Proposal of Isobutyl Alcohol as Entrainer To Separate Mixtures

Article pubs.acs.org/jced

Proposal of Isobutyl Alcohol as Entrainer To Separate Mixtures Formed by Ethanol and Water and 1‑Propanol and Water Jordi Pla-Franco, Estela Lladosa, Sonia Loras,* and Juan B. Montón Departamento de Ingeniería Química, Escuela Técnica Superior de Ingeniería, Universitat de València, 46100 Burjassot, Valencia, Spain ABSTRACT: Isobutyl alcohol (IBA) has been proposed as a solvent to carry out the dehydration of ethanol and 1-propanol by means of a nonconventional distillation process. In this way, isobaric vapor−liquid equilibrium (VLE) and vapor−liquid−liquid equilibrium (VLLE) data at atmospheric pressure have been obtained for the ternary systems ethanol (1a) + water (2) + IBA (3) and 1-propanol (1b) + water (2) + IBA (3). Then, these data have been correlated to obtain a set of parameters capable of estimating VLE and VLLE. According to the results achieved, distillation sequences to separate water and ethanol or 1-propanol have been proposed. Finally, a study about the minimum amount of entrainer required to break the azeotrope of the mixture ethanol/1-propanol + water is carried out.

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INTRODUCTION During the last decades, and due to the problem of global warming, some alternatives to fossil fuels have been proposed in order to reduce the total amount of CO2 emissions into the atmosphere.1 Of all these new proposals, the use of biofuels such as bioethanol or biopropanol is very interesting since both compounds are obtained from renewable energy sources.2−4 However, their production presents some difficulties. For example, conventional distillation cannot be used to completely dehydrate bioethanol and biopropanol because of the binary azeotrope appearing in the systems formed by water and either of the two alcohols. To solve this problem, membrane technology5 or nonconventional distillation processes have been proposed, including extractive distillation (ED)6 and azeotropic distillation (AD).7 These distillations are based on the addition of a third compound called entrainer that alters the relative volatility of the components of the original mixture. Its main advantage is the possibility of carrying out dehydration of large amounts of ethanol or 1-propanol, unlike the membranes processes.8 Furthermore, although the total energy consumption in these distillations is high, it can be reduced by choosing the proper entrainer and operating conditions. Therefore, both ED and AD are recommended techniques for dehydration of alcohols. In an ED process, the entrainer does not form an azeotrope with any of the components of the mixture to be separated. On the other hand, in an AD process, addition of the third compound forms an additional azeotrope to carry out the desired separation. Depending on the number of phases present in the new azeotrope, the AD will be homogeneous (a unique liquid phase) or heterogeneous (two liquid phases). The heterogeneous AD is more interesting because the presence of the two liquid phases allows separation with a decanter. Anyway, in both cases it is essential to know the vapor−liquid equilibrium (VLE) © 2017 American Chemical Society

data to properly design a distillation sequence. In addition, in cases with more than one liquid phase, the knowledge of the vapor−liquid−liquid equilibrium (VLLE) data is also required. In previous works,9−13 several organic compounds were proposed as the entrainer in an ED (glycerol, ethylene glycol, 2-methoxyethanol) or AD (diisopropyl ether, isobutyl alcohol, propyl acetate) process for removing water from ethanol or 1-propanol. Results showed that isobutyl alcohol (IBA) does not form a ternary azeotrope with ethanol and water, only a binary azeotrope with water. In addition, its heterogeneous region comprises a very limited composition range. For this reason, IBA was discarded as entrainer in an AD process to separate water an ethanol. However, a process combining AD and ED can be proposed. Because of this, isobaric VLE data at 101.3 kPa have been obtained for the ternary systems ethanol (1a) + water (2) + IBA (3) and 1-propanol (1b) + water (2) + IBA (3). VLLE data at 101.3 kPa have been determined only for the 1-propanol system since data for ethanol mixtures were obtained in the previously mentioned work.12 VLE data at several temperatures for the system ethanol + water + IBA were obtained by Perelygin et al.14 and isobaric VLE at 101.3 kPa were also reported previously.15−17 No VLE data for the system 1-propanol + water + IBA have been found in the literature. The choice of IBA is in concordance with the principles of Green Chemistry, as IBA has a low toxicity (LD50 in rats equal to 2460 mg/kg) and can be obtained from biomass. Once VLE and VLLE experimental data are obtained, it is possible to correlate them in order to obtain parameters of the Special Issue: Memorial Issue in Honor of Ken Marsh Received: January 30, 2017 Accepted: May 11, 2017 Published: May 24, 2017 2697

DOI: 10.1021/acs.jced.7b00098 J. Chem. Eng. Data 2017, 62, 2697−2707

Journal of Chemical & Engineering Data

Article

NRTL18 and UNIQUAC19 local composition models. These parameters have been subsequently used to estimate the equilibrium and calculate the residue curve maps (RCMs) to design the distillation sequences of the separation processes. The last part of the work is a preliminary calculation of the total amount of IBA required to carry out the separation.

VLLE data. In this case, sampling of the vapor phase cannot be performed directly with a syringe, and instead a 6-port heated valve from Valco Instruments Co. injects the sample directly into the chromatograph. The tube through which the vapor phase sample flows is heated with a resistance whose temperature is controlled by a digital potentiometer. Value of this temperature is kept along all the tube over 20 °C higher than the dew temperature of the vapor sample to ensure that condensation does not occur. The sampling of the liquid phase is totally different and takes place once two consecutives vapor phase samples have the same composition. The heterogeneous liquid sample is taken from the separation chamber of the equipment. Then, the sample is placed on a heat-insulated test tube. Immediately, this tube is put in a Block Heater SBH200DC from Stuart Scientific to keep the temperature at a slightly lower value than the mixture bubble point. The equilibrium is reached between the two liquid phases after a time, usually 1 day. At this moment, one sample of each phase is taken and located in a vial. A small amount of 1-propanol is added to samples with ethanol to avoid the phase separation effects caused by the change in temperature. For this same reason, ethanol is added to 1-propanol samples. Analysis. The concentration of the samples was determined by gas chromatography (GC). The composition of all samples was obtained by means of a CE Instruments GC 8000 Top GC which uses a thermal conductivity detector (TCD) together with 80/100 Porapak Q 3 m × 1/8 in. In both systems the optimum operations conditions were as follows: injection temperature, 473 K; oven temperature, 373 K; detector temperature, 433 K; detector current, 220 mA; and helium flow rate, 40 mL/min. Prior to sample analysis, the chromatograph was calibrated gravimetrically with prepared standard solutions. Thus, peak area ratios could be converted to molar fractions with a standard deviation generally lower than 0.002.

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EXPERIMENTAL SECTION Chemicals. Table 1 lists the supplier and purity of all chemicals used in this work. In this sense, water (>0.990 mass fraction,

Table 1. Specifications of Chemical Samples chemical name

source

fraction purity

purification method

analysis method

ethanol 1-propanol water isobutyl alcohol

Aldrich Aldrich Merck Aldrich

0.999 0.990 0.990 0.990

none none none none

GCa GCa GCa GCa

a

Gas chromatography.

water for chromatography LiChrosolv) was supplied by Merck while ethanol (>0.999 mass fraction), 1-propanol (>0.990 mass fraction), and IBA (>0.990 mass fraction) were supplied by Sigma-Aldrich. No purification technique was required as no significant impurities were found after the chromatographic analysis. Water content was below 0.001 mass fraction in all organic compounds; even so, organic chemicals were dried over molecular sieves (Aldrich, type 4 Å, 1.6 mm pellets). Apparatus and Procedure. Both VLE and VLLE were measured by means of an all-glass dynamic-recirculating still (Labodest VLE 602/D) manufactured by Fischer Labor and Verfahrenstechnik (Germany), which is widely referenced in the literature.20,21 The mixture temperature is measured using a digital Hart Scientific thermometer model 1502A with a standard uncertainty equal to 0.02 K. This thermometer has a Pt 100 probe Hart Scientific model 5622 calibrated at the ENACaccredited Spanish Instituto Nacional de Técnica Aeroespacial using the ice and steam points of distilled water as references. The pressure and the heating power are measured and controlled via a Fisher M101 pressure control system, with an estimated value of the standard uncertainty equal to 0.1 kPa. The manometer is calibrated using the vapor pressure of water. In this kind of apparatus, the equilibrium is reached in its Cottrell circulation pump. To homogenize the boiling mixture to ensure the correct VLLE, the still has an ultrasonic homogenizer (Labsonic P with a titanium probe of 14 mm i.d. from Sartorius Stedim Biotech) coupled to its boiling flask, similar to that introduced by Gomis et al.22 Regardless of the type of data (VLE or VLLE), all experiments were carried out at a constant pressure of 101.3 kPa. The pressure controller of the equipment and a vacuum pump connected to the equilibrium apparatus were used to mantain a constant pressure. Experiments begin when the heating system is turned on. In addition, the liquid stirring system and the ultrasonic homogenizer are also activated. After a while, temperature and pressure values do not change significantly. Then sampling can begin, which varies depending on the number of liquid phases. For VLE data, sampling is done directly from the equipment using special syringes (0.5 μL). When two consecutive samples have the same composition (a difference in the mole fraction smaller than 0.001) in each of the phases, it is considered that equilibrium has been reached. This procedure is different for

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RESULTS AND DISCUSSION Experimental Data. Equilibrium temperature T, the liquid phase mole fraction of component i, xi, and the vapor phase mole fraction of component i, yi, were obtained for the ternary systems formed after adding IBA to each one of the mixtures to separate (ethanol + water and 1-propanol + water). Table 2 lists the experimental values of these variables for the system with ethanol while Tables 3 (VLE) and 4 (VLLE) do the same for the system with 1-propanol. In addition, the activity coefficient of component i, γi, has also been included in Tables 2 and 3. It was calculated assuming the nonideality of both phases. In this way, the liquid−liquid equilibrium (LLE) condition is expressed as follows: (γixi)I = (γixi)II

(1)

where the superscripts I and II represent the liquid phases. On the other hand, the vapor−liquid equilibrium (VLE) condition can be expressed as yi =

γiPio P Φi

xiI =

γiPio P Φi

xiII

(2)

where Poi is the vapor pressure of the component i, calculated by the equation and parameters reported in DIPPR tables19 and listed in Table 5, P is the equilibrium pressure, and Φi is the fugacity coefficients quotient multiplied by the Poynting 2698



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Table 2. VLE Data for Ethanol (1a) + Water (2) + IBA (3) Ternary System at 101.3 kPaa: Temperature (T), Mole Fraction of Liquid Phase (xi), Mole Fraction of Vapor Phase (yi), and Activity Coefficients (γi)

Table 3. VLE Data for 1-Propanol (1b) + Water (2) + IBA (3) Ternary System at 101.3 kPaa: Temperature (T), Mole Fraction of Liquid Phase (xi), Mole Fraction of Vapor Phase (yi) and Activity Coefficients (γi)

T (K)

x1a

x2

y1a

y2

γ1a

γ2

γ3

T (K)

x1b

x2

y1b

y2

γ1b

γ2

γ3

352.65 352.75 352.92 353.46 353.55 353.57 353.97 354.38 354.78 355.38 356.00 356.44 356.48 356.66 356.81 356.88 356.92 357.15 357.28 357.47 357.83 358.08 358.24 358.71 359.30 359.58 359.78 359.87 360.12 360.30 360.36 360.64 360.81 360.84 361.52 361.65 361.72 361.81 362.29 362.72 363.03 363.51 363.73 368.91 373.35

0.888 0.767 0.644 0.695 0.519 0.813 0.578 0.423 0.463 0.328 0.335 0.516 0.236 0.404 0.641 0.540 0.666 0.416 0.252 0.318 0.152 0.305 0.236 0.167 0.201 0.155 0.280 0.387 0.246 0.262 0.176 0.245 0.489 0.573 0.098 0.255 0.165 0.150 0.371 0.051 0.016 0.180 0.171 0.105 0.062

0.044 0.162 0.287 0.189 0.424 0.065 0.301 0.522 0.410 0.620 0.542 0.198 0.714 0.325 0.061 0.152 0.037 0.277 0.612 0.416 0.801 0.375 0.506 0.687 0.468 0.569 0.271 0.153 0.299 0.263 0.388 0.264 0.041 0.075 0.435 0.202 0.298 0.320 0.079 0.411 0.456 0.194 0.198 0.090 0.037

0.919 0.807 0.710 0.753 0.637 0.878 0.657 0.578 0.570 0.524 0.478 0.609 0.468 0.500 0.775 0.652 0.812 0.514 0.407 0.420 0.401 0.402 0.340 0.307 0.290 0.246 0.375 0.518 0.341 0.359 0.254 0.336 0.700 0.600 0.150 0.364 0.241 0.222 0.544 0.082 0.029 0.278 0.262 0.191 0.139

0.059 0.170 0.262 0.203 0.334 0.080 0.292 0.385 0.364 0.427 0.437 0.272 0.463 0.373 0.107 0.224 0.074 0.348 0.475 0.438 0.493 0.436 0.496 0.534 0.513 0.552 0.410 0.274 0.426 0.408 0.508 0.430 0.092 0.184 0.566 0.371 0.487 0.499 0.195 0.588 0.584 0.398 0.424 0.291 0.189

0.989 1.002 1.044 1.005 1.135 0.997 1.035 1.225 1.088 1.381 1.205 0.980 1.646 1.020 0.991 0.987 0.995 1.000 1.302 1.057 2.085 1.032 1.121 1.407 1.081 1.177 0.986 0.982 1.008 0.990 1.041 0.979 1.016 0.742 1.059 0.983 1.004 1.014 0.987 1.067 1.189 0.996 0.981 0.976 1.040

2.906 2.262 1.953 2.251 1.643 2.571 1.991 1.488 1.764 1.335 1.526 2.560 1.204 2.119 3.228 2.702 3.666 2.277 1.398 1.883 1.085 2.031 1.701 1.324 1.827 1.599 2.478 2.927 2.304 2.491 2.096 2.582 3.546 3.868 1.993 2.804 2.486 2.364 3.686 2.095 1.855 2.922 3.024 3.790 5.128

0.991 0.989 1.231 1.122 1.503 1.012 1.221 1.918 1.455 2.573 1.836 1.082 3.595 1.208 1.011 1.026 0.975 1.134 2.183 1.331 5.545 1.229 1.534 2.579 1.374 1.670 1.081 1.016 1.140 1.084 1.204 1.038 0.954 1.322 1.279 1.019 1.056 1.093 0.963 1.228 1.447 1.001 0.954 1.002 0.979

360.91 361.08 361.23 361.28 361.38 361.41 361.55 361.84 361.87 362.02 362.35 362.52 362.96 363.04 363.13 363.16 363.23 363.45 363.91 364.14 364.90 365.10 365.22 365.54 365.77 366.36 366.51 367.98 368.34 368.88 369.50 369.84 370.59

0.445 0.360 0.244 0.545 0.278 0.385 0.371 0.259 0.464 0.640 0.375 0.169 0.563 0.115 0.270 0.735 0.464 0.165 0.347 0.069 0.246 0.028 0.124 0.073 0.845 0.064 0.359 0.762 0.028 0.261 0.129 0.195 0.188

0.513 0.598 0.722 0.409 0.568 0.451 0.597 0.468 0.372 0.315 0.346 0.427 0.271 0.413 0.342 0.221 0.269 0.341 0.272 0.345 0.261 0.304 0.271 0.269 0.114 0.233 0.180 0.071 0.183 0.135 0.136 0.202 0.112

0.396 0.367 0.341 0.431 0.263 0.312 0.371 0.213 0.360 0.485 0.288 0.134 0.433 0.092 0.209 0.564 0.360 0.128 0.272 0.054 0.198 0.023 0.100 0.060 0.681 0.055 0.303 0.696 0.027 0.237 0.122 0.165 0.185

0.577 0.601 0.618 0.545 0.624 0.592 0.597 0.617 0.551 0.492 0.559 0.633 0.482 0.637 0.578 0.415 0.497 0.596 0.518 0.627 0.524 0.600 0.564 0.569 0.299 0.544 0.426 0.199 0.492 0.381 0.415 0.482 0.353

1.274 1.451 1.977 1.116 1.330 1.138 1.397 1.136 1.070 1.039 1.039 1.066 1.016 1.055 1.016 1.006 1.014 1.006 0.998 0.989 0.987 1.000 0.977 0.984 0.956 0.998 0.974 0.997 1.039 0.959 0.976 0.863 0.976

1.761 1.564 1.324 2.058 1.690 2.018 1.528 1.994 2.238 2.348 2.398 2.186 2.582 2.231 2.437 2.707 2.655 2.491 2.669 2.524 2.713 2.647 2.779 2.792 3.437 2.992 3.018 3.395 3.209 3.306 3.496 2.696 3.475

1.391 1.637 2.576 1.111 1.556 1.240 2.107 1.295 1.127 1.055 1.116 1.166 1.016 1.136 1.082 0.939 1.050 1.086 1.051 1.029 1.033 1.026 1.004 1.006 0.862 0.985 1.008 1.017 0.973 0.987 0.960 0.882 0.964

a

Standard uncertainties u. u(P) = 0.1 kPa, u(T) = 0.02 K, u(x) = 0.002, and u(y) = 0.002.

Table 4. VLLE Data for 1-Propanol (1b) + Water (2) + IBA (3) Ternary System at 101.3 kPaa: Mole Fraction of OrganicRich Phase (xiI), Mole Fraction of Aqueous Rich Phase (xiII), Mole Fraction of Vapor Phase (yi) and Temperature (T) organic-rich phase, xiI

a

Standard uncertainties u. u(P) = 0.1 kPa, u(T) = 0.02 K, u(x) = 0.002 and u(y) = 0.002.

ϕî ϕio

exp

−ViL(P − Pio) RT

vapor phase, yi

xI1b

xI2

xII1b

xII2

y1b

y2

T/K

0.011 0.043 0.082 0.091 0.096

0.549 0.554 0.638 0.733 0.746

0.001 0.006 0.015 0.031 0.042

0.972 0.971 0.957 0.927 0.905

0.010 0.042 0.090 0.130 0.144

0.662 0.657 0.651 0.649 0.644

362.68 362.30 362.29 362.55 362.10

a Standard uncertainties u. u(P) = 0.1 kPa, u(T) = 0.02 K, u(x) = 0.002 and u(y) = 0.002.

correction factor: Φi =

aqueous-rich phase, xiII

The fugacity coefficient of component i in the vapor phase can be calculated from the second virial coefficient, B:

(3)

N

where ϕ̂ i is the fugacity coefficient of component i in the vapor phase, ϕoi is the fugacity coefficient of pure saturated liquid i and VLi is the molar liquid volume of component i calculated by the Rackett equation.24

ln ϕî =

P (2 ∑ y Bij − B) RT j = 1 j

(4)

where Bij is the cross second virial coefficient. 2699



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Table 5. Vapor Pressure Parametersa ethanol (1a) 1-propanol (1b) water (2) IBA (3)

Ai

Bi

Ci

Di

Ei

74.475 79.463 73.649 187.79

−7164.3 −8294.9 −7258.2 −12955.0

−7.3270 −8.9096 −7.3037 −24.285

3.1340 × 10−6 1.8197 × 10−6 4.1653 × 10−6 1.4262 × 10−5

2 2 2 2

a

Parameters and vapor pressure equation obtained from DIPPR tables.23 Vapor pressure equation: ln P°(Pa) = A + B/T(K) + C ln T(K) + D (T(K))E.

Figure 1. Experimental VLE and VLLE data for the ternary system ethanol (1a) + water (2) + IBA (3) at 101.3 kPa. VLE data: (●) liquid phase; (▲) vapor phase. VLLE data taken from a previous work:12 (○) organic liquid phase; (□) aqueous liquid phase; (△) vapor phase; (---) tie lines. (★) binary azeotrope with data taken from refs 29 and 30.

With respect to the fugacity coefficient of a pure saturated liquid, ϕi°, it can be obtained assuming as standard state the pure component at the pressure and temperature of the solution ln ϕio =

Bii Pio RT

The thermodynamic consistency of VLE data was checked, in both systems, with the Wisniak and Tamir modification27 of the McDermott−Ellis test.28 This test is based on parameters Da and Dmaxa, obtained as follows: 3

Da =

(5)

ln γi = ln

yP i xiPio

P + 2RT

+

(Bii −

n

−

a

i=1

b

(8)

⎛ ⎞ 1 1 1 1 + xib)⎜⎜ + + + ⎟⎟Δx yi xib yi ⎠ ⎝ xia a b

3

+

3

∑ (x i

a

i=1 3

+

∑ (x i

a

⎛ ΔP ⎞ ⎟ + 2 ∑ |ln γ − ln γ |Δx + xib)⎜ ib ia ⎝ P ⎠ i=1 + xib)Bj [(Ta + Cj)−2 + (Tb + Cj)−2 ]ΔT

i=1

n

(9)

where the subscripts a and b refer to two consecutive experimental points. Δx, ΔP, and ΔT are the standard uncertainty of the liquid phase mole fraction, the pressure and the temperature, respectively. To be thermodynamically consistent, Da has to be smaller than Dmaxa in all experimental points. As this condition is

(6)

where δij is obtained from

δij = 2Bij − Bjj − Bii

a

∑ (x i

Pio)

∑ ∑ yj yk (2δji − δjk)

+ xib)(ln γi − ln γi )

3

Dmaxa =

RT

j=1 k=1

a

i=1

where Bii is the second virial coefficient of the pure component. The virial coefficients Bii and Bij were estimated by the method of Hayden and O’Connell25 using the molecular parameters suggested by Prausnitz et al.26 Given eqs 1−5, the follow expression is used to obtain the activity coefficients: ViL)(P

∑ (x i

(7) 2700



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Figure 2. Experimental VLE and VLLE data for the ternary system 1-propanol (1b) + water (2) + IBA (3) at 101.3 kPa. VLE data: (●) liquid phase; (▲) vapor phase. VLLE data: (○) organic liquid phase; (□) aqueous liquid phase; (△) vapor phase; (---) tie lines. (★) binary azeotrope with data taken from refs 30 and 31.

Table 6. Parameters for NRTL and UNIQUAC Models Obtained for the Ternary System Ethanol (1a) + Water (2) + IBA (3) UNIQUACa model

NRTL model component

Δgij

Δgji

correlation type

i+j

(J/mol)

(J/mol)

VLE + VLLEb

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

373.49 −2627.45 12555.58 555.16 2164.42 6824.01 1318.92 −2997.28 12992.01

4882.75 4902.04 −1179.74 3784.91 −1893.04 1250.41 4757.45 13415.05 −1355.42

VLE

VLLEb

a

Δuij

Δuji

αij = αji

(J/mol)

(J/mol)

0.30 0.30 0.30 0.30 0.30 0.30 0.30 0.30 0.30

225.65 −461.54 3427.14 1831.34 −382.30 −517.08 288.10 −2410.53 3231.17

1164.27 795.61 −906.77 −442.41 480.24 3441.27 1190.63 13468.34 −584.84

Volume and surface parameters from DECHEMA.34 bExperimental data taken from Pla-Franco et al.12

parameters have to be capable of estimating correctly the VLE data and the binodal curve. For this purpose, the suitable way is correlate jointly both kinds of data.26 So, the overall objective function is the sum of two objective functions: one used to correlate VLE data (FO1) and other one used to correlate VLLE data (FO2), according to the following equations:

satisfied in both studied systems, ternary VLE data can be considered thermodynamically consistent. Experimental data corresponding to the ethanol (1a) + water (2) + IBA (3) and 1-propanol (1b) + water (2) + IBA (3) systems are plotted in Figures 1 and 2, respectively. Additionally, Figure 1 also includes VLLE data for the ethanol system obtained in a previous work.12 According to the plots, neither of the two systems present a ternary azeotrope. Thus, only binary azeotropes are plotted, although data are taken from literature.29−31 The heterogeneous zone in both systems exhibits a type I behavior.32 As can be seen that the heterogeneous zone covers a small range of compositions. Data Correlation. Experimental data have been correlated in order to obtain the parameters of the NRTL and UNIQUAC models. It is always preferable to obtain a unique set of parameters for each model and system. This means that the same

⎛ exp Ti − Ticalc ⎜ + ∑⎜ Tiexp i=1 ⎝ N

FO1 =

⎛

N

FO2 =

Tiexp − Ticalc + Tiexp ⎝

∑ ⎜⎜ i=1

2

+

2

k=1 j=1

2701

exp

∑ ∑ |x k ji

⎞ exp calc ⎟ y y | − | ∑ ji ji ⎟ ⎠ j=1 2

(10)

2

∑ |yjexp − yjcalc | j=1

⎞ exp − x k j |⎟⎟ i ⎠

i

i

(11) DOI: 10.1021/acs.jced.7b00098 J. Chem. Eng. Data 2017, 62, 2697−2707


Article

Table 7. Parameters for NRTL and UNIQUAC Models Obtained for the Ternary System 1-Propanol (1b) + Water (2) + IBA (3) UNIQUACa model

NRTL model component

Δgij

Δgji

correlation type

i+j

(J/mol)

(J/mol)

VLE + VLLE

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

917.83 −5272.20 11581.38 83.48 1746.54 10726.98 1831.51 −5316.10 11586.45

5113.56 12361.77 −277.04 8027.03 −1752.85 220.33 4254.76 13299.73 −285.85

VLE

VLLE

a

Δuij

Δuji

αij = αji

(J/mol)

(J/mol)

0.30 0.30 0.30 0.30 0.30 0.30 0.30 0.30 0.30

−1599.04 4417.89 2154.53 985.51 726.85 1385.60 −1269.70 48596.52 2288.47

3715.48 −2076.95 336.57 1049.53 −732.25 1084.54 3495.90 −2557.53 198.97

Volume and surface parameters from DECHEMA.34

Table 8. Correlation Statistics for the NRTL and UNIQUAC Models from the System Ethanol (1a) + Water (2) + IBA (3) Correlated data

AADT/Ka

AADy1ab

AADy2b

0.99 0.30 0.25 0.26

0.0385 0.0120 0.0113 0.0064

0.0440 0.0535 0.0112 0.0437

1.01 0.31 0.25 0.21

0.0312 0.0053 0.0112 0.0051

0.0360 0.0512 0.0115 0.0357

AADxI1ac

AADxI2c

AADxII1ac

AADxII2 c

0.0308

0.0005

0.0018

0.0140

0.0005

0.0015

0.0043

0.0339

0.0005

0.0031

0.0029

0.0138

0.0005

0.0016

NRTL Model VLE + VLLE

VLE VLLE

VLE VLLE VLE + VLLE

VLE VLLE

VLE VLLE

0.0058 0.0029 UNIQUAC Model

a

Average absolute deviation in temperature. bAverage absolute deviation in vapor phase composition. cAverage absolute deviation in liquid phase composition.

Table 9. Correlation Statistics for the NRTL and UNIQUAC Models from the System 1-Propanol (1b) + Water (2) + IBA (3) Correlated data

AADT/Ka

AADy1bb

AADy2b

1.54 0.25 0.32 0.27

0.0211 0.0250 0.0063 0.0396

0.0481 0.0366 0.0126 0.0521

2.25 0.18 0.32 0.27

0.0322 0.0259 0.0065 0.0210

0.0671 0.0315 0.0131 0.0295

AADxI1bc

AADxI2c

AADxII1bc

AADxII2 c

0.0234

0.0021

0.0025

0.0160

0.0014

0.0025

0.0055

0.0268

0.0005

0.0047

0.0035

0.0185

0.0009

0.0038

NRTL Model VLE + VLLE

VLE VLLE

VLE VLLE VLE + VLLE VLE VLLE

VLE VLLE

0.0077 0.0034 UNIQUAC Model

a

Average absolute deviation in temperature. bAverage absolute deviation in vapor phase composition. cAverage absolute deviation in liquid phase composition.

where N is the total number of points, superscripts exp and calc refer to experimental and calculated values, respectively, and the subscript k refers to the number of the liquid phase. The nonrandomness parameter αij of the NRTL model has a fixed value of 0.3, as Prausnitz recommended in one of his works.33 The volume and surface parameters used in the UNIQUAC model were taken from DECHEMA.34 Tables 6 and 7 show the binary interaction parameters obtained after correlation for the ethanol and 1-propanol system, respectively. In this way, deviations are listed in Tables 8 (ethanol mixtures) and 9 (1-propanol mixtures). According to these results, neither of the models can estimate accurately the equilibrium in the two systems, with values above 0.02 for the average absolute deviation of the vapor phase composition, indistinctly of the nature of the data. Probably, these problems arise due the presence the

heterogeneous zone. As the overall objective function is the sum of eqs 10 and 11, the correlation sacrifices some accuracy in the vapor phase composition in order to estimate better the LLE. In fact, results improve when the correlation is made only using VLE data. According to the values of the deviations shown in Tables 8 and 9, the new parameters (called from here VLE parameters and listed in Tables 6 and 7) give good results in VLE estimation with either of the two models, as any value is above 0.015. Evidently, the downside is that the VLE parameters do not estimate immiscibility. For this reason, they can be used only in applications where the presence of the heterogeneous zone is not important and can be avoided. Moreover, these VLE parameters can also be used in modern processes simulators, as they give the possibility to use different sets of parameters in function of the process equipment. Thus, a set of parameters to describe the VLE 2702



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the IBA + ethanol face. This means a reversal of the relative volatility between ethanol and water. The RCM of the other system, formed by 1-propanol (1b) + water (2) + IBA (3), is shown in Figure 5. A comparison of both figures shows that they have common characteristics. In this way, the RCM has two distillation regions delimited by the separatrix connecting the two binary azeotropes (water + 1-propanol or + IBA). The stable nodes are the water and IBA vertices, while the unstable node is the 1-propanol + water azeotrope. The other binary azeotrope and the 1-propanol vertex are saddle points. According to Figure 5, the form of the residue curves also indicates a reversal of the volatility, just as in the ethanol system. The similarity between ethanol and 1-propanol RCMs allows to propose a common two-column distillation sequence. Figure 6 shows this process. The feed stream, called F1, has a composition near the azeotropic point. This means that it has a mole fraction of 0.80 for ethanol and 0.35 for 1-propanol. This stream enters the first distillation column, called extractive column, at a stage located about the middle of the column. In addition to the mixture to separate, there is another feed stream formed solely by the entrainer. Usually, this second stream enters the column at a stage located near the condenser in order to achieve higher concentration gradients. If the proportion of these two streams is adequate, the composition profile of the extractive column will be in the region II of the RCM plotted in Figures 4 and 5. According to the residue curves of this zone, the overhead stream of the extractive column, D1, is composed mainly by water and IBA (in the two cases there is a reverse in relative volatility), with a composition near the heterogeneous binary azeotrope. Because of this, D1 has two liquid phases that can be separated using a decanter. The aqueous stream which leaves the decanter is sent to the water treatment unit and the organic stream (composed mainly by IBA) is recycled as the reflux. On the other hand, the bottom stream B1 is formed by ethanol or 1-propanol and the remaining IBA. This stream is the feed of the second distillation column, called entrainer recovery column. Here, the desired alcohol (ethanol or 1-propanol) is obtained as overhead product. The entrainer almost pure leaves the column as the bottom stream B2, which is then recirculated to the extractive column. The process also includes a fresh entrainer stream (makeup) to cover the possible losses. Solvent Effect. Figures 4 and 5 also include the isovolatility line corresponding to components to separate, built from the union of all points where the relative volatility between these components is equal to the unity. According to Laroche et al.,36 the edge intersected by the isovolatility line corresponds to the component which leaves the extractive column in the overhead stream. This statement coincides with the isovolatility lines of Figures 4 and 5, where in both cases these lines intersect the water + IBA edge. Isovolatility lines are also important because they provide information about the minimal amount of solvent required to carry out the separation. In this way, the point where the isovolatility line ends (considering the azeotrope composition as the starting point) corresponds to the minimum mole fraction of entrainer, x3, needed to break the azeotrope. According to Figures 4 and 5, x3 is equal to 0.578 for ethanol mixtures and 0.387 for the case of 1-propanol, a lower value. Another way to see the changes in relative volatility caused by the addition of the entrainer are the solvent-free diagrams. In these figures, y−x curves are plotted for several fixed solvent compositions although considering only VLE data between components to separate. Figures 7 and 8 show the solvent-free diagrams for the ethanol and 1-propanol system, respectively.

data is used in the distillation column while another set, related to VLLE, are used in the decanter. In this way, Tables 6 and 7 also include a set of parameters obtained after correlating exclusively the VLLE data, whose deviations are shown in Tables 8 and 9. To compare the VLE data of the system ethanol (1a) + water (2) + IBA (3) obtained in this work with those from literature at 101.3 kPa,15−17 these reported VLE data have been estimated using the NRTL model with the VLE parameters given in Table 6. The estimation has been made with a bubble temperature calculation and the average deviations between calculated and experimental values for temperature and vapor phase composition are shown in Table 10 and Figure 3. VLE data of Andiappan Table 10. Estimation of VLE Data from Literature15−17 Using NRTL Model with VLE Parameters from Table 6 for the System Ethanol (1a) + Water (2) + IBA (3)

a

ref

AAD T/Ka

AADy1ab

AADy2b

this work Kharin et al.15 Suska et al.16 Andiappan and McLean17

0.25 0.47 1.85 2.62

0.0113 0.0190 0.0369

0.0112 0.0166 0.0413

Average absolute deviation in temperature. deviation in vapor phase composition.

b

Average absolute

and McLean17 are T−x data, so only deviations for temperature can be calculated. According to the results of this comparison, a good agreement with the data of Kharin et al.15 is obtained; however, the deviations for the other data16,17 are higher, especially for values of x1 < 0.2. It could be due to the high values of δT/δx1 and δT/δy1 of the vapor−liquid equilibrium in this region since the difference between the boiling points of the pure components is large. By other hand, Suska et al.16 had not obtained VLE data for x3 > 0.01 and Andiappan and McLean17 had not measured vapor phase compositions, so the comparison cannot be made so accurately. Process Design. The best way to design a distillation sequence is to use one of the existing tools that make use of the equilibrium data. In this sense, residue curve maps (RCMs) are very useful constructions that allow an analysis of the system. A RCM is a triangular composition diagram where in which residue curves representing the variation of the liquid phase composition in a batch distillation are plotted. In addition, and according to Perry et al.,35 the shape adopted by the residue curves indicates if there is a reversal of the relative volatility of the components to separate. This occurs when a point of inflection is observed in the residue curves when they approach the vertex representing the entrainer. The local composition model used to obtain the RCMs is NRTL because it was the model with lower deviations (see Tables 7 and 8). As the immiscibility region has a very small range of compositions, all residue curves are calculated with the VLE parameters of Tables 6 and 7. On the other hand, binodal curves have been obtained with the VLLE parameters. Figure 4 shows the RCM for the ternary system formed by ethanol (1a) + water (2) + IBA (3). The map is divided into two distillation regions separated by the line (called separatrix) which links the two binary azeotropes: one of these azeotropes, formed by ethanol and water, is the unstable node (the lower boiling temperature in the system). The saddle points of the system are the other azeotrope, formed by water and IBA, and the pure ethanol vertex. The others diagram vertices (water and IBA) are the stable nodes of the system. As it can be seen in Figure 4, all residue curves approaching the IBA vertex are inflected toward 2703



Article

Figure 3. Deviations between calculated and experimental values for temperature and vapor phase composition using the NRTL model with VLE parameters from Table 6 for different literature15−17 VLE data of the system ethanol (1a) + water (2) + IBA (3): (●) Kharin et al.;15 (△) Suska et al.;16 (×) Andiappan and McLean.17

■

CONCLUSIONS With the aim to assess the role as entrainer of the IBA, experimental isobaric VLE data were obtained for the ternary systems ethanol (1a) + water (2) + IBA (3) and 1-propanol (1b) + water (2) + IBA (3). The two systems are partially miscible so isobaric VLLE data were also determined. However, it was only measured for the system with 1-propanol, as the VLLE data of the ethanol system were obtained in a previous work. After experimental determination, all data (VLE and VLLE) of each system were correlated using NRTL and UNIQUAC to obtain a set of parameters to estimate the equilibrium. However, values of deviations were too high and two new sets of parameters were obtained correlating separately the VLE and VLLE data. IBA does not form a ternary azeotrope in any of the two studied systems. However, it forms a heterogeneous azeotrope with water that allows the proposal of a two-column distillation sequence, the design of which is carried out using the RCM calculated with NRTL. IBA reverses the relative volatility between water and 1-propanol or ethanol. Thus, the overhead stream in the first column has a composition near the water + IBA binary azeotrope, without ethanol or 1-propanol. Since this azeotrope is heterogeneous, the obtained stream can be separated by decantation once it has condensed. Ethanol or 1-propanol, depending on the case, leaves the column together with IBA

All VLE data have been estimated with the NRTL model using VLE parameters shown in Tables 6 and 7. As IBA inverts the volatility of the components to separate, y−x curves are closer to the diagonal x = y as the IBA compositions increase. In fact, the addition of entrainer shifts to the left the azeotropic point in the solvent-free diagram until it disappears. The solvent fixed compositions are the same in both cases: 0.2, 0.4, and 0.8. The first value is below the minimum molar fraction of entrainer to break the azeotrope in both cases. For this reason, y−x curves at this solvent concentration continue to cross the diagonal x = y, indicating the azeotrope. This behavior changes for the 1-propanol system when the entrainer concentration increases until x3 = 0.4, where the curve does not cross the diagonal. However, the curve is very close to the diagonal, so it would be difficult to carry out the separation. At this composition the azeotrope has not disappear for the ethanol system, although the azeotropic point is located at a low ethanol composition. Finally, there are no azeotropes at x3 = 0.8. Furthermore, the curves are sufficiently far from the diagonal to be able to perform the separation as described in the distillation sequence of Figure 5, especially in the case of 1-propanol. Anyway, design of the extractive distillation sequences requires a complete simulation of the process in order to study their economic viability. 2704



Article

Figure 4. Residue curve map for the system ethanol (1a) + water (2) + IBA (3) at 101.3 kPa. Binary azeotropes (●), residue curves (−), binodal curve (− − −), and ethanol + water isovolatility line (···) have been calculated using the NRTL model with the parameters obtained in the VLE correlation and given in Table 6.

Figure 5. Residue curve map for the system 1-propanol (1b) + water (2) + IBA (3) at 101.3 kPa. Binary azeotropes (●), residue curves (−), binodal curve (− − −), and 1-propanol + water isovolatility line (···) have been calculated using the NRTL model with the parameters obtained in VLE correlation and given in Table 7.

forming the bottom stream. The desired alcohol is obtained as the overhead product of the second column, where it is separated from IBA. Finally, a study about the minimal amount of entrainer capable to break the azeotrope is done using isovolatility lines

and solvent-free diagrams. According to the results, the total amount of IBA required to break the azeotrope is higher for the ethanol than for 1-propanol. However, in any case IBA could be recovered almost completely and recirculated to the extractive 2705



Article

column, so that the quantity of fresh entrainer (makeup) would be relatively very small. An economic study must be carried out to analyze the feasibility of both processes.

■

AUTHOR INFORMATION

Corresponding Author

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

Sonia Loras: 0000-0002-3863-2201 Funding

Financial support from the Universitat de València (Project No. UV-INV-AE15-340195) is gratefully acknowledged. J. PlaFranco is deeply grateful for the grant BES-2011-04636 received from the Ministerio de Economiá y Competitividad.

Figure 6. Proposed two-column distillation sequence for dehydration of ethanol or 1-propanol using IBA as entrainer.

Notes

The authors declare no competing financial interest.

■

REFERENCES

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Figure 7. VLE data plotted on a solvent-free basis for the system ethanol (1a) + water (2) + IBA (3) at 101.3 kPa for x3 = 0.00 (−), 0.20 (), 0.40 (---) and 0.80 (··). All values have been calculated using the NRTL model with the VLE parameters given in Table 6.

Figure 8. VLE data plotted on a solvent-free basis for the system 1-propanol (1b) + water (2) + IBA (3) at 101.3 kPa for x3 = 0.00 (−), 0.20 (), 0.40 (---) and 0.80 (··). All values have been calculated using the NRTL model with the VLE parameters given in Table 7. 2706



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