Thermodynamics in Separation for the Ternary System 1, 2

Article pubs.acs.org/IECR

Thermodynamics in Separation for the Ternary System 1,2Ethanediol + 1,2-Propanediol + 2,3-Butanediol Yang Zhong, Yanyang Wu,* Jiawen Zhu, Kui Chen, Bin Wu, and Lijun Ji Chemical Engineering Research Center, East China University of Science & Technology, Shanghai 200237, China S Supporting Information *

ABSTRACT: The thermodynamic data in distillation, i.e., isobaric vapor−liquid equilibrium data, have been measured for the three binary systems of 1,2-ethanediol + 1,2-propanediol, 1,2-ethanediol + 2,3-butanediol, and 1,2-propanediol + 2,3-butanediol at 101.3 kPa. The accuracy of the experimental data was confirmed by both Herington test and Van Ness test. The thermodynamic properties of the vapor phase have been calculated with the Hayden−O’Connell equation in consideration of nonideality. The liquid activity coefficients have been satisfactorily correlated with the Wilson, NRTL, and UNIQUAC models, while the estimated results from the UNIFAC model are not satisfactory. The corresponding binary interaction parameters of the three models Wilson, NRTL, and UNIQUAC were calculated and used to obtain the residual curves of the ternary system 1,2ethanediol + 1,2-propanediol + 2,3-butanediol. Based on all of the preceding results, a two-column distillation process has been designed to obtain the required products.

1. INTRODUCTION Glycols, including ethanediol (EG), propanediol (PG), and butanediol (BD), etc., are important chemical raw materials and widely used in many fields such as chemical engineering, pharmaceutical, textile, polymer materials, and so on. Glycols can be produced via chemical synthesis and catalytic hydrocracking of sorbitol obtained from cereal enzymolysis.1 The catalytic hydrocracking of sorbitol provides the possibility of an energy substitute in a time of energy crisis.2 Because the product obtained from catalytic hydrocracking is a mixture of EG, PG, and BD, etc., it is essential to separate and purify such glycols. Apparently, the separation of the three diols is not an easy thing due to high similarities of their chemical structure, boiling point, polarities, and so on. It is hard to find an appropriate solvent for extraction and an adsorbent for adsorption. Distillation as the most commonly used method for liquid solution has been considered in this study. However, it needs to be studied in detail and designed carefully because of their adjacent and high boiling points. Thermodynamics data are essential for distillation design, operation, and optimization. However, no available data for isobaric vapor−liquid equilibrium (VLE) of the binary mixtures 1,2-EG + 1,2-PG, 1,2-EG + 2,3-BD, and 1,2-PG + 2,3-BD at 101.3 kPa have been found in the open literature until now. The present work has focused on isobaric VLE of the three binary mixtures. The Hayden−O’Connell equation3 was used to calculate the fugacity coefficients considering the nonideality in the vapor phase, and the Wilson,4 NRTL,5 and UNIQUAC6 models were used to describe the nonideality in the iquid phase. The accuracy of experimental data was confirmed by both Herington test7 and Van Ness test.8 The interaction parameters for Wilson, NRTL, and UNIQUAC models of the constituent binaries were regressed to calculate the residue curve map (RCM) of the ternary system 1,2-EG + 1,2-PG + 2,3-BD with Aspen Plus. According to RCM, purified product can be acquired after distillation with certain ternary feed. © 2014 American Chemical Society

Based on the above, a feasible distillation process has been designed in this work.

2. EXPERIMENTAL SECTION 2.1. Materials. 1,2-EG (>99.9 wt %), 1,2-PG (>99.9 wt %) ,and 2,3-BD (>99.0 wt %) were provided by Sinopharm Chemical Reagent Co., Ltd. All chemicals were used directly without any further treatment in this study. The purity of these materials was detected and assured by gas chromatography with flame ionization detector (FID). The boiling points and densities of these chemicals were measured and shown in Table 1 in comparison with the literature data. Table 1. Densities at 293.75 K and Boiling Points at 101.3 kPa densities (kg/m3)

boiling points (K)

component

experimental

literature

experimental

literature

1,2-EG 1,2-PG 2,3-BD

1113 1035 1001

1134a 1038c 994d

469.71 460.22 453.08

470.45b 460.45c 453.85e

a Taken from ref 9. bTaken from ref 10. cTaken from ref 11. dTaken from ref 12. eTaken from the Aspen properties databank.

2.2. Methods. In this work, the modified Ellis recirculation equilibrium still, described by Walas,13 was used to take measurements. The temperature was measured with a precise mercury thermometer, which was calibrated by a standard mercury thermometer. The uncertainty in the temperature measurements was ±0.1 K. Received: Revised: Accepted: Published: 12143

May 17, 2014 June 20, 2014 June 30, 2014 June 30, 2014 dx.doi.org/10.1021/ie501968g | Ind. Eng. Chem. Res. 2014, 53, 12143−12148

Industrial & Engineering Chemistry Research

Article

The pressure was maintained at 101.3 kPa with the aid of a pressure control system, which includes one vacuum pump, one pressure sensor, one electromagnetic valve, one electronic relay, and one reservoir. The reservoir was linked between the pressure gauge and the system. At first, the air went into the system by opening the atmosphere vent. It was regarded as 101.3 kPa when the pressure gauge level on both sides became even. When the pressure in the reservoir was not 101.3 kPa, the electronic relay started working; meanwhile the electromagnetic valve opened, and then the pressure in the reservoir was adjusted to 101.3 kPa by the vacuum pump removing or injecting air. When the pressure of the reservoir reached 101.3 kPa again, the electronic relay stopped working and then the electromagnetic valve was closed. The uncertainty in the pressure was ±0.05 kPa. The vapor (condensate) and liquid compositions were analyzed by a gas chromatograph (GC-9790) with a FID, all produced by Zhejiang Fuli Co., Ltd. The GC column was a SE30 capillary column (50 m × 0.25 mm i.d. × 0.25 μm). Highpurity nitrogen (99.9999% purity) was used as the carrier gas at a constant flow rate of 30 mL/min. Column temperature was controlled by first-order ascent function. Each sample was measured at least three times to ensure accuracy. The uncertainty in mass fraction was within ±0.1%.

ln φi = (2 ∑ yj Bij − Bm)P /RT

Bm =

ln(PiS/Pa) = A + B /T + C ln T + DT E

Table 3. Coefficients for Vapor Pressure (P°/Pa)a A B C D E a

a

1,2-PG

2,3-BD

460.75 626 6.1 0.239 0.28 2.0986 3.154 3.0824 2.784 1.1065

453.85 611 5.1 0.267 0.27 3.6275 3.371 3.7561 3.32 1.1055

1,2-EG

1,2-PG

2,3-BD

84.09 −10411 −8.198 1.65 × 10−18 6

212.80 −15420 −28.109 2.16 × 10−05 2

112.25 −11021 −12.792 8.86 × 10−06 2

Taken from the Aspen properties databank.

3.2. Thermodynamic Consistency Test. The thermodynamic consistency test for binary systems was based on the Gibbs−Duhem equation15 and verified with the Herington test and Van Ness test. As Herington suggested, if |D − J| < 10, the pertinent experimental isobaric VLE data can be considered thermodynamically consistent. D and J were expressed as follows: 1

D = 100

|∫ ln(γ1/γ2) dx1| 0 1

∫0 |ln(γ1/γ2)| dx1 J = 150

Tmax − Tmin Tmin

(5)

(6)

where Tmax and Tmin are the highest and lowest temperature in the system, respectivly. In this work, the values of |D − J| for the three binary systems of 1,2-EG + 1,2-PG, 1,2-EG + 2,3-BD, and 1,2-PG + 2,3-BD were 8.02, 4.80, and 0.85, respectively, which meant that VLE results for these systems passed the Herington test. Meanwhile, the experimental data have also been inspected by the Van Ness test modified by Fredenslund,16 and the equation is expressed as

Table 2. Physical Properties of Pure Component Used in This Worka 470.45 720 8.2 0.191 0.262 2.4103 2.564 2.4087 2.248 0.5068

(4)

where T is the absolute temperature and A, B, C, D, and E are the component specific coefficients for vapor pressure, which are listed in Table 3.

(1)

1,2-EG

(3)

j

where Bii and Bij are the second virial parameter and interactive virial parameter, respectively, and were calculated by the Hayden−O’Connell equation due to the polarity of measured systems. The saturation vapor pressure was calculated by the following equation:14

where yi and xi are the mole fractions of component i in vapor and liquid phases, respectively. P is the total pressure, R the gas constant, and T the temperature. Ps stands for the saturation vapor pressure of pure component. φ̑ Vi and φSi are the fugacity coefficients of component i in vapor phase and pure state, respectively. γi is the activity coefficient of component i in liquid phase. VLi is the molar volume of pure liquid i. The physical properties used in this work can be seen in Table 2. The fugacity coefficients in vapor phase were calculated by the following equations:

Tb/K Tc/K Pc/MPa Vc/(m3·kmol−1) Zc DM/D RD/A r q ω

∑ ∑ yyi j Bij i

3. RESULTS AND DISCUSSION 3.1. Experimental Data. The isobaric VLE data and calculated activity coefficients for three binary systems 1,2-EG + 1,2-PG, 1,2-EG + 2,3-BD, and 1,2-PG + 2,3-BD were listed in the Supporting Information as Tables S1−S3. The activity coefficients γi were calculated with the following equation: ⎧ V L(P − P S) ⎫ i ⎬ yi φȋ V P = xiγiPiSφiS exp⎨ i RT ⎩ ⎭

(2)

j

N

AAD(y) = (1/N ) ∑ |yiexp − yical | i=1

(7)

where N is the number of experimental data; the superscript exp denotes experimental and cal denotes calculated with the thermodynamic model. AAD(y) between calculated and experimental y must be smaller than 0.01 according to the test. And the AAD(y) for systems 1,2-EG + 1,2-PG, 1,2-EG + 2,3-BD, and 1,2-PG + 2,3-BD were calculated and listed in

Taken from the Aspen properties databank. 12144

dx.doi.org/10.1021/ie501968g | Ind. Eng. Chem. Res. 2014, 53, 12143−12148


Article

Table 4. Correlated Interaction Parameters, AAD, and ARD% between Calculated and Experimental Values model

A12a/(J·mol−1)

A21a/(J·mol−1)

Wilson NRTL UNIQUAC

288.6609 −271.5009 −118.6912

513.0628 313.1567 106.1660

Wilson NRTL UNIQUAC

−922.8968 344.3529 −257.2794

−704.2941 −224.4070 170.5381

Wilson NRTL UNIQUAC

1146.0345 −203.0119 82.9210

888.7044 251.0647 −99.1084

α

AADb(y1)

1,2-EG (1) + 1,2-PG (2) 0.0007 0.3 0.0007 0.0006 2,3-BD (1) + 1,2-EG (2) 0.0027 0.3 0.0027 0.0028 2,3-BD (1) + 1,2-PG (2) 0.0009 0.3 0.0009 0.0009

AADb(T)

ARD%c(γ1)

ARD%c(γ2)

0.0309 0.0348 0.0406

0.1993 0.2508 0.2316

0.3723 0.3808 0.3192

0.0267 0.0265 0.0274

0.7700 0.7600 0.7700

0.9199 0.9212 0.9179

0.0186 0.0182 0.0186

0.1974 0.1923 0.1997

0.2827 0.2899 0.2837

cal N exp Wilson, Aij = (λij − λii)/R; NRTL, Aij = (gij − gii)/R; UNIQUAC, Aij = (Uij − Uii)/R. bAAD(y) = (1/N)ΣNi=1|yexp i − yi |; AAD(T) = (1/N)Σi=1|Ti − cal c N exp cal exp N exp cal exp Ti |. ARD%(γ1) = (100/N)Σj=1|γ1,j − γ1,j |/γ1,j ; ARD%(γ2) = (100/N)Σj=1|γ2,j − γ2,j |/γ2,j . a

Table 4. As can be seen from this table, the experimental data passed the Van Ness test. 3.3. Data Regression. In this work, the experimental vapor−liquid equilibrium data were correlated by minimizing object function F,17 which is defined by the sum of squared differences between calculated and experimental data as follows: ⎡⎛ exp cal ⎞2 ⎛ P exp − P cal ⎞2 ⎢⎜ T j − T j ⎟ j ⎜ j ⎟ F = ∑ ⎢⎜ exp ⎟ + ⎜ σ exp ⎟ σ T P ⎠ ⎝ ⎠ j = 1 ⎢⎝ ⎣ N

⎛ T exp − T cal ⎞2 j,i j,i ⎟ + + ∑ ⎜⎜ exp ⎟ σ , x i ⎝ ⎠ i=1 C

⎛ y exp − y cal ⎞2 ⎤ ⎥ ∑ ⎜⎜ j ,i exp j ,i ⎟⎟ ⎥ σy , i i=1 ⎝ ⎠ ⎥⎦ C

(8)

where N is the number of the experimental data. C is the number of components. σT, σP, σx, and σy refer to calculated standard deviations for T, P, x, and y, respectively. The regression was carried out by Aspen Plus V7.2 chemical process simulator. The interaction parameter α in NRTL was fixed at 0.3, as recommended by Renon and Prausnitz.5 The Wilson, NRTL, and UNIQUAC models were used to regress interaction parameters, and the average absolute deviation (AAD) and the percent average relative deviation (ARD%) between the calculated and experimental values are listed in Table 4. The comparison indicated that NRTL model was more suitable to calculate the systems in this study than others did, while all of the models fit the experimental VLE data in a similar way. In the view of industrial application, the calculation of VLE data could be done with all three models. The experimental VLE data of the three binary systems were plotted in Figures 1−3, and the solid lines in these figures were calculated by the NRTL model. Besides the preceding experimental way, the VLE data were also estimated with the UNIFAC model. The estimated thermodynamics, including boiling temperature Tcal, vapor composition y1,cal, activity coefficients γ1,cal, and γ2,cal were calculated from UNIFAC model, while the liquid compositions x1 were set as experimental values. The estimated values of Tcal, y1,cal, γ1,cal, and γ2,cal of the three systems were listed in the Supporting Information as Tables S4−S6, and the deviations of these values are shown in Table 5. As we can see from the table, the estimated results are not satisfactory. Then the UNIFAC

Figure 1. T−x−y diagram for 2,3-BD(1) + 1,2-EG(2) at 101.3 kPa: (●, ○) experimental data; () NRTL equation.

Figure 2. T−x−y diagram for 2,3-BD (1) + 1,2-PG (2) at 101.3 kPa: (●, ○) experimental data; () NRTL equation.

model was not used to design subsequent distillation in this study. 3.4. RCMs of Ternary System 1,2-EG + 1,2-PG + 2,3BD. The interaction parameters regressed in this study were 12145



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Under total reflux or high reflux ratio, the residue curves become equivalent to the distillation trajectories in tray or packed columns.21 It can be used to preliminarily estimate the feasible distillation region by adding a material balance line to the RCM. As shown in Figure 4, all of the curves were in the same area, so there is no distillation boundary. The liquid component in the boiler along the distillation tends to be 1,2EG within the last liquid drop. RCMs are closely related to composition profiles in distillation columns. According to the composition to be separated in the ternary system, feed composition is set as point F in Figure 4, in which 1,2-EG is 23%, 1,2-PG 20%, and 2,3-BD 57%, respectively. The relevant distillation was conducted under high reflux ratio. If the overhead was pure 2,3-BD, the bottom composition could be at point A. Similarly, if the bottom was pure 1,2-EG, the overhead composition could be at point C. The areas ABF̑ and CFD̑ surrounded by the two material balance lines AF and FC and the distillation line through F were defined as possible product composition regions. According to Figure 4, normal distillation could be taken to separate the ternary system in this study. 3.5. Distillation Process Design. The separation process was designed with Aspen Plus. As shown in Figure 5, the

Figure 3. T−x−y diagram for 1,2-EG (1) + 1,2-PG (2) at 101.3 kPa: (●, ○) experimental data; () NRTL equation.

Table 5. AAD and ARD% between Calculated and Experimental Data with UNIFAC Model system

AAD(y1)

AAD(T)

ARD%(γ1)

ARD%(γ2)

1,2-EG (1) + 1,2-PG (2) 2,3-BD (1) + 1,2-EG (2) 1,2-BD (1) + 1,2-PG (2)

0.0129 0.0167 0.0013

1.6614 1.7862 0.6397

7.4945 8.1922 0.5821

9.3451 6.3725 0.2330

used to estimate the boiling point of the ternary system and to verify the feasibility of separation. RCMs play an important role in ternary diagrams pertaining to separations.18 Residue curves are defined as the trace of liquid composition that remained in a still with time in a batch distillation.19 Its trajectories flow from the lowest to the highest boiling compounds.20 During the early stages of distillation design, residue curves are often used to provide insight into the behavior of the distillation. In this work the RCMs at 101.3 kPa were computed by Aspen Plus with NRTL−HOC equation based on the regressed interaction parameters. As illustrated in Figure 4, there existed no binary or ternary azeotropic point in the system 1,2-EG + 1,2-PG + 2,3-BD.

Figure 5. Flow diagram for design.

distillate (product 1) from column 1 was crude 2,3-BD and the bottom flow (mixture) is comprised of 1,2-PG and 1,2-EG. Then the mixture was sent to column 2 for further separation, and the overhead (product 2) was 1,2-PG and the bottom (product 3) was 1,2-EG. In this design, feed concentrations were set as that obtained from practical industry. The purities of target products were set as 98% for both 2,3-BD and 1,2-PG, and 99% for 1,2-EG, which are the common concentrations in commercial use. Column 1 was designed with 61 stages in consideration of product purity, reflux ratio, stage number, and the heat duty of condenser and reboiler. The specific operation parameters were taken as feed at the 30th stage and mass reflux ratio 5.9. Column 2 was designed with 60 stages with mass reflux ratio 8.1 and optimum feed stage as the 31st. The sensitivity analysis indicated that the ratio of distillate to feed exerted a great influence on the distillate purity, which is shown in Figure 6a,b. In order to concentrate 2,3-BD in product 1 and decrease 2,3-BD in the mixture, the distillate to feed fraction in column 1 was set as 0.579, based in Figure 6. Similarly, for balancing 1,2-PG in product 2 and product 3 in column 2, the distillate to feed fraction in column 2 was set as 0.451. The results for all of the streams in column 1 and column 2 are shown in Table 6.

Figure 4. RCs for the 1,2-EG + 1,2-PG + 2,3-BD ternary system at 101.3 kPa simulated by Aspen Plus using NRTL−HOC model with the parameters given in Table 4. 12146



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ASSOCIATED CONTENT

S Supporting Information *

Tables of experimental VLE data at 101.3 kPa and of estimated VLE data with UNIFAC for three binary systems of 1,2-EG (1) + 1,2-PG (2), 2,3-BD (1) + 1,2-EG (2), and 2,3-BD (1) + 1,2PG (2). This material is available free of charge via the Internet at http://pubs.acs.org.

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AUTHOR INFORMATION

Corresponding Author

*Tel.: +86 21 64253914. Fax: +86 21 64253914. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The financial support of the National Natural Science Foundation of China (Grant No. 21106039) is gratefully acknowledged.

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(1) Tronconi, E.; Ferlazzo, N.; Forzatti, P.; Pasquon, I.; Casale, B.; Marini, L. A mathematical model for the catalytic hydrogenolysis of carbohydrates. Chem. Eng. Sci. 1992, 47 (9), 2451−2456. (2) Lancia, A.; Musmarra, D.; Pepe, F. Vapor-liquid equilibria for mixtures of ethylene glycol, propylene glycol, and water between 98 and 122°C. J. Chem. Eng. Jpn. 1996, 29 (3), 449−455. (3) Hayden, J. G.; O’Connell, J. P. A generalized method for predicting second virial coefficients. Ind. Eng. Chem. Process Des. Dev. 1975, 14 (3), 209−216. (4) Wilson, G. M. Vapor-liquid equilibrium. XI. A new expression for the excess free energy of mixing. J. Am. Chem. Soc. 1964, 86 (2), 127− 130. (5) Renon, H.; Prausnitz, J. M. Local compositions in thermodynamic excess functions for liquid mixtures. AIChE J. 1968, 14 (1), 135−144. (6) Abrams, D. S.; Prausnitz, J. M. Statistical thermodynamics of liquid mixtures: A new expression for the excess Gibbs energy of partly or completely miscible systems. AIChE J. 1975, 21 (1), 116−128. (7) Herington, E. F. G. Tests for the Consistency of Experimental Isobaric Vapor-Liquid Equilibrium Data. J. Inst. Pet. 1951, 37, 457− 470. (8) Van Ness, H. C.; Byer, S. M.; Gibbs, R. E. Vapor-Liquid equilibrium: Part I. An appraisal of data reduction methods. AIChE J. 1973, 19 (2), 238−244. (9) Chiavone-Filho, O.; Proust, P.; Rasmussen, P. Vapor−liquid equilibria for glycol ether + water systems. J. Chem. Eng. Data 1993, 38 (1), 128−131. (10) Lide, D. R. CRC Handbook of Chemistry and Physics, 87th ed.; CRC Press, Taylor and Francis: Boca Raton, FL, USA, 2007. (11) Doherty, M. F.; Malone, M. F. Conceptual Design of Distillation Systems; McGraw-Hill: New York, 2001. (12) Knowlton, J. W.; Schieltz, N. C.; Macmillan, D. Physical chemical properties of the 2,3-BDs. J. Am. Chem. Soc. 1946, 68 (2), 208−210. (13) Walas, S. M. Phase Equilibria in Chemical Engineering; Butterworths: London, U.K., 1985. (14) Daubert, T. E.; Danner, R. P. Physical and Thermodynamic Properties of Pure Chemicals; Hemisphere: New York, 1989. (15) Smith, J. M.; Van Ness, H. C.; Abbott, M. M. Introduction to Chemical Engineering Thermodynamics, 6th ed.; McGraw-Hill: New York, 2001. (16) Fredenslund, A., Gmehling, J., Rasmussen, P. Vapor−Liquid Equilibria Using UNIFAC; Elsevier: Amsterdam, 1977. (17) Prausnitz, J.; Anderson, T.; Grens, E.; Eckert, C.; Hsieh, R.; O’Connell, J. Computer Calculations for Multicomponent Vapor-Liquid

Figure 6. Relation between distillate to feed fraction and mass fraction of (a) 2,3-BD in product 1 and mixture in column 1 and (b) 1,2-PG in product 2 and product 3 in column 2.

Table 6. Specifications for Flows in Column 1 and Column 2 column 1 component mass fraction 2,3-BD 1,2-PG 1,2- EG mass flow (kg/h) temperature (K)

■

column 2

feed

mixture

product 1

0.57 0.20 0.23 100.00 293.00

0.98 0.02 0.00 57.80 453.22

0.00 0.45 0.55 42.20 465.67

product 2 product 3

0.02 0.98 0.00 19.03 460.07

REFERENCES

0.00 0.01 0.99 23.17 469.64

CONCLUSION The thermodynamics was studied in detail in order to make an appropriate distillation process design in this paper. The isobaric VLE data for the three binary mixtures 1,2-EG + 1,2PG, 1,2-EG + 2,3-BD, and 1,2-PG + 2,3-BD were measured at 101.3 kPa. Reasonable thermodynamic consistency was obtained according to both Herington and Van Ness tests. The experimental data were regressed by Wilson, NRTL, and UNIQUAC models, while the estimated results from the UNIFAC model are not satisfactory. The three regressed binary groups of interaction parameters were used to calculate residue curves. Based on all of the preceding results, a two-column distillation process was designed and optimized to obtain the required products with specified feed concentration. Product amounts of 2,3-BD, 1,2-PG, and 1,2-EG can be up to 98%, 98%, and 99%, respectively, after the designed process. 12147



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and Liquid-Liquid Equilibria; Prentice-Hall: Englewood Cliffs, NJ, USA, 1980. (18) Partin, L. R. Use Graphical Techniques to Improve Process Analysis. Chem. Eng. Prog. 1993, 89 (1), 43−48. (19) Doherty, M. F.; Perkins, J. D. On the Dynamics of Distillation Processes I: The Simple Distillation of Multi-Component NonReacting, Homogeneous Liquid Mixtures. Chem. Eng. Sci. 1978, 33 (3), 281−301. (20) Lucia, A.; Taylor, R. The Geometry of Separation Boundaries: I. Basic Theory and Numerical Support. AIChE J. 2006, 52 (2), 583− 594. (21) Taylor, R.; Baur, R.; Krishna, R. Influence of Mass Transfer in Distillation: Residue Curves and Total Reflux. AIChE J. 2004, 50 (12), 3134−3148.

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