Isobaric VaporâLiquidâLiquid Equilibria for the Ternary Systems

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Isobaric Vapor−Liquid−Liquid Equilibria for the Ternary Systems Ethanol + Water + Propyl Acetate and 1‑Propanol + Water + Propyl acetate Jordi Pla-Franco, Estela Lladosa, Sonia Loras,* and Juan B. Montón Departamento de Ingeniería Química, Escuela Técnica Superior de Ingeniería, Universitat de València, Burjassot, 46100 Valencia, Spain ABSTRACT: In order to break azeotropes appearing in mixtures of water and either ethanol or 1-propanol, an azeotropic distillation process is proposed using propyl acetate as entrainer. Therefore, isobaric vapor−liquid− liquid equilibrium (VLLE) data at atmospheric pressure have been obtained for the ternary systems ethanol + water + propyl acetate and 1-propanol + water + propyl acetate. Data correlation has been carried out in four different ways and a comparison between the results obtained in each correlation has been done. The set of parameters obtained from the best correlation has been used in the VLLE estimation. Effects of using propyl acetate as entrainer in the azeotropic distillation have been discussed, and a three-column distillation sequence to separate 1-propanol and water has been proposed.

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INTRODUCTION Systems formed by water and some alcohols such as ethanol or 1-propanol exhibit azeotropy under certain conditions.1,2 For example, in the case of mixtures of ethanol and water, an azeotrope appears at a temperature of 351.33 K and an ethanol molar fraction of 0.893 at 101.3 kPa.3 Because of this behavior, the conventional distillation process is limited to preconcentration, approaching the composition of the binary azeotrope. Therefore, the use of other techniques is inevitable. Examples of these alternative processes are pervaporation,4−6 adsorption,7−9 liquid− liquid extraction, 10 pressure-swing distillation, 11 extractive distillation12−15 (ED), and azeotropic distillation16 (AD). In a case of large scale production, ED and AD are the preferred methods for carrying out alcohol dehydration.17 Although these techniques have high energy consumption, the other options have some critical drawbacks.18 In the AD process, separation is based on the azeotropic property of the system. In this technique, a compound known as entrainer is added to the original mixture to achieve a desired azeotropic condition. Depending on the system deviations from Raoult’s Law, the azeotropic system can be homogeneous, with a unique liquid phase, or heterogeneous, with two liquid phases. In the latter case, a detailed study of the heterogeneous region is essential to plan the distillation sequence. The best way to dispose of accurate heterogeneous zone information is by experimental vapor−liquid−liquid equilibrium (VLLE) data acquisition. This paper continues our thermodynamic research, which is focused on the separation of alcohol and water mixtures using various solvents. In a first work,19,20 an ED process to separate ethanol and water using glycerol or ethylene glycol as solvent was proposed. In a second work,21 behavior of 2-methoxyethanol © XXXX American Chemical Society

(commercially called methyl cellosolve) as entrainer in an ED process to separate ethanol + water and 1-propanol + water mixtures was discussed. In our recent work,22 an AD process to separate mixtures of ethanol and water using diisopropyl ether as solvent was proposed and analyzed. Now, the use of propyl acetate as solvent in an azeotropic distillation process is proposed in order to separate ethanol + water and 1-propanol + water mixtures. As noted above, it is very important to have VLLE data to analyze the behavior of the potential candidate entrainer. In this way, isobaric VLLE data at 101.3 kPa for the ternary systems ethanol (1a) + water (2) + propyl acetate (3) and 1-propanol (1b) + water (2) + propyl acetate (3) were obtained. As far as we know, neither LLE nor VLLE data have been previously published for the ternary system ethanol + water + propyl acetate. However, the ternary system 1-propanol + water + propyl acetate has attracted greater attention. However, most published studies23,24 about this system are based on data of the liquid−liquid equilibrium (LLE). Only Smirnova et al.25 obtained isobaric VLLE data at several pressures. As there is no consensus in the literature about the best way to obtain parameters of the NRTL26 and UNIQUAC27 local composition models, four different procedures were used to correlate the equilibrium data obtained in this work. Results of each method were compared and binary parameters for the mentioned models were obtained using the correlation type that gave best results. These parameters were subsequently Received: February 27, 2014 Accepted: April 23, 2014

A

dx.doi.org/10.1021/je500191j | J. Chem. Eng. Data XXXX, XXX, XXX−XXX

Journal of Chemical & Engineering Data

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coming from the separation chamber of the equipment is collected in a test tube once equilibrium has been reached. The test tube is completely sealed and coated with an isolating material. In this tube, the dispersed liquid phases are separated into two phases (after approximately 1 day) at constant temperature (slightly lower than the mixture bubble point to prevent partial vaporization of the sample). For that purpose the tube is placed in a Block Heater SBH200DC from Stuart Scientific. A sample of each liquid phase was taken and placed in a vial with a small amount of an internal standard (1-propanol in the case of ethanol+water+propyl acetate mixture and ethanol in the case of 1-propanol + water + propyl acetate mixture). The addition of the internal standard prevents phase separation effects when changing the temperature after separation of phases. The composition of the sampled liquid phases and vapor phase were determined using a CE Instruments GC 8000 Top GC after calibration with gravimetrically prepared standard solutions. A thermal conductivity detector (TCD) was used together with 80/100 Porapak Q 3 m × 1/8 in. The GC response peaks were treated with Chrom-Card for Windows. In both cases, optimum operation conditions were: injection temperature, 473 K; oven temperature, 373 K; detector temperature, 433 K; detector current, 220 mA; helium flow rate, 40 mL/min. Very good peak separation was achieved under these conditions and calibration analyses were used to convert the peak area ratio to the mass composition of the sample. The standard deviation in the mole fraction was usually less than 0.002.

used to describe the azeotropic distillation used to dehydrate ethanol or 1-propanol.

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EXPERIMENTAL SECTION Chemicals. Ethanol (>0.995 mass fraction, for analysis ACS) and water (>0.990 mass fraction, water for chromatography LiChrosolv) were supplied by Merck. On the other hand, both 1-propanol (>0.995 mass fraction) and propyl acetate (>0.990 mass fraction) were supplied by Sigma-Aldrich. All chemicals used in the experiments are listed in Table 1. The reagents were

Table 1. Specifications of Chemical Samplesa chemical name ethanol 1-propanol water propyl acetate a

source

initial mass fraction purity

purification method

Merck Aldrich Merck Aldrich

0.9950 0.9900 0.9900 0.9900

none none none none

final mass fraction purity

analysis method GCa GCa GCa GCa

Gas chromatography.

used without further purification after chromatography failed to show any significant impurities. The water content, determined using a Karl Fischer volumetric automatic titrator (Metrohm, 701 KF Titrino), was small in all chemicals (0.990 mass fraction). Analysis. Sampling was carried out using a specific method depending on the sample equilibrium phase. The sampling of the vapor phase was carried out using a VW Type, 6-port heated valve from Valco Instruments Co., which injected the sample directly into the gas chromatograph (GC). In order to prevent vapor condensation, the wall of the connecting tube is overheated by a resistance tape controlled by a digital potentiometer. This assembly keeps the temperature 20 °C higher than the dew temperature of the vapor sample during all the vapor recirculation process. Due the presence of two phases in the liquid phase, sampling of the liquid requires a more complex system. In this case, a small amount of the liquid

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RESULTS AND DISCUSSION Experimental Data. Ternary VLLE data were obtained for the two studied systems: ethanol + water + propyl acetate and 1-propanol + water + propyl acetate. All measurements were made under isobaric conditions at 101.3 kPa. Experimental values for each system are listed in Tables 2 and 3, respectively. Table 2. VLLE Data for Ethanol (1a) + Water (2) + Propyl Acetate (3) Ternary System at 101.3 kPaa organic rich phase, xIi

aqueous rich phase, xIIi

vapor phase, yi

xI1a

xI2

xII1a

xII2

y1a

y2

T

0.217 0.200 0.221 0.207 0.186 0.145 0.088 0.162 0.067 0.035

0.465 0.396 0.523 0.413 0.353 0.273 0.219 0.300 0.196 0.160

0.081 0.070 0.098 0.071 0.060 0.043 0.030 0.050 0.024 0.011

0.907 0.919 0.884 0.919 0.931 0.947 0.964 0.943 0.969 0.982

0.262 0.256 0.273 0.242 0.221 0.178 0.199 0.185 0.093 0.056

0.428 0.429 0.427 0.432 0.438 0.451 0.445 0.449 0.488 0.504

351.74 351.78 351.62 351.90 352.63 352.77 352.80 352.93 353.84 354.70

K

a

u(T) = 0.02 K, u(x) = 0.002, and u(y) = 0.002.

In these tables, the superscripts I and II indicate the organicrich phase and the aqueous-rich phase, respectively. Isobaric tie lines of the two liquid phases and vapor phase mole fractions corresponding to each system have been plotted in Figures 1 (for the ethanol + water + propyl acetate system) and 2 (for the 1-propanol + water + propyl acetate system). In order to compare acquired experimental data, the Smirnova et al.25 data have also been plotted in Figure 2. As can be seen, there are no significant differences between our data and the data of Smirnova et al. B



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concentration in the vapor phase is increased. Thus, finding a ternary heterogeneous azeotrope in the 1-propanol + water + propyl acetate system is expected. Data correlation. Generally, four ways to correlate the equilibrium data obtained from VLLE measurements can be found in the literature. The first method (here, called correlation VLLE) is to correlate directly the obtained VLLE data with a suitable algorithm such as the Britt-Luecke36 or Deming.37 This procedure has been used by us in a recent work22 with good results. The second method (here, called correlation VLE) consists in correlating VLE data extracted from the experimental VLLE data. In this way, VLE data are obtained after matching each liquid composition with its corresponding vapor composition (so the number of experimental points VLE will be twice the number of experimental points VLLE). The third method (called here correlation LLE) is similar to the above but using LLE data instead of VLE data. The LLE data are obtained by discarding the phase vapor mole fractions of the VLLE data. Finally, the last method (called here correlation VLE + LLE) involves a simultaneous VLE and LLE correlation. In this work, we assumed the nonideality of the liquid and vapor phase. In this way, the liquid−liquid equilibrium (LLE) condition is expressed as follows:34

Table 3. VLLE Data for 1-Propanol (1b) + Water (2) + Propyl Acetate (3) Ternary System at 101.3 kPaa organic rich phase, xIi xI1b

xI2

aqueous rich phase, xIIi xII1b

xII2

vapor phase, yi y1b

y2

T K

0.139 0.122 0.064 0.170 0.215 0.032 0.236 0.243 0.253 0.262 0.255 a

0.2698 0.2526 0.2084 0.2994 0.3665 0.1906 0.4042 0.4305 0.4589 0.5091 0.5924

0.013 0.010 0.007 0.016 0.021 0.002 0.027 0.030 0.032 0.035 0.042

0.984 0.987 0.990 0.981 0.975 0.994 0.970 0.967 0.964 0.961 0.952

0.090 0.081 0.051 0.108 0.132 0.024 0.147 0.161 0.174 0.183 0.207

0.521 0.525 0.528 0.524 0.524 0.534 0.526 0.525 0.529 0.533 0.539

355.37 355.39 355.44 355.50 355.53 355.63 355.69 355.83 356.08 356.14 356.74

u(T) = 0.02 K, u(x) = 0.002, and u(y) = 0.002.

These figures are the projection on the ternary composition diagram of the intersection of the single liquid−liquid envelope with the VLE surface.34 According to these representations, propyl acetate and water are partially miscible. On the other hand, water and propyl acetate are completely miscible with either of the two alcohols. Consequently, the LLE of the two systems exhibits a type I behavior.35 It is possible to know by the ternary composition diagram if a ternary system has a heterogeneous ternary azeotrope. If some vapor phase points are above and some others are below its tie line, there is a ternary heterogeneous azeotrope. In Figure 1, all the experimental vapor phase points are above its tie lines, so the ethanol + water + propyl acetate system has not a ternary heterogeneous azeotrope. In Figure 2, the vapor points are below its tie lines at very low concentrations of 1-propanol in the vapor phase. This condition is reversed when the 1-propanol

(γixi)I = (γixi)II

(1)

where I and II represent the liquid phases, γi the activity coefficient and xi the liquid phase mole fraction of the component i. On the other hand, the vapor−liquid equilibrium (VLE) condition can be expressed as yi =

γiPio ΦiP

xiI =

γiPio ΦiP

xiII

(2)

Poi

where is the saturated vapor pressure of the component i, calculated by the equation and parameters reported in DIPPR

Figure 1. Experimental VLLE diagram for the ethanol (1a) + water (2) + propyl acetate (3) ternary system at 101.3 kPa: ●, liquid phases; vapor phase; ■, binary azeotrope; ★, homogeneous ternary azeotrope; - - - - -, tie lines. C

△,



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Figure 2. Experimental VLLE diagram for the 1-propanol (1b) + water (2) + propyl acetate (3) ternary system at 101.3 kPa: ●, liquid phases (data from this work); △, vapor phase (data from this work); × , liquid phases (data from Smirnova et al.25); ▲, vapor phase (data from Smirnova et al.25); ■, binary azeotrope; ★, heterogeneous ternary azeotrope; - - - - -, tie lines.

Table 4. Vapor Pressure Parametersa Aib ethanol (1a) 1-propanol (1b) water (2) propyl acetate (3)

74.475 79.463 76.649 98.623

Bib −7164.3 −8294.9 −7258.2 −8038.4

Cib −7.3270 −8.9096 −7.3037 −12.4522

Dib 3.1340 1.8197 4.1653 8.8612

× × × ×

The standard state for the calculation of activity coefficients is the pure component at the pressure and temperature of the solution. Equation 5 is valid from low to moderate pressures when the virial equation of state truncated after the second coefficient is adequate to describe the vapor phase of the pure components and their mixtures, and pure components in state liquid are considered incompressible over the pressure range under consideration. The molar virial coefficients Bii and Bij were estimated by the method of Hayden and O’Connell40 using the molecular parameters suggested by Prausnitz et al.41 According to results reported in the literature,42−44 none of parameters obtained by correlation VLE seems to give satisfactory predictions of the liquid boiling envelope in most systems of type I. Other results45 confirm that parameters obtained by correlation LLE cannot predict correctly the vapor phase composition. Prausnitz et al.41 studied this topic and concluded that the best way to obtain parameters to estimate correctly the VLLE is with a correlation VLE + LLE. Anyway, a comparison between all methods has been done in order to determine the best method. In a first step, VLLE data have been treated by each one of the four methods and various sets of NRTL and UNIQUAC models parameters have been obtained. Values of obtained parameters are listed in Table 5 for the ethanol + water + propyl acetate system and Table 6 for the 1-propanol + water + propyl acetate system. Later, these parameters have been used to estimate VLLE data. Finally, deviations between experimental and estimated values of composition and temperature have been calculated. These deviations are shown in Tables 7 and 8 for the ethanol + water + propyl acetate system and the 1-propanol + water + propyl acetate system, respectively. Generally, all correlations estimate correctly the temperatures. Only deviation values obtained for the ethanol system

Eib −6

10 10−6 10−6 10−6

2 2 2 2

a

Vapor pressure equation: ln P°(Pa) = A + B/T(K) + C ln T(K) + D (T(K))E bParameters and vapor pressure equation obtained from DIPPR tables.35

tables38 and listed in Table 4, P is the equilibrium pressure, yi the molar vapor fraction and Φi is Φi =

0⎞ ⎛ L P − Pi ⎜ ⎟ − exp V i ϕi s ⎝ RT ⎠

ϕi

(3)

where ϕi is the fugacity coefficient of component i in the vapor phase, ϕis is the fugacity coefficient of a pure saturated liquid i, and ViL is the molar liquid volume of component i calculated by the Rackett equation.39 When the calculation of the activity coefficient was required, the following equation was used: ln γi = ln +

yP i xiPi0 P 2RT

+

(Bii − ViL)(P − Pi0) RT

∑ ∑ yyi k (2δji − δjk)

(4)

where T and P are the temperature and pressure in the equilibrium, respectively, Bii and Bjj are the second virial coefficients of the pure gases, Bij the cross second virial coefficient, and

δij = 2Bij − Bjj − Bii

(5) D



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Table 5. Parameters for Various GE Models Obtained from Various Types of Correlation for the Ternary System Ethanol (1a) + Water (2) + Propyl Acetate (3) UNIQUACa model

NRTL model correlation type

component i + j

VLLE

1a + 2 1a + 3 2+3 1a + 2 1a + 3 2+3 1a + 2 1a + 3 2+3 1a + 2 1a + 3 2+3

VLE

LLE

VLE + LLE

a

Δgij

Δgji

(J/mol)

(J/mol)

−2323.94 3016.88 12911.97 46.33 2603.73 12787.46 1564.42 −786.48 14202.32 1117.03 −668.10 14139.98

8850.87 1249.07 5971.46 5031.38 −284.98 4688.95 4806.23 5160.44 5541.69 4930.97 4679.94 5572.47

αij = αji 0.30 0.30 0.34 0.30 0.30 0.34 0.30 0.30 0.34 0.30 0.30 0.34

Δuij

Δuji

(J/mol)

(J/mol)

−2470.38 −1866.10 1166.66 111753.51 −1562.05 1110.35 1042.25 −2856.67 1170.56 742.26 −2673.38 1209.95

6267.48 5742.39 3758.06 −1834.02 3558.82 3489.15 −826.96 6508.07 3991.00 −216.71 6268.35 3684.64

Volume and surface parameters from DECHEMA.46

Table 6. Parameters for Various GE Models Obtained from Various Types of Correlation for the Ternary System 1-Propanol (1b) + Water (2) + Propyl Acetate (3)a UNIQUACa model

NRTL model correlation type

component i + j

Δgij (J/mol)

(J/mol)

1b + 2 1b + 3 2+3 1b + 2 1b + 3 2+3 1b + 2 1b + 3 2+3 1b + 2 1b + 3 2+3

−2214.26 5069.17 14496.27 −799.15 2175.38 14571.07 −2149.05 4166.37 14542.98 −2213.08 5135.23 14488.98

11004.15 −680.43 5602.94 9172.94 145.68 4958.75 10682.84 −574.23 5614.91 11021.12 −690.80 5592.32

VLLE

VLE

LLE

VLE + LLE

a

Δgji

αij = αji

Δuij

Δuji

(J/mol)

(J/mol)

0.30 0.30 0.34 0.30 0.30 0.34 0.30 0.30 0.34 0.30 0.30 0.34

−1884.65 −941.67 1944.66 350.93 4867.55 1952.04 −1944.92 −1275.95 1942.09 −1872.19 −973.83 1952.48

4608.76 2047.21 2710.70 1498.39 −1898.25 2450.43 4645.66 2471.08 2711.05 4564.43 2051.19 2673.57

Volume and surface parameters from DECHEMA.46

Table 7. Correlation Statistics for the NRTL and UNIQUAC Models Obtained with VLLE Data from the System Ethanol (1a) + Water (2) + Propyl Acetate (3) correlation type

AAD Ta

AAD xI1ab

AAD xI2b

AAD xII1ab

AAD xII2 b

AAD y1ac

AAD y2c

∑AADd

0.0033 0.0093 0.0033 0.0030

0.0179 0.0155 0.0374 0.0257

0.0135 0.0093 0.0396 0.0329

0.0728 0.1413 0.1200 0.1006

0.0044 0.0064 0.0049 0.0048

0.0440 0.0223 0.0916 0.0661

0.0134 0.0074 0.0462 0.0335

0.1003 0.1173 0.1924 0.1608

K NRTL Model VLLE VLE LLE VLE + LLE

0.20 0.18 0.35 0.31

0.0112 0.0224 0.0105 0.0103

0.0228 0.0760 0.0241 0.0240

VLLE VLE LLE VLE + LLE

0.09 0.06 0.41 0.31

0.0123 0.0200 0.0120 0.0151

0.0208 0.0536 0.0321 0.0355

0.0041 0.0088 0.0051 0.0047 UNIQUAC Model 0.0054 0.0076 0.0056 0.0058

a

Average absolute deviation in temperature. bAverage absolute deviation in liquid phase composition. cAverage absolute deviation in vapor phase composition. dSum of the average absolute deviations in temperature and liquid and vapor phase composition, ∑AAD = AAD xI1a + AAD xI2 + AAD xII1a + AAD xII2 + AAD y1 + AAD y2.

vapor phase mole fractions and the binodal curve, respectively. However, and according to Tables 7 and 8, LLE correlations do not have the lowest values for deviations of the liquid phase mole fractions. Thus, the correlation LLE does not ensures the

with correlations LLE are above 0.35 K. Anyway, these deviations are still acceptable values. Regarding the estimation of concentrations, it can be expected that correlation VLE and correlation LLE will have the best estimations of the E



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Table 8. Correlation Statistics for the NRTL and UNIQUAC Models Obtained with VLLE Data from the System 1-Propanol (1b) + Water (2) + Propyl Acetate (3) correlation type

AAD Ta

AAD xI1bb

AAD xI2b

AAD xII1bb

AAD xII2 b

AAD y1c

AAD y2c

∑AADd

0.0006 0.0010 0.0007 0.0006

0.0099 0.0087 0.0144 0.0094

0.0172 0.0161 0.0131 0.0174

0.0586 0.0813 0.0600 0.0582

0.0007 0.0020 0.0009 0.0008

0.0238 0.0104 0.0289 0.0243

0.0062 0.0059 0.0087 0.0068

0.0429 0.0690 0.0509 0.0450

K NRTL Model VLLE VLE LLE VLE + LLE

0.19 0.20 0.18 0.19

0.0093 0.0198 0.0091 0.0093

0.0208 0.0345 0.0218 0.0207


0.12 0.12 0.14 0.12

0.0052 0.0106 0.0047 0.0053

0.0061 0.0384 0.0066 0.0068

0.0008 0.0012 0.0009 0.0008 UNIQUAC Model 0.0009 0.0017 0.0011 0.0010

a

Average absolute deviation in temperature. bAverage absolute deviation in liquid phase composition. cAverage absolute deviation in vapor phase composition. dSum of the average absolute deviations in temperature and liquid and vapor phase composition, ∑AAD = AAD xI1b + AAD xI2 + AAD xII1b + AAD xII2 + AAD y1 + AAD y2.

Table 9. Experimental and Calculated Binary and Ternary Azeotropic Data for the System Ethanol (1a) + Water (2) + Propyl Acetate (3) at 101.3 kPa azeotrope (1a + 2) correlation type

x1a

(1a + 3) T

x1a

K

VLLE VLE LLE VLE + LLE VLLE VLE LLE VLE + LLE a

0.893a

351.33a

0.795 0.839 0.726 0.743

349.82 351.07 348.86 349.52

0.795

345.55

(2 + 3) T

x2

K 0.922b

0.991

(1a + 2 + 3) T

x1a

x2

K

Experimental data 351.02b 0.520c NRTL model 0.502 351.54 0.513 0.507 0.507 UNIQUAC model 0.514 0.520 0.511 0.515

T K

355.55c

0.607d

0.227d

350.82d

355.05 355.56 355.32 355.32

0.645 0.571 0.571

0.238 0.291 0.284

350.72 348.67 349.25

0.573

0.171

352.81

0.401 0.393

0.401 0.398

357.88 356.24

355.66 355.87 355.45 355.72

Taken from Kurihara et al.3 bTaken from Ortega et al.47 cTaken from Smirnova et al.48 dFrom this work.

data were obtained from the literature.3,47−49 On the other hand, data corresponding to ternary azeotropes were measured experimentally in this work. To determine temperature and concentrations of the ternary azeotropes, a fully automated unit Fischer Labodest HMS 500 AC was used. This appliance is equipped with a Fischer SPALTROHR-Column, which has a separation efficiency of up to 90 theoretical plates. To reach the azeotrope conditions, a sample of the studied mixture (water, propyl acetate and ethanol or 1-propanol) was introduced into the reboiler of the column. First, the column worked under total reflux conditions, and after some time, the reflux ratio was changed to a high value (>50). Finally, a sample with the azeotropic composition was obtained as column top product. Figures 1 and 2 show the azeotropes present in each system. According to the figures, the ternary azeotrope of the ethanol + water + propyl acetate system is homogeneous whereas the azeotrope of the 1-propanol + water + propyl acetate system is heterogeneous. The values of concentrations and temperature of these azeotropes are listed in Tables 9 and 10 for the ternary systems ethanol + water + propyl acetate and

best prediction of the binodal curve. On the other hand, the parameters obtained with a correlation VLE have the best predictions of the vapor phase mole fractions. The last column in Tables 7 and 8 lists the values for the sum of all composition deviations. For the ethanol + water + propyl acetate system the lowest value corresponds to the parameters calculated from a correlation VLLE for the NRTL and UNIQUAC models. However, in the latter case, the parameters of the correlation VLE also have an acceptable value. These results are quite different from those obtained in the 1-propanol system. Here, correlation VLE has the highest value for the sum of all composition deviations in both models, whereas correlation VLLE has low values. In conclusion, the best way to estimate the VLLE data is with parameters models obtained from correlations VLLE. In addition to deviations between experimental and calculated values, the ability of the thermodynamic model to predict azeotropes is also very important, especially if the equilibrium data are used to design an azeotropic distillation process. In this case, the two studied systems have three binary azeotropes and one ternary azeotrope. Binary azeotropes F



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Table 10. Experimental and Calculated Binary and Ternary Azeotropic Data for the System 1-Propanol (1b) + Water (2) + Propyl Acetate (3) azeotrope (1b + 2) correlation type

x1b

(1b + 3) T

x1b

K

a

(2 + 3) T

x2

K

0.432a

360.10a

0.613b


0.334 0.397 0.342 0.334

364.13 361.47 363.98 363.99

0.615 0.602 0.612 0.616


0.359 0.391 0.359 0.360

364.72 362.03 364.97 364.62

0.562 0.688 0.556 0.567

(1b + 2 + 3) T

x1a

x2

K

Experimental data 367.63b 0.520c NRTL Model 364.37 0.507 367.11 0.514 365.41 0.507 364.30 0.507 UNIQUAC Model 367.54 0.525 365.10 0.530 368.47 0.525 367.82 0.525

T K

355.55c

0.081d

0.525d

355.36d

355.26 355.63 355.34 355.35

0.032 0.067 0.017 0.033

0.505 0.508 0.506 0.505

355.27 355.52 355.28 355.21

0.120

0.516

356.02

356.13 356.44 356.15 356.16

Taken from Demirel et al.49 bTaken from Ortega et al.47 cTaken from Smirnova et al.48 dFrom this work.

Figure 3. Comparison between experimental and calculated VLLE data for the ethanol (1a) + water (2) + propyl acetate (3) ternary system at 101.3 kPa. Experimental data: ●, liquid phases; △, vapor phase. Calculated data with parameters listed in Table 5: NRTL model using parameters obtained by correlation VLE; - - - -, UNIQUAC model using parameters obtained by correlation VLE.

1-propanol + water + propyl acetate, respectively. In addition to the experimental data, Tables 9 and 10 also report calculated data. According to calculated azeotropic data from Table 9, all cases make a good prediction of the heterogeneous binary azeotrope involving water and propyl acetate. On the other hand, the ethanol−propyl acetate azeotrope is predicted only with one combination (correlation VLE and NRTL model). The prediction of ethanol−water azeotrope largely depends on the model used. Thus, all estimations made with the NRTL model can predict the azeotrope in contrast to the case of the UNIQUAC model, where the azeotrope is only predicted with parameters obtained by a correlation VLE. Finally, the best estimation of the ternary azeotrope is clearly made with

the NRTL model with parameters obtained by correlation VLE. In summary, for the system ethanol + water + propyl acetate and considering azeotropic predictions the best model is NRTL with parameters from a correlation VLE. Table 10 shows the estimated azeotropic data corresponding to the 1-propanol + water + propyl acetate system. In contrast to the previous studied system, parameters obtained by all correlations predict the presence of the three binary azeotropes. The best model to predict the binary water + propyl acetate azeotrope is the UNIQUAC model. On the other hand, the NRTL model is better to predict the 1-propanol + propyl acetate azeotrope. In the case of the 1-propanol + water azeotrope, the best predictions are reached with parameters obtained in a G



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Figure 4. Comparison between experimental and calculated VLLE data for the 1-propanol (1b) + water (2) + propyl acetate (3) ternary system at 101.3 kPa. Experimental data: ●, liquid phases; △, vapor phase. Calculated data with parameters listed in Table 6: , NRTL model using parameters obtained by correlation VLE; - - - -, UNIQUAC model using parameters obtained by correlation VLE.

Figure 5. Residue curve map for the system ethanol (1a) + water (2) + propyl acetate (3) at 101.3 kPa. Binodal curve and residue curves have been calculated using the NRTL model with the parameters obtained in correlation VLE and given in Table 5.

the accuracy in azeotropic estimations is crucial to design properly the azeotropic distillation process so this is considered the determinant factor. Therefore, model parameters used in the VLLE data estimation must be taken from a VLE correlation in order to ensure a correct azeotropic prediction. Estimated VLLE data with the NRTL and UNIQUAC models are plotted in Figures 3 and 4 for the ethanol + water + propyl acetate ternary system and the 1-propanol + water + propyl

correlation VLE, especially if the used model is UNIQUAC. Predictions of the ternary heterogeneous azeotrope are closer to the experimental values if the parameters used in the VLLE estimation are obtained from a correlation VLE. In conclusion, the best azeotropic estimations for this system are achieved with parameters obtained by a VLE correlation. As seen so far, the best type of correlation to obtain the model parameters depends on the considered factor. However, H



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Figure 6. Residue curve map for the system 1-propanol (1b) + water (2) + propyl acetate (3) at 101.3 kPa. Binodal curve and residue curves have been calculated using the UNIQUAC model with the parameters obtained in correlation VLE and given in Table 6.

Figure 7. Proposed three-column distillation sequence for the separation of water and 1-propanol mixtures by an azeotropic distillation process.

correlation VLE, as it is the combination with the best azeotropic predictions. On the other hand, UNIQUAC model with parameters calculated by the correlation VLE has been used to build the RCM that corresponds to the 1-propanol system. As expected, both figures have similar characteristics. There are three distillation regions where pure components are the stable node and the ternary azeotrope is the unstable node. The binary azeotropes are the saddle points which delimit the three zones. In both cases, the alcohol (ethanol or 1-propanol) can be obtained as the bottom stream if the mixture to be separated is located in the region II. In this case, the top product is the

acetate ternary system, respectively. As can be seen, deviations between calculated and experimental values of mole fractions are high in all phases, especially at high concentrations of 1-propanol. Residue Curve Map. Residue curve maps (RCMs) are important tools for analyzing and planning a distillation sequence. Figures 5 and 6 show the RCM corresponding to the ethanol + water + propyl acetate ternary system and the 1-propanol + water + propyl acetate ternary system, respectively. The model used to build the RCM of the ethanol system has been the NRTL model with parameters calculated by the I



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ternary azeotrope. The main difference between the two systems is the type of azeotrope. One ternary azeotrope is homogeneous (ethanol + water + propyl acetate) and the other is heterogeneous (1-propanol + water + propyl acetate). As the heterogeneous azeotropic distillation is based on the presence of a heterogeneous ternary azeotrope; propyl acetate is not a suitable solvent for the ethanol + water system. In this way, henceforth only the case studio of the ternary system 1-propanol + water + propyl acetate will be analyzed. Besides, the possibility of using propyl acetate as a separation agent for an ethanol dehydration process should not be discarded, despite the fact that the region II of ethanol RCM is very small (see Figure 5). A three-column distillation sequence to separate 1-propanol and water by an azeotropic distillation process can be proposed based on information extracted from the RCM plot in Figure 6. Figure 7 shows the flow diagram of this process. In the first column (preconcentration column), a mixture of water and 1-propanol is distillated until near the binary azeotropic point. Then, propyl acetate is added to reach the region II in the RCM (see Figure 6). The minimum amount of propyl acetate that can be added to reach the region II is approximately 10 mol % of the initial amount of water and 1-propanol. On the other hand, the addition of propyl acetate at amounts above 28 mol % of the initial mixture results a stream located in region III. The resulting stream is the feed stream of the second column (azeotropic column). Here, the 1-propanol is obtained as the bottom product. The top product is the ternary heterogeneous azeotrope. The condensed top product goes to a decanter where the stream is split into the water-rich-phase stream and the organic-rich-phase stream. The resulting water-rich-phase stream is composed almost entirely of water and it must be treated depending on later use. On the other hand, the organic-richphase stream is the feed of the third column (entrainer recovery column). In this column, the propyl acetate is obtained as the bottom product. The top product will have the same composition as the top product of the azeotropic column. In this way, it is recycled to the decanter. However, this sequence is just a predesign based on the residue curve maps and, therefore, there are many factors that have not been considered. Moreover, other alternatives such as a process without the preconcentration column must be analyzed. Anyway, each design of an azeotropic distillation process requires a complete simulation of the process in order to asses their viability.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Tel.: +34 963544317. Fax: +34 963544898. Funding

Financial support from the Ministerio de Ciencia y Tecnologiá of Spain, through project No. CTQ2010-18848 is gratefully acknowledged. J. P.-F. is deeply grateful for the grant BES-201104636 received from the Ministerio de Economiá y Competitividad. Notes

The authors declare no competing financial interest.

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REFERENCES

(1) Speight, J. G. Lange’s Handbook of Chemistry, 16th Edition; McGraw-Hill: New York, 2005. (2) Rosseau, R. W. Handbook of Separation Process Technology; WileyIEEE: New York, 1987. (3) Kurihara, K.; Nakamichi, M.; Kojima, K. Isobaric vapor-liquid equilibria for methanol + ethanol + water and the three constituent binary systems. J. Chem. Eng. Data 1993, 38, 446−449. (4) Kanti, P.; Srigowri, J.; Madhuri, B.; Smitha, S.; Sridhar, S. Dehydration of ethanol through blend membranes of chitosan and sodium alginate by pervaporation. Sep. Purif. Technol. 2004, 40, 259− 266. (5) Ma, Y.; Wang, J. H.; Tsuru, T. Pervaporation of water/ethanol mixtures through microporous silica membranes. Sep. Purif. Technol. 2009, 66, 479−485. (6) Ling, W. S.; Thian, T. C.; Bhatia, S. Synthesis, characterization and pervaporation properties of microwave synthesized zeolite A membrane. Sep. Purif. Technol. 2010, 71, 192−199. (7) Lu, L.; Shao, Q.; Juang, L. Simulation of adsorption and separation of ethanol−water mixture with zeolite and carbon nanotube. Fluid Phase Equilib. 2007, 261, 191−198. (8) Hong, J.; Voloch, M.; Ladisch, M. R.; Tsao, G. T. Adsorption of ethanol−water mixtures by biomass materials. 3. Biotechnol. Bioeng. 1982, 24, 725−730. (9) Simo, M.; Bown, C. J.; Hlavacek, V. Simulation of pressure swing adsorption in fuel ethanol production process. Comput. Chem. Eng. 2008, 32, 1635−1649. (10) Oleffmanm, R. D.; Stephenson, S. K.; Franqui, P.; Cline, J. L. Extraction of ethanol with higher alcohol solvents and their toxicity to yeast. Purif. Technol. 2008, 63, 444−451. (11) Arifeen, N.; Wang, R.; Kookos, I. K.; Webb, C.; Koutinas, A. A. Process design and optimization of novel wheat-based continuous bioethanol production system. Biotechnol. Prog. 2007, 23, 1394−1403. (12) Kotai, B.; Lang, P.; Modla, G. Batch extractive distillation as a hybrid process: separation of minimum boiling azeotropes. Chem. Eng. Sci. 2007, 62, 6816−6826. (13) Seiler, M.; Kholer, D.; Arlt, W. From Alcohols to Hyperbranched Polymers: The Influence of Differently Branched Additives on the Vapor−Liquid Equilibria of Selected Azeotropic Systems. Purif. Technol. 2003, 30, 179−197. (14) Calvar, N.; Gonzalez, B.; Gomez, E.; Dominguez, A. Vapor− liquid equilibria for the ternary system ethanol + water + 1-ethyl-3methylimidazolium ethylsulfate and the corresponding binary systems containing the ionic liquid at 101.3 kPa. J. Chem. Eng. Data 2008, 53, 820−825. (15) Ligero, E. L.; Ravagnani, T. M. K. Dehydration of ethanol with salt extractive distillation -a comparative analysis between processes with salt recovery. Chem. Eng. Process. 2003, 42, 543−552. (16) Vasconcelos, C. J. G.; Wolf-Maciel, M. R. Dynamic and control of high purity heterogeneous azeotropic distillation process. Comput.Aided Chem. Eng. 2000, 8, 217−222. (17) Ravagnani, M. A. S. S.; Reis, M. H. M.; Maciel Filho, R.; WolfMaciel, R. R. Anhydrous ethanol production by extractive distillation: A solvent case study. 1. Process Saf. Environ. Prot. 2000, 88, 67−73.

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CONCLUSIONS Isobaric VLLE data were measured for the ternary systems ethanol + water + propyl acetate and 1-propanol + water + propyl acetate at 101.3 kPa. The data obtained in these experiments show the presence of a homogeneous ternary azeotrope in the ethanol + water + propyl acetate system. For this reason, use of the propyl acetate as a solvent in an azeotropic distillation process to separate water and ethanol is discarded. On the other hand, a heterogeneous ternary azeotrope appears in the 1-propanol + water + propyl acetate system. VLLE data estimation is carried out using the NRTL and UNIQUAC local composition models. Before that, the best way to obtain parameters of these models is previously discussed. According to the results, the only way to estimate all the azeotropes appearing in each one of the studied systems is with parameters obtained from a VLE correlation. Finally, a threecolumn distillation sequence to separate water and 1-propanol is proposed. In this process, 1-propanol is obtained as bottom product in the azeotropic column. J



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(18) Huang, H. J.; Ramaswamy, S.; Tschirner, U. W.; Ramarao, B. V. A review of separation technologies in current and future biorefineries. Sep. Purif. Technol. 2008, 62, 1−21. (19) Pla-Franco, J.; Lladosa, E.; Loras, S.; Montón, J. B. Phase equilibria for the ternary systems ethanol, water + ethylene glycol or + glycerol at 101.3 kPa. Fluid Phase Equilib. 2013, 341, 54−60. (20) Pla-Franco, J.; Lladosa, E.; Loras, S.; Montón, J. B. Corregendum to “Phase equilibria for the ternary systems ethanol, water + ethylene glycol or + glycerol at 101.3 kPa. Fluid Phase Equilib. 2013, 356, 374−374. (21) Pla-Franco, J.; Lladosa, E.; Loras, S.; Montón, J. B. Evaluation of 2-methoxyethanol as entrainer in ethanol-water and 1-propanol-water mixtures. J. Chem. Eng. Data 2013, 58, 3504−3512. (22) Pla-Franco, J.; Lladosa, E.; Loras, S.; Montón, J. B. Thermodynamic analysis and process simulation of ethanol dehydration via heterogeneous azeotropic distillation. Ind. Eng. Chem. Res. 2014, 53 (14), 6084−6093. (23) Smirnova, I. A.; Marachevskii, A. G.; Storonkin, A. B. Vestn. Leningr. Univ., Ser. 4. Fiz. Khim. 1963, 18 (22), 97−97. (24) Cehreli, S.; Ozmen, D.; Dramur, U. (Liquid + liquid) equilibria of (water + 1-propanol + solvent) at T = 298.2 K. Fluid Phase Equilib. 2006, 239, 156−160. (25) Smirnova, N. A.; Morachevskii, A. G.; Storonkin, A. V. Liquidvapor and liquid-liquid-vapor equilibrium in the propyl alcohol propyl acetate − water system: I isobaric conditions. Fiz. Khim. 1959, 14, 70−80. (26) Renon, H.; Prausnitz, J. M. Local compositions in thermodynamic excess functions for liquid mixtures. AIChE J. 1968, 14, 135−144. (27) 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, 116−128. (28) Arce, A.; Martínez-Ageitos, J.; Rodil, E.; Soto, A. Thermodynamic behavior of ethanolqmethanolq2-ethoxy-2-methylpropane system. Physical properties and phase equilibria. Fluid Phase Equilib. 1999, 165, 121−139. (29) Gascón, I.; Martín, S.; Artigas, H.; López, M. C.; Lafuente, C. Isobaric vapour−liquid equilibrium of binary and ternary mixtures containing cyclohexane, n-hexane, 1,3-dioxolane and 1-butanol at 40.0 and 101.3 kPa. Chem. Eng. J. 2002, 88, 1−9. (30) Loras, S.; Aucejo, A.; Muñoz, R.; de la Torre, J. Isobaric vapor− liquid equilibrium for binary and ternary mixtures of 2-methyl-2propanol C methyl 1,1-dimethylpropyl ether C 2,2,4-trimethylpentane. Fluid Phase Equilib. 2000, 175, 125−138. (31) Loras, S.; Fernández-Torres, M. J.; Gomis-Yagües, V.; RuizBeviá, F. Isobaric vapor−liquid equilibria for the system 1-pentanol−1propanol−water at 101.3 kPa. Fluid Phase Equilib. 2001, 180, 205− 210. (32) Gomis, V.; Font, A.; Saquete, M. D.; García-Cano, J. LLE, VLE and VLLE data for the water−n-butanol−n-hexane system at atmospheric pressure. Fluid Phase Equilib. 2012, 316, 135−140. (33) Gomis, V. R.; Ruiz, F.; Asensi, J. C. The application of ultrasound in the determination of isobaric vapour-liquid-liquid equilibria data. Fluid Phase Equilib. 2000, 172, 245−259. (34) Walas, S. M. Phase Equilibria in Chemical Engineering; Butterworth-Heinemann: Boston, 1984. (35) Sørensen, J. M.; Magnussen, T.; Rasmussen, P.; Fredenslund, A. Liquid-liquid equilibrium data: Their retrieval, correlation and prediction Part I: Retrieval. Fluid Phase Equilib. 1979, 2, 297−309. (36) Britt, H. I.; Luecke, R. H. The Estimation of parameters in nonlinear, implicit models. Technometrics 1973, 15, 233−247. (37) Deming, W. E.; Stephan, F. F. On a least squares adjustment of a sampled frequency table when the expected marginal totals are known. Ann. Math. Stat. 1940, 11, 367−484. (38) Daubert, T. E.; Danner, R. P. Physical and Thermodynamic Properties of Pure Chemicals: Data Compilation; Taylor & Francis: Bristol, PA, 1989. (39) Rackett, H. G. Equation of State for Saturated Liquids. J. Chem. Eng. Data 1970, 15, 514−517.

(40) Hayden, J.; O’Connell, A. A generalized method for predicting second virial coefficients. Ind. Eng. Chem. Process Des. Dev. 1975, 14, 209−216. (41) Prausnitz, J.; Anderson, T.; Gren, E.; Eckert, C.; Hsieh, R.; O’Connell, J. Computer Calculation for Multicomponent Vapor-Liquid and Liquid-Liquid Equilibria; Prentice Hall: Englewood Cliffs, NJ, 1980. (42) Sørensen, J. M.; Magnussen, T.; Rasmussen, P.; Fredenslund, A. Liquid-liquid equilibrium data: Their retrieval, correlation and prediction Part II: Correlation. Fluid Phase Equilib. 1979, 2, 297−309. (43) van Zandijjcke, F.; Verhoeye, L. The vapour-liquid equilibrium of ternary systems with limited miscibilities at atmospheric pressure. J. Appl. Chem. Biotechnol. 1974, 25, 709−729. (44) Cha, T. H.; Prausnitz, J. M. Thermodynamic method for simultaneous representation of ternary vapor-liquid and liquid-liquid equilibria. Ind. Eng. Chem. Process Des. Dev. 1985, 24, 551−555. (45) Sugi, H.; Katayama, T. Ternary liquid-liquid and miscible binary vapor-liquid equilibrium data for the two systems n-hexane-ethanolacetonitrile and water-acetonitrile-ethyl acetate. J. Chem. Eng. Jpn. 1978, 11, 167−172. (46) Gmehling, J.; Onken, U. Vapor−Liquid Equilibrium Data Collection; DECHEMA: Frankfurt, 1977. (47) Ortega, J.; González, C.; Peña, J.; Galván, S. Thermodynamic study on binary mixtures of propyl ethanoate and and alkan-1-ol (C2C4). Isobaric vapor−liquid equilibria and excess properties. Fluid Phase Equilib. 2000, 170, 87−111. (48) Smirnova, N. A.; Morachevski, A. G. Liquid-vapor equilibrium and mutual solubility of components in the propyl acetate-water system. Zh. Fiz. Khim. (J. Phys. Chem. USSR) 1960, 34, 2546−2553. (49) Demirel, Y. Estimation of the entropy of vaporization at the normal boiling point for azeotropic mixtures containing water, alcohol or acetic acid. Thermochim. Acta 1999, 339, 79−85.

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