibria 248 (2006) 206–210
2. Experimental 2.1. Chemicals All chemicals used were “for analysis” grade and obtained from Merck. No further purification was needed, since ethanol and n-heptane have nominal purities above 99.8 mass%, and 1propanol above 99.5 mass%. The Karl–Fischer titration method was used to determine the water content in mass% for n-heptane, ethanol and 1-propanol that was around 0.005, 0.08 and 0.07, respectively. The water used was purified using a Milli Q-Plus system.
207
controlled by a potentiometer so that the vapour becomes unsaturated and condensation is avoided. (b) For the sampling of the liquid phase in the heterogeneous region, a small amount of the liquid coming from the separation chamber of the instrument was diverted to a tube using a solenoid valve. In this tube the dispersed liquid phases enter and separate into two layers at their bubble point. A sample of each layer was taken and placed in a vial with a small amount of 1-propanol as an internal standard. (c) In the homogeneous region, liquid samples were withdrawn from the liquid coming from the separator chamber with a syringe and put into a vial with a small amount of internal standard.
2.2. Apparatus and procedures An all-glass dynamic recirculating still with an ultrasonic homogenizer (Braun Labsonic P) coupled to the boiling flask was used to determine the VLLE data. This is a commercial unit (Labodest model 602) built in Germany by Fischer Labor und Verfahrenstechnik and modified by Gomis et al. [1]. The use of ultrasonic sound on the boiling flask causes the emulsification of the two liquid phases throughout the still and thus prevents the oscillations in temperature and flow rate of systems with two liquid phases. Visual, observation indicates that during operation the emulsion is maintained throughout the apparatus. Evaporation is carried on with an electrical immersion heater located concentrically in a flow heater. Before entering the separation chamber the vapour liquid mixture passes an extended contact line (Cottrell pump), which guarantees an intense phase exchange. The phases merge together on the thermometer, which measures its boiling point. The separated gas and liquid phases are condensed and returned to a mixing chamber, where they are stirred by a magnetic stirrer and returned again to the immersion heater, where the ultrasound homogenizer disperses the liquid phases. The construction of the separation chamber prevents an entrainment of liquid drops and partial condensation of the vapour phase. Regular circulation of both liquid and vapour phases and the simultaneous mixing of the back flowing circulation streams in the mixer chamber, allows a quick establishment of the equilibrium. For VLE determinations, the apparatus was used without further modification since it allows good mixing and separation of the vapour and liquid phases once they reach equilibrium. A Pt-100 sensor was employed to measure the equilibrium temperatures. The probe was connected to a Cropico thermometer (model 3002) with an uncertainty of 0.006 K according to the calibration certificate (scale ITS 90 [2]). A Fischer M101 control system was used to measure and control the pressure and the heating power. The pressure in the still was 101.3 kPa, measured and controlled with an accuracy of 0.1 kPa. In order to guarantee the correct operation of the equipment, the boiling point of water was measured and checked with that in the bibliography. Sampling was carried out using three different methods: (a) Gaseous samples were injected into the GC through an UW Type, 6-port valve from Valco Instruments Co. The connecting tube walls were superheated with a resistance tape
All analytical work was carried out by gas chromatography on a Shimadzu GC-14B coupled with a personal computer by using the Shimadzu CLASS-VP Chromatography Data System. Separation of the components was obtained on a 2 m × 3 mm column packed with Porapak Q 80/100. The oven temperature was 453 K and the helium flow rate was 40 mL/min. Detection was carried out by different techniques depending on the composition of the samples: TCD –Thermal Conductivity detector – for organic and aqueous samples (analyzing water, ethanol and heptane) and FID – Flame Ionization Detector – for aqueous samples (analyzing ethanol and heptane). The temperature of the detector was 473 K and the current for the TCD was 100 mA. Water in the organic phase was also determined using the Karl–Fischer titration method. An internal standard was used to obtain quantitative results in the analysis of the liquid phases. For this reason, 1-propanol, which is completely miscible in water, ethanol and heptane, was also added to the sample vials. Furthermore, the addition of the standard prevents phase split when changing the temperature after the separation of the phases. The accuracy of the mole fraction measurements was estimated at ±0.002 for all the compounds except for the water in the liquid organic phase and heptane in the aqueous phase, where the accuracy was approximately ±0.005. 3. Results and discussion The experimental VLLE data of the studied ternary system are reported in Table 1 and the corresponding activity coefficients are reported in Table 2. VLE data and the activity coefficients for the homogeneous region are shown in Table 3. In these tables, the bubble point (Tb (K)), the composition (mol fraction) of the liquid phases (xi ) and the vapour phase (yi ), as well as the activity coefficients (γ i ) are presented. Table 1 also reports the data for the binary heterogeneous azeotrope water–n-heptane (BIN). Fig. 1 represents the graphical ternary diagram that includes the vapour line and the non-isothermal binodal curve. This figure represents the projection on the ternary composition diagram of the intersection of the single liquid–liquid envelope with the VLE surface. It looks like an isothermal liquid–liquid envelope but it is not. This figure also shows the presence of a ternary azeotrope in the heterogeneous region. The points
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Fig. 1. Equilibrium diagram (mol%) for the ternary system water + ethanol + nheptane at 101.3 kPa.
of the vapour line corresponding to the liquid–liquid equilibrium lines 1–10 are above the tie line. However, the points corresponding to lines 11, 12 and 13 are under the tie line. Therefore, there is a point between lines 10 and 11 where the composition of the vapour coincides with that of a liquid heterogeneous mixture. The composition and temperature of the ternary azeotrope was determined by numerical interpolation as x1 = 0.205, x2 = 0.432, and x3 = 0.363, and 341.83 K, respectively. The composition of the two liquid phases in equilibrium is x1 = 0.030, x2 = 0.195, and x3 = 0.775 in the organic-rich phase, and x1 = 0.341, x2 = 0.614, and x3 = 0.045 in the aqueous-rich phase. Although the azeotropic temperature determined in this work is similar to that reported by Fritzweiler et al. in 1933 [3], the compositions differ significantly (azeotropic composition: x1 = 0.3625, x2 = 0.3035, x3 = 0.3340, T = 341.15 K). The ternary VLE and VLLE experimental data were tested by the point-to-point L–W Wisniak [4] method and found to be thermodynamically consistent. Calculations of the test and the activity coefficients shown in Tables 2 and 3 were done using the program PRO-VLE 2.0 [5]. All the values of L/W are between
Table 1 Vapour–liquid–liquid equilibrium data (mole fraction) for the ternary system water (1) + ethanol (2) + n-heptane (3) at 101.3 kPa Organic phase
BIN 1 2 3 4 5 6 7 8 9 10 11 12 13
Aqueous phase
Vapour phase
Tb (K)
x1
x2
x3
x1
x2
x3
y1
y2
y3
0.001 0.002 0.002 0.003 0.003 0.006 0.007 0.008 0.013 0.018 0.025 0.032 0.036 0.049
0.000 0.019 0.026 0.038 0.048 0.060 0.075 0.084 0.100 0.123 0.164 0.209 0.248 0.294
0.999 0.979 0.971 0.959 0.949 0.934 0.919 0.908 0.887 0.859 0.812 0.759 0.716 0.657
0.9998 0.914 0.845 0.786 0.718 0.672 0.612 0.543 0.487 0.429 0.402 0.320 0.282 0.235
0.000 0.086 0.154 0.213 0.280 0.325 0.383 0.448 0.497 0.547 0.569 0.629 0.646 0.646
0.0002 0.0003 0.0006 0.0010 0.002 0.003 0.006 0.009 0.015 0.025 0.029 0.050 0.072 0.119
0.439 0.313 0.284 0.271 0.261 0.253 0.246 0.238 0.230 0.222 0.215 0.201 0.195 0.190
0 0.257 0.314 0.335 0.357 0.368 0.381 0.394 0.406 0.415 0.423 0.434 0.442 0.449
0.561 0.429 0.403 0.394 0.382 0.379 0.373 0.368 0.365 0.362 0.362 0.364 0.363 0.362
352.39 345.47 343.98 343.34 342.97 342.77 342.60 342.50 342.34 342.25 342.22 341.87 341.89 341.91
Table 2 Activity coefficients of the VLLE data for the ternary system water (1) + ethanol (2) + n-heptane (3) at 101.3 kPa Tb (K)
BIN 1 2 3 4 5 6 7 8 9 10 11 12 13
352.39 345.47 343.98 343.34 342.97 342.77 342.60 342.50 342.34 342.25 342.22 341.87 341.89 341.91
Organic phase
Aqueous phase
γ1
γ2
γ3
γ1
γ2
γ3
970.6 455.16 440.04 287.74 281.6 137.69 115.61 98.31 58.88 41.22 28.78 21.35 18.4 13.16
– 17.16 16.25 12.17 10.42 8.67 7.23 6.7 5.84 4.87 3.72 3.04 2.61 2.23
1.023 0.98 0.97 0.98 0.97 0.99 1 1 1.01 1.05 1.11 1.21 1.27 1.38
0.971 1 1.04 1.1 1.18 1.23 1.32 1.45 1.57 1.73 1.79 2.14 2.35 2.75
– 3.79 2.74 2.17 1.79 1.6 1.42 1.26 1.17 1.09 1.07 1.01 1 1.02
5110.8 956.18 939.8 941.22 462.05 307.69 182.75 100.51 56.23 37.5 30.98 17.97 12.65 7.61
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Table 3 Vapour–liquid equilibrium data (mole fraction) and activity coefficients (γ i ) for the ternary system water (1) + ethanol (2) + n-heptane (3) at 101.3 kPa Tb (K)
343.28 344.70 345.18 347.22 348.64 350.90 348.93 345.82 345.26 347.81 348.61 346.57 350.10 347.56 345.89 345.12 344.38 343.42 343.09 342.89 342.82 342.80 342.77 342.74 342.69 342.68 342.80 342.84 342.86 342.51
Liquid phase
Vapour phase
Activity coefficients
x1
x2
x3
y1
y2
y3
γ1
γ2
γ3
0.174 0.144 0.144 0.138 0.128 0.074 0.227 0.232 0.321 0.315 0.411 0.465 0.071 0.071 0.069 0.066 0.089 0.081 0.074 0.069 0.064 0.061 0.059 0.052 0.048 0.042 0.034 0.034 0.024 0.027
0.716 0.786 0.794 0.831 0.853 0.922 0.762 0.732 0.653 0.669 0.580 0.527 0.919 0.889 0.865 0.845 0.809 0.730 0.679 0.636 0.596 0.567 0.522 0.469 0.431 0.388 0.345 0.299 0.251 0.224
0.109 0.070 0.062 0.031 0.019 0.004 0.012 0.036 0.026 0.016 0.009 0.008 0.009 0.040 0.066 0.090 0.102 0.189 0.248 0.294 0.339 0.372 0.419 0.479 0.521 0.570 0.621 0.667 0.725 0.749
0.145 0.121 0.119 0.116 0.112 0.071 0.177 0.167 0.193 0.208 0.237 0.236 0.080 0.076 0.074 0.074 0.096 0.099 0.104 0.107 0.110 0.113 0.117 0.122 0.125 0.129 0.126 0.132 0.142 0.169
0.510 0.575 0.603 0.685 0.755 0.906 0.707 0.592 0.527 0.616 0.603 0.529 0.870 0.767 0.700 0.671 0.621 0.576 0.556 0.544 0.538 0.533 0.527 0.519 0.513 0.507 0.507 0.497 0.485 0.461
0.345 0.303 0.278 0.199 0.132 0.024 0.116 0.242 0.281 0.176 0.160 0.235 0.050 0.157 0.225 0.255 0.283 0.325 0.340 0.349 0.352 0.354 0.357 0.359 0.362 0.364 0.368 0.371 0.373 0.370
2.67 2.54 2.45 2.29 2.25 2.25 1.97 2.07 1.77 1.75 1.48 1.41 2.73 2.88 3.09 3.33 3.31 3.90 4.55 5.06 5.63 6.07 6.51 7.71 8.57 10.11 12.14 12.69 19.32 20.72
0.98 0.95 0.97 0.97 0.99 1.00 1.02 1.01 1.03 1.06 1.16 1.22 1.00 1.00 1.01 1.02 1.01 1.08 1.14 1.20 1.27 1.32 1.42 1.56 1.69 1.85 2.07 2.34 2.72 2.94
7.51 9.91 10.07 13.48 14.04 10.73 20.96 14.72 24.13 22.67 35.70 63.04 9.58 8.16 7.51 6.45 6.40 4.10 3.32 2.87 2.52 2.31 2.07 1.83 1.70 1.56 1.44 1.35 1.25 1.21
0.97 and 0.98. The test did not reveal any significant inconsistency in the data. Vapour pressures of the three components were calculated using the Antoine equation, with parameters Ai , Bi , and Ci taken from literature [6,7] and shown in Table 4. VLE data correlation. The 13 ternary experimental VLLE data obtained (treated as 26 VLE data), the 30 experimental VLE data of the homogeneous region, together with binary data from the bibliography [8,9] have been correlated using the UNIQUAC and NRTL models. Only some points of each binary have been used to do the correlation. All the calculations have been made using CHEMCAD V [10]. Fig. 2 shows the graphical representation of three different experimental tie lines compared with those obtained using UNIQUAC and NRTL. Results of the tie lines obtained by UNIFAC (original) are also shown. The slope of the tie lines calculated using NRTL and UNIFAC is similar to the experimental one in the bottom part of the heterogeneous
region. In the top part, the experimental tie line coincides better with those obtained with UNIQUAC. In all cases, it can be observed that the experimental heterogeneous region is smaller than the calculated one that even contains some homogeneous experimental points. It is important to notice that the calculation of the equilibrium compositions of this region close to the
Table 4 Antoine equation parametersa of the pure substances Compound
A
B
C
Temperature range (K)
Water Ethanol n-Heptane
7.19621 7.2371 6.0273
1730.63 1592.864 1268.115
−39.724 −46.966 −56.25
+274/+373 +293/+366 –
a
Antoine equation: log(P) = A − B/[T + C], with: P (Pa) and T (K).
Fig. 2. Experimental and calculated LLE at 101.3 kPa. Parameters of the models obtained from correlation of VLE data.
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simulate the behaviour of the decanter since the prediction of the experimental data is better. However, none of these methods properly correlate and approximate the top zone of the nonisothermal solubility curve, which is important for the simulation of the azeotropic distillation. Acknowledgements The authors thank the European Community (FEDER Funds) and the DGICYT of Spain for the financial support of project PPQ2003-03385. References
Fig. 3. Experimental and calculated LLE at 101.3 kPa. Parameters of the models obtained from correlation of LLE data.
azeotrope is important for the simulation of ethanol dehydration. As seen in Fig. 2, the use of these models may lead to inaccurate predictions of the equilibrium and important differences in the simulation results. LLE data correlation. In order to improve these calculations, model parameters have been determined regressing exclusively LLE data. The non-isothermal VLLE experimental data treated as 13 LLE experimental tie lines have been correlated using UNIQUAC and NRTL models. Fig. 3 shows the representation of three experimental tie lines and those obtained by means of UNIFAC-LLE (original), UNIQUAC and NRTL and the fact that there is not too much improvement correlating only the tie lines. The top part of the heterogeneous region is not correlated with sufficient precision although these parameters could be used to
[1] V. Gomis, F. Ruiz, J.C. Asensi, Fluid Phase Equilib. 172 (2000) 245–259. [2] B.W. Mangum, G.T. Furukawa, US Department of Commerce, National Institute of Standards and Technology, Springfield, 1990. [3] J. Gmehling, J. Menken, J. Krafczyk, K. Fischer, Azeotropic Data, VCH, Weinheim, 1994. [4] J. Wisniak, Ind. Eng. Chem. Res. 32 (1993) 1531–1533. [5] M. Elly, M. Landa, J. Wisniak, PRO-VLE 2.0. Vapor–Liquid Equilibria Computer Program, Ben-Gurion University of the Negev, Beer-Sheva, 2003. [6] C.L. Yaws, Chemical Properties Handbook, McGraw-Hill, New York, 1999. [7] R.M. Felder, R.W. Rousseau, Principios Elementales de los Procesos Qu´ımicos, 2nd ed., Addison-Wesley Iberoamericana, Wilmington, 1986. [8] J.D. Raal, R.K. Code, D.A. Best, J. Chem. Eng. Data. 17 (1972) 211–221, VLE data extracted from: J. Gmehling, U. Onken, Chemistry Data Series, DECHEMA, Dortmund, 1977. [9] V.N. Stabnikov, B.Z. Matyushev, T.B. Protsyuk, N.M. Yushchenko, Pishch. Prom. (Kiev) 15 (1972) 49, VLE data extracted from: J. Gmehling, U. Onken, Chemistry Data Series, DECHEMA, Dortmund, 1977. [10] CHEMCAD V, Process Flow Sheet Simulator, Chemstations Inc., Houston, 2002.