Article Cite This: J. Chem. Eng. Data XXXX, XXX, XXX−XXX
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Isothermal Vapor−Liquid Equilibrium Measurements for the Pentan2-one + Propan-1-ol/Butan-1-ol System within 342−363 K Shivan Mavalal, Thavashni Chetty, and Kuveneshan Moodley*
J. Chem. Eng. Data Downloaded from pubs.acs.org by UNIV OF LOUISIANA AT LAFAYETTE on 04/15/19. For personal use only.
Thermodynamics Research Unit, School of Engineering, University of KwaZulu-Natal, Howard College Campus, King George V Avenue, Durban 4041, South Africa ABSTRACT: The aim of this work was to measure and model highprecision isothermal binary vapor−liquid equilibrium (VLE) phase data for the pentan-2-one + propan-1-ol/butan-1-ol system to aid in the separation process design. P−T−x−y equilibrium measurements were obtained for the pentan-2-one + propan-1-ol system at 342.34, 352.05, and 361.65 K and for the pentan-2-one + butan-1-ol system at 342.95, 352.65, and 362.55 K at subatmospheric pressures. The data were successfully modeled using the γ−Φ approach by employing the Wilson and nonrandom two-liquid activity coefficient models and the Peng and Robinson equation of state for the vapor correction. Thermodynamic consistency testing with the point and area tests was also performed for the experimental VLE data. The data sets passed both the point test with 0.01 tolerance and the area test with 10% tolerance.
1. INTRODUCTION Most industrial chemical engineering separation processes such as distillation, extraction, absorption, and adsorption rely on accurate phase equilibrium data for effective design, optimization, and simulation (Avoseh et al.1). The design of separation processes for some specific systems in industry become difficult because experimental data comprising these systems are scarce and rarely available in literature. Nonacid oxygenated hydrocarbons including alcohols, aldehydes, and ketones are found in the product stream of the Fischer− Tropsch process (de Klerk2). This includes pentan-2-one, propan-1-ol, and butan-1-ol. The separation of these products further valorizes the process as these components can be appropriately repurposed, sold, or disposed of. As an example, the chemical nature of ketones prevents them from being used as a safe solvent in many applications. Hence, they must be removed from alcohol mixtures to make the alcohols more widely useable. The phase equilibrium for the pentan-2-one + propan-1-ol system is not well studied in the literature apart from an azeotropic study at one temperature (366.62 K) by Lecat.3 For the pentan-2-one + butan-1-ol system, only one isobaric study by Seo et al.4 was found in the literature. Isothermal vapor− liquid equilibrium (VLE) data are preferred for the design of high-purity separation procedures such as enhanced distillation, as they allow for tray to tray energy balance calculations via the heat of mixing. To improve the literature on these systems, isothermal P−x−y phase equilibria data have been measured for the pentan-2-one + propan-1-ol system at 342.34, 352.05, and 361.65 K and for the pentan-2-one + butan-1-ol system at 342.95, 352.65, and 362.55 K using a dynamic apparatus designed by Raal and Mühlbauer.5 Temperatures were selected to be within the range considered in the © XXXX American Chemical Society
common industrial processes for these separations as they form part of light naptha.6,7 The VLE data were processed using the combined (γ−Φ) approach. Thermodynamic consistency testing was conducted using both the area and point tests.
2. THEORY 2.1. Modeling Approach. The γ−Φ approach has been reviewed in detail by Raal and Mühlbauer5 and is used extensively for modeling low-pressure VLE data.1,8−14 An activity coefficient is used to account for liquid-phase nonideality as an ideal solution reference state can be assumed at low pressures, whereas vapor-phase nonideality is accounted for by the fugacity coefficient in solution via an equation of state yi ΦiP = xiγiPisat
(1)
where xi and yi are the liquid and vapor compositions of component i, P is the pressure, Psat i is the saturation pressure of component i, and γi is the activity coefficient of component i. The vapor correction factor Φi is given by ÅÄÅ ÑÉ ϕî ÅÅ −ViL(P − Pisat) ÑÑÑ Å ÑÑ Φi = sat expÅÅ ÑÑ ÅÅÅ RT ÑÑÖ ϕî Ç (2) where ϕ̂ i and ϕ̂ sat i are the fugacity coefficient in solution and fugacity coefficient in solution at saturation for component i, VLi is the molar volume of the liquid, R is the universal gas constant, and T is the temperature. Received: January 29, 2019 Accepted: March 29, 2019
A
DOI: 10.1021/acs.jced.9b00100 J. Chem. Eng. Data XXXX, XXX, XXX−XXX
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Table 1. Chemical Suppliers and Purities refractive index (RI) at 293.15 K and 0.101 MPaa componentc
CAS RN.
supplier
experimental
literatureb
minimum stated mass fraction purity
GC peak relative area (mass fraction purity)
propan-1-ol butan-1-ol pentan-2-one
71-23-8 71-36-3 107-87-9
Sigma-Aldrich Sigma-Aldrich Merck
1.3848 1.3989 1.3893
1.3850 1.3988 1.3895
≥0.995 ≥0.995 ≥0.990
0.9999 0.9999 0.9999
a
Standard uncertainties u are u(RI) = 0.001, u(T) = 0.01 K, u(P) = 0.002 MPa. bHaynes28 at 293.15 K. cPurified by molecular sieving.
Figure 1. Schematic of the apparatus of Joseph et al.21 used in this work. 1Equilibrium chamber. 2Liquid sampling port. 3Temperature measurement. 4Boiling chamber. 5Variable heat supply to boiler. 6Heater cartridge and sleeve. 7Magnetic stirrer and bead. 8Boiler drain valve. 9Condensate drain valve. 10Vapor condensate sampling point. 11Condenser. 12Coolant line to condenser. 13Coolant bath and controller. 14Chiller. 15Pressure measurement. 16Isolation valves. 17Ballast tank. 18Cold trap. 19Vacuum pump. ∂ 2(GM − ∑ xiGi)/RT
Usually the virial equation of state is employed to determine ϕ̂ i and ϕ̂ sat i as it does not require binary interaction parameters
∂xi 2 ÄÅ ÑÉ ÅÄÅ ÅÅ τ G (1 − G ) Å τ G τ12G12 ÑÑÑ 21 ÑÑ − (1 − 2x1)ÅÅÅÅ 21 21 = −2ÅÅÅÅ 21 21 + Ñ ÅÅ (x1 + G21x 2)2 ÅÅÅ x1 + G21x 2 x 2 + G12x1 ÑÑÑÖ ÅÇ Ç ÄÅ ÉÑ ÅÅ τ G (1 − G )2 τ G (1 + G21) ÑÑÑÑ ÅÅ 21 21 21 (2 ) x x + + 12 12 Ñ 1 2 Å ÅÅ 2 Ñ Ñ (x 2 + G12x1) ÑÑÖ ÅÅÇ (x1 + G21x 2)3 ÑÉ τ G (1 + G12)2 ÑÑÑÑ 1 + 12 12 >0 ÑÑ + (x 2 + G12x1)3 ÑÑÑÖ x1x 2 (5)
for application. However, cubic equations of state can also be employed.11 Stability tests are used to ensure that the two-liquid phase formation does not occur in VLE measurements. For azeotropic systems with close boiling components, this test is necessary as the gradient of the P−x curve is small near the azeotrope and the two-liquid phase formation may be overlooked. The criterion for stability is ∂ 2(GM − ∑ xiGi) /RT ∂xi 2
where τij and Gij are the model parameters defined in the original NRTL model.15 2.2. Model Selection. In this work, the Wilson16 and NRTL15 activity coefficient models were employed to account for liquid-phase non-idealities, and the Peng and Robinson17 equation of state was used to account for the vapor-phase nonideality. The virial equation of state with the Hayden and O’Connell18 correlation was initially considered to account for the vapor-phase nonidealities of the ketone−alcohol systems measured in this work, as was done previously.19 However, the equation provided a poor correlation of the vapor phase of the systems considered in this work. Additionally, the nonrandom lattice fluid equation of state with the hydrogen bonding equation of state (NLF-HB EOS)20 was considered, as was used in the work of Seo et
>0 (3)
where GM is the Gibbs energy of the mixture and Gi are the pure component Gibbs energies. GM is given by GM = RT ∑ xi ln(xiγi)
(4)
For the nonrandom two-liquid (NRTL) model, eq 3 becomes B
DOI: 10.1021/acs.jced.9b00100 J. Chem. Eng. Data XXXX, XXX, XXX−XXX
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al.4 Seo et al.,4 however, extended the fitting procedure of the NLF-HB EOS by including an additional regressed binary interaction parameter to account for steric hindrances of hydrogen bond formation. This addition is purely empirical. Hence, the procedure was not explored further in this work. The selection of the activity coefficient models and cubic equation of state used in this work is based on the superior performance of these model combinations, in comparison to the other models considered, in correlating the VLE data presented in this work, as assessed by root mean square deviation (rmsd) in pressure, absolute average deviation (AAD) in vapor composition, and the quality of the prediction of the experimental activity coefficients and relative volatilities.
ability. The procedures outlined in the NIST JCGM22 guide were used for all uncertainty calculations.
4. RESULTS AND DISCUSSION The pure component vapor pressures of the chemicals used were determined using the dynamic method. These results are presented in Table 2 along with a comparison to literature and Table 2. Experimental Vapor Pressure Data and Comparison to Literaturea P/kPa literature
3. EXPERIMENTAL SECTION 3.1. Materials. The pentan-2-one used was sourced from Merck and the alcohols were sourced from Sigma-Aldrich. Supplier mass purities were stated to be >99 mass % and the chemicals were stored under a molecular sieve for 48 h before use. The purities of the pure chemicals were confirmed by performing gas chromatography (GC) sampling, Karl-Fischer titration, and refractive index measurements. For the GC analysis, a Shimadzu GC 2014 was used, fit with a POROPAKQ column (2 m × 2.2 mm). Helium was used as the carrier gas and a thermal conductivity detector was employed. GC conditions were optimized to be T = 513.15 K for the injector, and column and detector temperature with a carrier gas flow of 30 mL/min. This same GC and conditions were used for the analysis of equilibrium samples. GC results for the pure components provided relative GC peak areas of >99.99%. A Karl-Fischer (MKS 500) apparatus was used to determine the water content of the chemicals used. These were found to be less than 0.0004 mol fraction. Refractive indices of the pure chemicals were determined at T = 293.15 K using an ATAGO RX-7000α refractometer (sodium D-line = 589 nm) with a supplier uncertainty of 0.0001. These results along with other chemical identifiers are presented in Table 1. 3.2. Equipment and Uncertainties. The dynamic apparatus and procedure of Joseph et al.21 were used to conduct the VLE measurements. In Figure 1, a schematic of this apparatus is shown. Pressure was maintained by an automatic pressure controller (ABB F080) that used vacuum and atmospheric air to maintain the desired pressure in the still. The temperature was controlled using a voltage supply to a heater cartridge within the boiling chamber of the apparatus. Temperature loss is prevented by employing insulation. The temperature was measured using a type-A Pt-100 probe. This was calibrated using a WIKA CTB 9100 temperature standard. The standard combined uncertainty for temperature was found to be 0.05 K. The pressure was measured using a WIKA P-10 transducer which was calibrated using a WIKA CPH 6000 standard. The supplier uncertainty was stated as 0.05 kPa. The standard combined uncertainty in pressure was found to be 0.11 kPa. The area ratio method of Raal and Mühlbauer5 was used to calibrate the thermal conductivity detector of the gas chromatograph. Standard mixtures were prepared gravimetrically using a Mettler-Toledo mass balance (model AB204-S) with an uncertainty of 0.0001 g. The standard combined uncertainty in composition was calculated to be ±0.004 mol fraction. All uncertainties were calculated by the propagation of errors of type A and B and included supplier uncertainty, uncertainty from calibration, and uncertainty from repeat-
component
T/K
experimental
Antoine correlation of Poling et al.23
propan-1-ol
342.34
31.5
31.50
352.05
49.5
48.52
361.65
73.2
72.31
342.34
32.7
33.44
342.95 352.05
34.3 47.8
34.21 47.53
352.65 361.65
48.7 66.2
48.54 65.82
362.55 342.95
68.1 13.4
67.79 13.32
352.65
21.4
21.33
362.55
33.4
33.31
pentan-2-one
butan-1-ol
measured data from the literature 31.12 (342.11 K)14 49.92 (352.12 K)29 72.00 (361.64 K)10 33.04 (342.1 K)30 46.57 (351.46 K)31 64.58 (361.12 K)32 13.33 (342.92 K)33 21.03 (343.10 K)34 33.33 (362.52 K)33
a Standard combined uncertainties uc are uc(T) = 0.05 K, uc(P) = 0.11 kPa.
Antoine prediction using the parameters reported in Poling et al.23 A close correlation between the experimental and literature data is observed, with minor differences being attributed to uncertainties in the experimental temperature and pressure measurements. These data confirm the accuracy of the temperature and pressure measurements. The isobaric data of Seo et al.4 were measured in a pressure and temperature range beyond that considered in work. Additionally, isothermal VLE data at the conditions considered in this work were not found in the literature. Hence, P−T−x−y comparisons with the data obtained in this work were not possible. The results of the vapor−liquid equilibrium data measurements are presented in Tables 3 and 4 and graphically in Figures 2 and 3. The data were modeled via the γ−Φ approach with the Wilson6 and NRTL15 activity coefficient models with the Peng−Robinson17 equation of state (PR-EOS) for the vapor phase. It can be observed in Table 4 that the activity coefficients are less than 1 for some compositions in the dilute regions. This is indicative of a negative deviation from Raoult’s law and is attributed to strong attractive forces in the liquid phase between the ketone and alcohol. A similar behavior has been reported.1,4 The area test of Redlich and Kister24 and the point test of Christiansen and Fredenslund25 for thermodynamic consisC
DOI: 10.1021/acs.jced.9b00100 J. Chem. Eng. Data XXXX, XXX, XXX−XXX
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Table 3. Vapor−Liquid Equilibrium Data for the Pentan-2-one (1) + Propan-1-ol (2) Systema T/K = 342.34
T/K = 352.05
P/kPa
x1
y1
ln γ1
31.50 32.60 33.20 33.78 34.40 34.80 35.30 36.40 38.00 38.70 38.95 39.00 38.98 38.60 38.32 37.95 37.24 36.58 35.60 34.70 34.38 34.05 33.55 32.70
0.0000 0.0118 0.0205 0.0309 0.0425 0.0527 0.0694 0.1174 0.2116 0.2918 0.3982 0.4497 0.4620 0.5880 0.6293 0.6809 0.7579 0.8168 0.8861 0.9351 0.9499 0.9641 0.9810 1.0000
0.0000 0.0384 0.0600 0.0850 0.1091 0.1219 0.1446 0.2042 0.2999 0.3570 0.4167 0.4502 0.4571 0.5490 0.5905 0.6303 0.7008 0.7610 0.8408 0.9018 0.9231 0.9428 0.9666 1.0000
1.1729 1.0891 1.0454 0.9940 0.9002 0.8110 0.6606 0.4990 0.3703 0.2201 0.1772 0.1652 0.0973 0.0949 0.0717 0.0518 0.0413 0.0324 0.0231 0.0215 0.0181 0.0109 0.0000
T/K = 361.65
ln γ2
P/kPa
x1
y1
ln γ1
0.0000 0.0071 0.0114 0.0124 0.0159 0.0238 0.0296 0.0411 0.0688 0.1092 0.1812 0.2128 0.2220 0.2937 0.2956 0.3335 0.3789 0.4154 0.4575 0.5107 0.5155 0.5434 0.6270
49.50 51.40 52.20 52.65 53.50 53.70 54.40 55.30 56.50 57.20 57.80 57.81 57.80 57.00 56.40 55.70 54.60 53.20 51.50 49.90 49.40 48.90 48.30 47.80
0.0000 0.0150 0.0247 0.0342 0.0507 0.0570 0.0781 0.1205 0.2045 0.2908 0.3941 0.4461 0.4577 0.5827 0.6311 0.6846 0.7532 0.8178 0.8862 0.9399 0.9548 0.9699 0.9869 1.0000
0.0000 0.0438 0.0672 0.0862 0.1158 0.1230 0.1538 0.1942 0.2635 0.3184 0.3847 0.4260 0.4295 0.5229 0.5605 0.6096 0.6750 0.7460 0.8295 0.9068 0.9280 0.9509 0.9780 1.0000
1.1453 1.0914 1.0204 0.9383 0.8854 0.8075 0.6230 0.4209 0.2702 0.1658 0.1440 0.1263 0.0677 0.0467 0.0369 0.0234 0.0151 0.0084 0.0071 0.0044 0.0030 0.0013 0.0000
ln γ2
P/kPa
x1
y1
ln γ1
0.0000 0.0080 0.0085 0.0064 0.0067 0.0089 0.0086 0.0233 0.0552 0.1049 0.1704 0.1908 0.2058 0.2750 0.3058 0.3313 0.3732 0.4044 0.4441 0.4470 0.4636 0.4771 0.4943
73.20 76.50 77.80 78.80 79.50 80.00 80.70 81.40 82.30 82.90 82.90 82.55 82.50 81.20 80.00 78.50 76.40 74.10 71.40 69.50 68.70 67.90 66.90 66.20
0.0000 0.0145 0.0254 0.0357 0.0469 0.0561 0.0796 0.1207 0.2064 0.2940 0.3961 0.4426 0.4533 0.5772 0.6348 0.6874 0.7525 0.8201 0.8906 0.9355 0.9516 0.9693 0.9870 1.0000
0.0000 0.0396 0.0565 0.0718 0.0845 0.0989 0.1226 0.1603 0.2320 0.3004 0.3770 0.4130 0.4182 0.5202 0.5607 0.6110 0.6758 0.7529 0.8375 0.8989 0.9226 0.9500 0.9773 1.0000
1.1499 0.9599 0.8748 0.7720 0.7556 0.6296 0.4909 0.3345 0.2467 0.1754 0.1514 0.1395 0.1003 0.0652 0.0526 0.0357 0.0272 0.0141 0.0087 0.0061 0.0052 0.0006 0.0000
ln γ2 0.0000 0.0183 0.0285 0.0355 0.0423 0.0425 0.0497 0.0600 0.0844 0.1152 0.1557 0.1720 0.1818 0.2301 0.2736 0.2885 0.3130 0.3296 0.3710 0.3977 0.4067 0.4146 0.4674
a
Standard combined uncertainties uc are uc(T) = 0.05 K, uc(P) = 0.11 kPa, uc(xi) = 0.005, uc(yi) = 0.005.
Table 4. Vapor−Liquid Equilibrium Data for the Pentan-2-one (1) + Butan-1-ol (2) Systema T/K = 342.95 P/kPa
x1
y1
13.40 13.48 14.85 16.20 17.50 21.10 23.20 23.20 25.60 27.40 29.10 29.58 31.65 32.70 33.50 34.20 34.29 34.30
0.0000 0.0023 0.0299 0.0592 0.0866 0.1931 0.2602 0.2675 0.3760 0.4655 0.5615 0.6027 0.7471 0.8473 0.9168 0.9896 0.9943 1.0000
0.0000 0.0112 0.1253 0.2206 0.2910 0.4765 0.5520 0.5579 0.6422 0.6959 0.7531 0.7740 0.8482 0.9016 0.9428 0.9920 0.9956 1.0000
T/K = 352.65 ln γ1
0.6533 0.5957 0.5663 0.5389 0.4174 0.3612 0.3441 0.2428 0.1775 0.1292 0.1021 0.0466 0.0144 0.0044 −0.0005 0.0010 0.0000
ln γ2
P/kPa
x1
y1
0.0000 −0.0031 −0.0008 0.0014 0.0136 0.0214 0.0472 0.0440 0.0912 0.1514 0.2010 0.2276 0.3488 0.4522 0.5408 0.6744 0.6835
21.40 21.78 23.21 25.08 26.40 31.61 32.38 34.68 38.20 40.30 41.42 44.20 45.51 47.10 47.58 48.57 48.68 48.70
0.0000 0.0065 0.0274 0.0587 0.0818 0.1885 0.2085 0.2716 0.3885 0.4701 0.5186 0.6635 0.7442 0.8566 0.8948 0.9895 0.9941 1.0000
0.0000 0.0280 0.1060 0.2000 0.2550 0.4440 0.4686 0.5270 0.6259 0.6751 0.7070 0.7870 0.8310 0.8988 0.9220 0.9916 0.9952 1.0000
T/K = 362.55 ln γ1
0.6603 0.6111 0.5628 0.5252 0.4245 0.4018 0.3234 0.2341 0.1726 0.1480 0.0737 0.0426 0.0147 0.0067 −0.0005 0.0007 0.0000
ln γ2
P/kPa
x1
y1
0.0000 −0.0043 −0.0031 −0.0040 0.0009 0.0120 0.0157 0.0510 0.0881 0.1438 0.1638 0.2680 0.3400 0.4403 0.4999 0.5937 0.6122
33.40 33.68 36.10 38.10 40.10 46.92 51.37 51.72 54.41 56.54 60.62 62.18 64.40 66.08 67.31 68.00 68.05 68.10
0.0000 0.0039 0.0298 0.0567 0.0825 0.1923 0.2839 0.2920 0.3587 0.4221 0.5617 0.6272 0.7315 0.8254 0.9138 0.9899 0.9945 1.0000
0.0000 0.0158 0.1059 0.1812 0.2390 0.4198 0.5200 0.5288 0.5880 0.6260 0.7130 0.7490 0.8110 0.8650 0.9299 0.9912 0.9952 1.0000
ln γ1 0.7019 0.6348 0.5807 0.5339 0.4082 0.3233 0.3187 0.2698 0.2081 0.1222 0.0865 0.0473 0.0168 0.0058 −0.0002 −0.0001 0.0000
ln γ2 0.0000 −0.0037 −0.0040 −0.0098 −0.0042 0.0091 0.0305 0.0301 0.0455 0.0912 0.1726 0.2259 0.3055 0.4251 0.4940 0.5727 0.5825
a
Standard combined uncertainties uc are uc(T) = 0.05 K, uc(P) = 0.11 kPa, uc(xi) = 0.005, uc(yi) = 0.005.
tency were also conducted using the ASPEN Plus V9 software package, with default consistency criteria of 10% tolerance for the area test and 0.01 for the point test. The results, presented in Table 5, confirm that the VLE data were thermodynamically consistent. The pentan-2-one + propan-1-ol system displays an azeotropic behavior at all temperatures measured in this work. In the temperature range considered here, the volatilities of the components change such that pentan-2-one transitions
from the more volatile component to the less volatile component. The pentan-2-one + butan-1-ol system does not exhibit an azeotrope in the range measured here. The behavior was found to be mostly conventional with a slight pinching of the VLE curve in the pentan-2-one-rich region. A comparable behavior has been reported for ketone/alcohol systems by Avoseh et al.,1 Mali and co-workers,8,9 and Moodley and Dorsamy.19 Conventional distillation would be suitable for this separation D
DOI: 10.1021/acs.jced.9b00100 J. Chem. Eng. Data XXXX, XXX, XXX−XXX
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for the PR-EOS. For this fitting procedure the pressure residual was minimized N
δP =
∑ (P exp − P calc)2 (6)
k=1 exp
where N is the number of measured data points and P and Pcalc are the measured and model-calculated pressures. The rmsd in pressure and AAD (δy1) in vapor composition were calculated according to the method of Van Ness and Abbott26 RMSD = Figure 2. Vapor−liquid equilibrium data (P−x−y) for the pentan-2one (1) + propan-1-ol (2) system at: (exp. x1, exp. y1, model x1, model y1): 342.34 K, (●, ○, − − −, -·-); 352.05 K, (■, □, - - -, -··-); 361.65 K, (▲, △, −, ···). The red lines represent the Wilson−PR model and the black lines are the NRTL−PR model.
δy1 =
δP N
(7)
abs(y1exp − y1calc )
(8) N The regressed model parameters and fitting deviations are presented in Table 6. The Wilson PR-EOS model performs
Table 6. Regressed Model Parametersa system pentan-2-one (1) + propan-1-ol (2)
pentan-2-one (1) + butan-1-ol (2)
parameter
Wilsona
NRTLb
Wilsona
NRTLb
a12 a21 b12/K b21/K α12,NRTL kij rmsd/kPa δy1b
−3.786 2.838 806.505 −900.000
−2.315 2.195 618.706 −149.547 0.300 1.167 0.0506 0.0049
−3.885 4.072 1227.183 −1529.147
−2.8827 2.732 1100.000 −817.930 0.3 0.216 0.0531 0.0035
0.992 0.0387 0.0039
0.050 0.0541 0.0028
a
Model parameters can be related to those described in the original works by the following expressions: aAij = exp(aij + bij/T), bτij = (aij + bij/T), Gij = exp(αij,NRTLτij). bAbsolute average deviation: δy1 = calc abs(yexp 1 − y1 )/N.
Figure 3. Vapor−liquid equilibrium data (P−x−y) for the pentan-2one (1) + butan-1-ol (2) system at: (exp. x1, exp. y1, model x1, model y1): 342.95 K, (●, ○, − − −, -·-); 352.65 K, (■, □, - - -, -··-); 362.55 K, (▲, △, −, ···). The red lines represent the Wilson−PR model and the black lines are the NRTL−PR model.
slightly better for the pentan-2-one + propan-1-ol system, whereas the performances of the two models are quite similar for the case of the pentan-2-one + butan-1-ol system. Because of the behavior of the pentan-2-one + propan-1-ol system, stability tests were conducted by evaluating eqs 3 and 4, to confirm that the two-phase formation did not occur in the experimental measurement range. These results are presented in Figure 4. It can be observed that two-phase splitting was not
Table 5. Results of Thermodynamic Consistency Tests Using the NRTL−PR Model calculated criterion system T/K = 342.34 T/K = 352.05 T/K = 361.65 T/K = 342.95 T/K = 352.65 T/K = 362.55
area test (%)
point test
consistency test result
Pentan-2-one (1) + propan-1-ol (2) 0.589 0.008 passed 2.675 0.010 passed 2.311 0.009 passed Pentan-2-one (1) + butan-1-ol (2) 0.610 0.006 passed 1.285 0.004 passed 5.902 0.008 passed
both tests both tests both tests
exhibited by the systems as
∂ 2(G M − ∑ xiGi) / RT ∂xi 2
is positive
everywhere, and GM/RT is always negative without convexity. The activity coefficient behavior for each system at the highest measured temperature is shown in Figures 5 and 6. The activity coefficients are reasonably well represented by the models for the pentan-2-one + propan-1-ol system, and very well represented in the case of the pentan-2-one + butan-1-ol system. This is attributed to the difficulty of fitting the highly nonideal pentan-2-one + propan-1-ol system, caused by the intersection of the pure component vapor pressure curves of the two components in the measurement range considered (yielding the azeotropic behavior). The asymmetry of the activity coefficient plots for each component of this system is also attributed to the highly nonideal behavior. It also seems
both tests both tests both tests
but may require a large number of stages if high-purity pentan2-one is desired. The Wilson and NRTL models with the PR-EOS describe the experimental data well. The PR-EOS was chosen over the virial equation of state as it described the vapor phase behavior more precisely. For the activity coefficient, a single set of temperature-dependent model parameters was regressed for each system, along with the binary interaction parameter, kij, E
DOI: 10.1021/acs.jced.9b00100 J. Chem. Eng. Data XXXX, XXX, XXX−XXX
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The regressed model parameters were used to predict the behavior of the pentan-2-one + propan-1-ol system under isobaric conditions at 80, 50, and 30 kPa to assess the effect of pressure on the position of the azeotrope under isobaric conditions. The position of the azeotrope shifts from approximately 0.65 mol fraction pentan-2-one at 80 kPa to 0.59 at 50 kPa and 0.52 kPa at 30 kPa. This suggests that pressure swing distillation is a viable option for the separation of these two components from a thermodynamic stand-point. The experimental and model calculated relative volatilities y / x1 y ij jjα12 = y1 / x zzz are presented in Figures 7 and 8 and were 2{ 2 k
Figure 4. Plots to show the stability of the pentan-2-one (1) + propan-1-ol (2) system using the NRTL−PR model at: 342.95 K, (− − −); 352.65 K, (- - -); 362.55 K, (−). The left axis and red lines represent ∂2(GM − ∑xiGi)/RT/∂xi2 vs x1; the right axis and black lines represent GM/RT vs x1.
Figure 7. α12 vs x1 for the pentan-2-one (1) + propan-1-ol (2) system at: (exp, model): 342.34 K, (●, − − −); 352.05 K, (■, - - -); 361.65 K, (▲, −). The red lines represent the Wilson−PR model and the black lines are the NRTL−PR model.
Figure 5. γi vs x1 for the pentan-2-one (1) + propan-1-ol (2) system at 361.65 K. γ1−x1: ▲experimental, - - - model; γ2−x1: ■ experimental, −model. The red lines represent the Wilson−PR model and the black lines are the NRTL−PR model.
Figure 8. α12 vs x1 for the pentan-2-one (1) + butan-1-ol (2) system at: (exp, model): 342.95 K, (●, − − −); 352.65 K, (■, - - -); 362.55 K, (▲, −). The red lines represent the Wilson−PR model and the black lines are the NRTL−PR model.
generally found to decrease with increasing temperature. The experimental and model-calculated relative volatilities correlate quite well and tend to converge in the pentan-2-one-rich region. Small deviations are attributed to the very high sensitivity of α12 to small differences between experimental and calculated composition values (Gmehling et al.27). The maximum deviations between experimental and calculated relative volatilities occurred in the dilute regions. For the pentan-2-one + propan-1-ol system the relative volatilities were between 2.85 and 14.88% for the propan-1-ol dilute region and
Figure 6. γi vs x1 for the pentan-2-one (1) + butan-1-ol (2) system at 362.55 K. γ1−x1: ●experimental, - - - - model; γ2−x1: ■ experimental, −model. The red lines represent the Wilson−PR model and the black lines are the NRTL−PR model.
from these figures that the infinite dilution activity coefficient can be reasonably well represented by the models used. F
DOI: 10.1021/acs.jced.9b00100 J. Chem. Eng. Data XXXX, XXX, XXX−XXX
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between 0.53 and 14.88 for the pentan-2-one dilute region. For the pentan-2-one + butan-1-ol system, the relative volatilities were between 0.38 and 7.00% for the propan-1-ol dilute region and between 1.68 and 5.85 for the pentan-2-one dilute region. In Figures 9−12, the plots of the excess Gibbs energy and excess enthalpy are presented. A positive GE and HE are
Figure 11. GE/RT vs x1 for the pentan-2-one (1) + butan-1-ol (2) system at: (exp, model): 342.95 K, (●, − − −); 352.65 K, (■, - - -); 362.55 K, (▲, −). The red lines represent NRTL−PR model. (×) calculated from the data of Seo et al.4 at 101.3 kPa.
Figure 9. GE/RT vs x1 for the pentan-2-one (1) + propan-1-ol (2) system at: (exp, model): 342.34 K, (●, − − −); 352.05 K, (■, - - -); 361.65 K, (▲, −). The red lines represent the NRTL−PR model.
Figure 12. HE/T2 vs x1 for the pentan-2-one (1) + butan-1-ol (2) system at: (exp, model): 342.95 K, (●, − − −); 352.65 K, (■, - - -); 362.55 K, (▲, −). The red lines represent the NRTL−PR model.
5. CONCLUSIONS The vapor−liquid equilibrium phase behavior of propan-1-ol and butan-1-ol with pentan-2-one was successfully determined using a dynamic apparatus operating at subatmospheric pressures. Consistency and stability tests were conducted and passed, which confirmed the validity of the results obtained. The VLE behavior was found to be nonideal with azeotropic behavior being observed for the pentan-2-one + propan-1-ol system. The Wilson PR-EOS and NRTL PR-EOS correlated the data well. The rmsd values in pressure and AAD values in vapor phase mole fraction were found to be within the experimental uncertainty. The lowest rmsd value for the pentan-2-one + propan-1-ol system was 0.0378 and was 0.0531 for the pentan-2-one + butan-1-ol system.
Figure 10. HE/T2 vs x1 for the pentan-2-one (1) + propan-1-ol (2) system at: (exp, model): 342.34 K, (●, − − −); 352.05 K, (■, - - -); 361.65 K, (▲, −). The red lines represent the NRTL−PR model. (×)literature data at 298.15 K,35 (−··−), prediction based on NRTL−PR model parameters from this work at 298.15 K.
observed for both systems. A reasonably good correlation between experimental and model-calculated excess Gibbs energies was observed. Experimental HE data for the pentan2-one + propan-1-ol system was found in the literature and compared to extrapolations by the model parameters that were regressed from the VLE data measured in this work. An excellent correlation is observed, as shown in Figure 10. In Figure 11, the GE/RT versus x1 plots for the pentan-2-one + butan-1-ol system are presented and compared to the data of Seo et al.4 It can be seen that the general shape of the GE/RT versus x1 plots conform to the literature data, and the trend of decreasing GE/RT with increasing temperature is preserved. Differences are attributed to the differences in measurement conditions.
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AUTHOR INFORMATION
Corresponding Author
*E-mail:
[email protected]. Phone: +27 31 2601519. ORCID
Kuveneshan Moodley: 0000-0003-1544-3624 Notes
The authors declare no competing financial interest. G
DOI: 10.1021/acs.jced.9b00100 J. Chem. Eng. Data XXXX, XXX, XXX−XXX
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ACKNOWLEDGMENTS This work is based upon research supported by the JW Nelson Fund awarded by the University of KwaZulu-Natal.
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DOI: 10.1021/acs.jced.9b00100 J. Chem. Eng. Data XXXX, XXX, XXX−XXX
