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Fluid Phase Equilibria 255 (2007) 121–130

Vapor–liquid equilibrium, densities, and interfacial tensions for thesystem ethyl 1,1-dimethylethyl ether (ETBE) + propan-1-ol

Andres Mejıa ∗, Hugo Segura ∗, Marcela Cartes, Pıa BustosDepartamento de Ingenierıa Quımica, Universidad de Conception POB 160 - C, Correo 3, Concepcion, Chile

Received 23 February 2007; received in revised form 6 April 2007; accepted 10 April 2007Available online 13 April 2007

bstract

Isobaric vapor–liquid equilibrium data at 50, 75, and 94 kPa have been determined for the binary system ETBE + propan-1-ol, in the temperatureange 325–368 K. The measurements were made in a vapor–liquid equilibrium still with circulation of both phases. Mixing volumes have been alsoetermined from density measurements at 298.15 K and 101.3 kPa and, at the same temperature and pressure, the dependence of interfacial tensionn concentration has been measured using the pendant drop technique. According to experimental results, the mixture presents positive deviationrom ideal behavior and azeotropy is present at 75 and 94 kPa. No azeotrope was detected at 50 kPa. The mixing volumes of the system are negative

ver the whole mole fraction range, and the interfacial tensions exhibit negative deviation from the linear behavior. The activity coefficients andoiling points of the solutions were well correlated with the mole fraction using the Wohl, Wilson, NRTL, UNIQUAC equations. Excess volumeata and interfacial tensions were correlated using the Redlich–Kister model.

2007 Elsevier B.V. All rights reserved.

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eywords: Vapor–liquid equilibrium; Densities; Interfacial tension; ETBE; Pro

. Introduction

Due to the environmental regulations placed by the Clean Airct of 1990 on automobile emissions, refineries worldwide have

ncreased continuously their use of alcohol- and ether-basedxygenates in gasoline blending. As well known, methanol,thanol, methyl tert-butyl ether (MTBE), ethyl tert-butyl etherETBE), tert-amyl methyl ether (TAME) and diisopropyl etherDIPE) are typical examples of oxygenates, and some of themave been used widely as commercial fuel additives. As followsrom the constant technological renewal joined to the accumu-ated experience in fuel production, alcohol- and ether-basedxygenates exhibit recognized advantages and inconveniencesor gasoline blending [1,2]. For example, alcohol–gasolinelends improve fuel’s octane rating and reduce carbon monoxide

missions in spark ignition engines. However, as polar sol-ents, alcohols show the inconvenience of driving off nonpolarydrocarbons from gasoline, thus increasing the emissions of

∗ Corresponding authors. Tel.: +56 41 2204534; fax: +56 41 2247491.E-mail addresses: amejia@udec.cl (A. Mejıa), hsegura@diq.udec.cl

H. Segura).

bisaRitI

378-3812/$ – see front matter © 2007 Elsevier B.V. All rights reserved.oi:10.1016/j.fluid.2007.04.009

-ol

azardous air pollutants [3]. Ethers act also as octane enhancerslthough, due to its inherent biochemistry and to the risksnvolved in its massive storage and distribution, they resultarmful to the environment [4]. Considering that ethers formuasi-ideal mixtures with hydrocarbons and dissolve alcohols inll proportion, they may be used as potentially attractive cosol-ents for improving the affinity of alcohols in gasoline mixtures.n fact, gasoline formulations containing mixtures of ethers andlcohols display interesting properties. For example, Weber deenezes et al. [5] reported a new gasoline formulation contain-

ng an azeotropic mixture of ETBE and ethanol that offers cleardvantages over pure ethanol (low volatility and low solubil-ty in water) and pure ETBE (higher octane number and lowerroduction cost) formulations.

An appropriate gasoline blending requires finding a balanceetween the combustion performance and the volatility of thenvolved fuel mixtures. Key parameters in the formulation andtorage of commercial gasoline are the distillation curve, whichffects the evolution of fuel’s combustion, together with the

eid’s vapor pressure (RVP) which is widely used as a volatility

ndicator. Both the distillation curve and RVP depend directly onhe vapor–liquid equilibrium (VLE) behavior of fuel mixtures.n addition, density and excess volume (VE) data are required to

1 Equi

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2

2

C(wial

ppsspda((wu5sDivia

2

2

mmcpflgaluwtbmPtc

TM

C

EP

22 A. Mejıa et al. / Fluid Phase

djust the engine injection system (pump and injectors), whilenterfacial tension (σ) affects the ignition time and the atomiza-ion of the fuel inside the engine’s combustion chamber [6]. Itollows then that VLE data and additional thermophysical prop-rties are required for predicting the properties of oxygenatednd related mixtures. Isothermal T–P–y measurements for sys-em ETBE + propan-1-ol have been reported previously by Ohnd Park at 333.15 K, together with density data measurementst 298.15 K [7]. According to these experimental results, theLE of the quoted mixture exhibits positive deviation from

deal behavior and no azeotrope is present. In addition, the VE

ata shows negative deviation from ideal behavior. Althoughot complete, due to the experimental determinations basedn headspace gas chromatography, the results of Oh and Parkre in good agreement with the general trends observed inhe VLE behavior of ether + alcohol mixtures (which alwayseviate positively from ideal behavior [7–11]). In addition, thexperimental results reported for VE data of ether + alcohol mix-ures usually exhibit negative deviation [7,8,10,11]. However, ithould be pointed out that depending on either the tempera-ure level or the alcohol, azeotropic behavior may be presentn these mixtures. Furthermore, experimental σ data is scarcer, simply, unavailable for alcohol + ether mixtures. The presentork was undertaken to measure complete VLE, VE and σ dataf the binary mixture ETBE + propan-1-ol for which scarce,ncomplete or no data have been reported previously in thexperimental range proposed in this contribution. Furthermore,onsidering the technical relevance of alcohol + ether azeotropicixtures in the formulation of gasoline, an additional objective

f the present experimental investigation is to detect evidencef azeotropy for the quoted system.

. Experimental

.1. Purity of materials

ETBE (>96.0 mass%) was purchased from TCl (Tokyohemical Industry Co. Ltd., Japan) and propan-1-ol

99.9 mass%) was purchased from Merck. Then, ETBE

as further purified to more than >99.9 mass% by rectification

n a 1 m height and 30 mm diameter Normschliffgeratebaudiabatic distillation column (packed with 3 mm × 3 mm stain-ess steel spirals), working at a 1:100 reflux ratio. Appropriate

haea

able 1ass% GC purities (mass%), refractive index (nD) at Na D line, densities (ρ), norma

omponent (purity/mass%) nD (298.15 K) ρ (298.15 K; g cm

Experimental Literature Experimental

TBE (>99.9) 1.37319 1.37290a 0.73562ropan-1-ol (>99.9) 1.38333 1.38370d 0.79954

a [26].b [31].c [32].d [33].e [34].f [35].

libria 255 (2007) 121–130

recautions were taken when handling ETBE in order to avoideroxide formation. Propan-1-ol was dried using 3 A molecularieves. After these steps, gas chromatography (GC) failed tohow any significant impurity. The properties and purity of theure components, as determined by GC, appear in Table 1. Theensities and refractive indexes of pure liquids were measuredt 298.15 K using an Anton Paar DMA 5000 densimeterAustria) and a Multiscale Automatic Refractometer RFM 81Bellingham + Stanley, England), respectively. Temperatureas controlled to ±0.01 K with a thermostated bath. Thencertainties in density and refractive index measurements are× 10−6 g cm−3 and ±10−5, respectively. The interfacial ten-

ions of pure fluids were measured at 298.15 K using a Pendantrop tensiometer IFT-10 (Temco, USA). The uncertainties

n interfacial tension are ±0.01 mN m−1. The experimentalalues of these properties and the boiling points are givenn Table 1 together with those given in the literature whenvailable.

.2. Apparatus and procedure

.2.1. Vapor–liquid equilibrium cellAn all-glass vapor–liquid equilibrium apparatus model 601,

anufactured by Fischer Labor und Verfahrenstechnik (Ger-any), was used in the equilibrium determinations. In this

irculation-method apparatus, the mixture is heated to its boilingoint by a 250 W immersion heater. The vapor–liquid mixtureows through an extended contact line (Cottrell pump) thatuarantees an intense phase exchange and then enters to a sep-ration chamber whose construction prevents an entrainment ofiquid particles into the vapor phase. The separated gas and liq-id phases are condensed and returned to a mixing chamber,here they are stirred by a magnetic stirrer, and returned again

o the immersion heater. The temperature in the VLE still haseen determined with a Systemteknik S1224 digital temperatureeter, and a Pt 100 � probe calibrated at the Swedish Statensrovningsanstalt. The accuracy is estimated as ±0.02 K. The

otal pressure of the system is controlled by a vacuum pumpapable of work under vacuum up to 0.25 kPa. The pressure

as been measured with a Fischer pressure transducer calibratedgainst an absolute mercury-in-glass manometer (22-mm diam-ter precision tubing with cathetometer reading), the overallccuracy is estimated as ±0.03 kPa.

l boiling points (Tb) and interfacial tensions (σ) of pure components

−3) Tb (101.33 kPa; K) σ (298.15 K; mN m−1)

Literature Experimental Literature Experimental Literature

0.73513a 345.84 345.86b 19.73 19.78c

0.79960d 370.30 370.30e 23.89 23.39f

Equi

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TsctcaewTamTiCgabattefpsDi

sa

dtdpiCiTistdtawitc

σ

wttneltdT

Istdar

ttpbibi2do

A. Mejıa et al. / Fluid Phase

On the average the system reaches equilibrium conditionsfter 2–3 h operation. Samples, taken by syringing 1.0 �L afterhe system had achieved equilibrium, were analyzed by GC on

Varian 3400 apparatus provided with a thermal conductiv-ty detector and a Thermo Separation Products model SP4400lectronic integrator. The column was 3 m long and 0.3 cmn diameter, packed with SE-30. Column, injector and detec-or temperatures were (353.15, 423.15, 493.15) K, respectively.ood separation was achieved under these conditions, and cali-ration analyses were carried out to convert the peak ratio to theass composition of the sample. The pertinent polynomial fit

ad a correlation coefficient R2 better than 0.99. At least threenalyses were made of each sample. Concentration measure-ents were accurate to better than ±0.001 in mole fraction.

.2.2. Density measurementsFor density measurements, samples of known mass were pre-

ared on an analytical balance (Chyo Balance Corp., Japan)ith an accuracy of ±10−4 g. Densities of the pure compo-ents and their mixtures were measured then using a DMA000 densitymeter (Anton Paar, Austria) with an accuracy of× 10−6 g cm−3. The density determination is based on mea-

uring the period of oscillation of a vibrating U-shaped tubelled with the liquid sample. The temperature of the apparatusas maintained constant to within ±0.01 K.

.2.3. Interfacial tension cellA pendant drop tensiometer model IFT-10, manufactured by

emco Inc. (USA) [12], was used in interfacial tension mea-urements. The pendant drop cell is a stainless steel cylindricalhamber (with an inner volume of ∼42 cm3), with two injec-ion orifices (one at the top and the other at the bottom of thehamber) where stainless steel needles are placed for gener-ting pendant (or sessile) drops and bubbles. The chamber isquipped with appropriately sealed borosilicate glass windowshich allow visualization of the inner space during operation.he light beam source, located at one side of the visualizationxis, is a quartz halogen bulb (SII P/N 240-350, Scientific Instru-ent Inc, USA) covered by a white diffuser made of teflon.he camera, located at opposite side of the visualization axis,

s a monochrome video camera model CS8320Bi (Toshiba Teli,orp., Japan) connected to a personal computer through a framerabber card. The temperature of the cell is measured by meansK-type thermocouple, and maintained constant to within ±1 Ky means of electric band heaters operated by a Watlow temper-ture controller model TC-211-K-989 (USA). The tensiometer,he light source and the camera are mounted on a free vibra-ion table (Vibraplane, model 2210, USA) in order to avoid theffect of noisy measurements due to external vibrations. Inter-acial tension measurements were made by analyzing images ofendant drops generated at the tip of a injection needle, which isubmerged into the fluid mixture that fills the cell, by using theROPimage [13] Advanced software version 1.5 (Rame-Hart

nstruments Co. USA).The experimental procedure for determining interfacial ten-

ion is as follows. The mixture to be analyzed is prepared bydding appropriate volumes of each pure fluid and then it is

iAee

libria 255 (2007) 121–130 123

egassed in an ultrasonic bath. After degasification, the concen-ration of the sample is measured by GC. The cell is heated to theesired experimental temperature and, later on, the mixture isumped through a stainless steel tube to the needle tip. The pumps a positive displacement ELDEX HP Series Model B-100-S-2E (USA). Initially, a small portion of the mixture is pumped

nto the chamber in order to saturate the vapor that fills the cell.hen, a liquid drop is generated at the tip of the needle and it

s maintained inside the chamber during 30 min. Using GC, thetability of the concentration of the liquid drop is checked. Afterhis equilibration step, a new liquid drop is generated and theirimensions are recorded (at least during 1 h) in order to checkhe stability of its geometry. The shape and volume of the dropre characterized by two diameters: the equatorial diameter (de),hich is the largest one, and the horizontal diameter (ds), which

s located at a distance de from the apex of the drop. Whenhe shape and volume of the drop are constant, the softwarealculates the interfacial tension according to the equation:

= �ρ gcd2e

H(1)

here σ is the interfacial tension (dyne cm−1 or mN m−1), gche local gravitational constant (≈981 cm seg−2), �ρ (g cm−3)he difference between the densities of the fluid at the tip of theeedle and the fluid that fills the chamber. In the case consid-red in this work, �ρ = ρL − ρV where ρ is the density of theiquid (L) and vapor phases (V), respectively, at the experimentalemperature, ρL is measured using the Anton Paar densitymeterescribed before, and vapor phase is assumed to be an ideal gas.he H constant is evaluated from:

1

H= f

(ds

de

)= f (S) (2)

n Eq. (2), the value of H as function of S corresponds to theolution of the Laplace’s Capillary Equation [14] and it wasaken from a numerical table. Usually, a least the images of tenifferent liquid drops are analyzed for each concentration. Inddition, for each drop, interfacial tension measurements wereepeated 20 times.

The chamber cleaning plays a key role in the accuracy ofhe pendant drop technique, since small impurity concentra-ions strongly affect σ measurements. Consequently, appropriaterecautions were taken when cleaning the chamber surfacesy replicating experimental a values of well characterized flu-ds. The appropriate diameter of sampling needles was selectedy generating drops of reproducible shapes and sizes for thenvolved mixtures. Usually, we tested needles from 1.00 to.00 mm i.d. Camera magnification, i.e. the relation betweenimensions and pixels, was calibrated by reproducing the valuef the diameter of a needle of known dimensions, at fixed light

ntensity and camera operation parameters (position and zoom).dditional details concerning to the pendant drop technique have

xtensively described by Rusanov and Prokhorov [15], Andreast al. [16], Ambwani and Fort [17].

124 A. Mejıa et al. / Fluid Phase Equilibria 255 (2007) 121–130

Table 2Experimental vapor–liquid equilibrium data for ETBE (1) + propan-1-ol (2) at50.00 kPa

T (K) x1 y1 γ1 γ2

352.76 0.000 0.000 1.000347.18 0.054 0.267 2.345 0.989343.40 0.100 0.408 2.178 0.997340.59 0.143 0.505 2.054 0.998338.14 0.191 0.579 1.918 1.008336.15 0.236 0.635 1.821 1.019334.53 0.284 0.678 1.707 1.037333.04 0.334 0.713 1.603 1.072332.12 0.371 0.733 1.530 1.106330.94 0.428 0.760 1.436 1.156329.91 0.482 0.784 1.362 1.213329.19 0.535 0.802 1.288 1.282328.48 0.584 0.819 1.236 1.359327.78 0.635 0.836 1.189 1.455327.26 0.683 0.852 1.147 1.560326.84 0.732 0.865 1.103 1.718326.34 0.771 0.881 1.086 1.821325.70 0.831 0.903 1.057 2.082325.36 0.909 0.935 1.014 2.602325.07 0.961 0.968 1.003 3.0173

3

3

v9w

TE7

T

333333333333333333333

Table 4Experimental vapor–liquid equilibrium data for ETBE (1) + propan-1-ol (2) at94.00 kPa

T (K) x1 y1 γ1 γ2

368.34 0.000 0.000 1.000363.28 0.056 0.229 2.254 0.995359.97 0.103 0.359 2.117 0.995357.19 0.145 0.443 2.007 1.018355.03 0.190 0.512 1.890 1.029353.28 0.237 0.567 1.772 1.043351.88 0.282 0.609 1.667 1.063350.46 0.333 0.652 1.575 1.084349.43 0.375 0.678 1.501 1.119348.42 0.427 0.706 1.417 1.166347.44 0.481 0.733 1.344 1.224346.69 0.537 0.756 1.272 1.295346.01 0.584 0.775 1.225 1.367345.30 0.637 0.796 1.179 1.468344.94 0.685 0.814 1.135 1.567344.46 0.733 0.833 1.102 1.700344.08 0.773 0.850 1.079 1.826343.58 0.832 0.878 1.052 2.050343.46 0.910 0.922 1.015 2.43733

f

γ

wvapor pressure. In Eq. (3), the vapor phase is assumed to be an

24.94 1.000 1.000 1.000

. Results and discussions

.1. Vapor–liquid equilibrium

The equilibrium temperature T, liquid-phase x andapor–phase y mole fraction measurements at P = 50, 75, and

4 kPa are reported in Tables 2–4 and in Figs. 1–4, togetherith the activity coefficients (γ i) that were calculated from the

able 3xperimental vapor–liquid equilibrium data for ETBE (1) + propan-1-ol (2) at5.00 kPa

(K) x1 y1 γ1 γ2

62.57 0.000 0.000 1.00057.21 0.055 0.244 2.334 0.99553.75 0.103 0.381 2.151 0.99351.23 0.145 0.465 2.009 1.00348.81 0.191 0.538 1.904 1.01646.84 0.238 0.594 1.791 1.03745.36 0.282 0.638 1.701 1.04744.03 0.333 0.674 1.589 1.07843.07 0.372 0.700 1.522 1.10241.73 0.429 0.729 1.434 1.16540.92 0.484 0.754 1.350 1.21640.19 0.536 0.773 1.279 1.29139.56 0.586 0.792 1.225 1.36238.85 0.637 0.811 1.181 1.46238.43 0.685 0.827 1.134 1.57637.94 0.732 0.846 1.104 1.69037.46 0.773 0.862 1.083 1.82436.96 0.831 0.888 1.054 2.04936.68 0.910 0.927 1.015 2.52236.60 0.961 0.964 1.002 2.87536.56 1.000 1.000 1.000

iil

F(s

43.46 0.961 0.962 1.002 2.76343.48 1.000 1.000 1.000

ollowing equation [18]:

i = yiP

xiP0i

(3)

here P is the total pressure and P0i is the pure component

deal gas and the pressure dependence of the liquid phase fugac-ty is neglected. These simplifications are reasonable due to theow pressures observed in the present VLE data. The tempera-

ig. 1. Boiling temperature diagram for the system ETBE (1) + propan-1-ol2). Experimental data at (�) 50.00 kPa; (�) 75.00 kPa; (�) 94.00 kPa; (—)moothed by fitting a three parameter Legendre polynomial.

A. Mejıa et al. / Fluid Phase Equilibria 255 (2007) 121–130 125

F5L

tc

l

w

a

F7L

Fig. 4. Activity coefficients for the system ETBE (1) + propan-1-ol (2) at94.00 kPa. (�) Experimental data; (—) smoothed by fitting a three parameterLegendre polynomial.

Table 5Antoine coefficients, Eq. (4)

Compound Ai Bi Ci Temperature range (K)

ig. 2. Activity coefficients for the system ETBE (1) + propan-1-ol (2) at0.00 kPa. (�) Experimental data; (—) smoothed by fitting a three parameteregendre polynomial.

ure dependence of the pure component vapor pressure P0i was

alculated using the Antoine equation

og P0i (kPa) = Ai − Bi

T (K) − C(4)

i

here the Antoine constants Ai, Bi and Ci are reported in Table 5.The calculated activity coefficients are reported in Tables 2–4

nd are estimated accurate to within ±2%. The results

ig. 3. Activity coefficients for the system ETBE (1) + propan-1-ol (2) at5.00 kPa. (�) Experimental data; (—) smoothed by fitting a three parameteregendre polynomial.

ETBEa 5.96651 −1151.7300 −55.060 328.15–370.15 KP

ro(ahof

f

wbtTop

TE

P

579

ropan-1-ol 6.82728 −1414.0237 −77.032

a [36].

eported in these tables indicate that, for the pressure rangef the measurements, the liquid phase of the mixture ETBE1) + propan-1-ol (2) deviates positively from ideal behaviornd azeotropy is present at 75 and 94 kPa. No azeotropeas been detected at 50 kPa. The azeotropic concentrationsf the measured binaries were estimated by fitting theunction

(x) = 100

(y − x

x

)(5)

here f(x) is an empirical interpolating function and x, y haveeen taken from the experimental data. Azeotropic concen-rations, as determined by solving f(x) = 0, are indicated in

able 6, from which it is concluded that the mole fractionf the azeotrope impoverishes in ETBE as pressure (or tem-erature) increases. The trend of the azeotropic concentration

able 6stimated azeotropic coordinates for the system ETBE (1) + propan-1-ol (2)

ressure (kPa) xAz1 TAz (K)

0 – –5 0.972 336.514 0.955 343.36

1 Equilibria 255 (2007) 121–130

iact

mNmtrevLantt

Table 7Consistency test statistics for the binary system ETBE (1) + propan-1-ol (2)

Pressure (kPa) L1a L2

a L3a 100 × �yb δP (kPa)c

50.00 1.0415 0.1679 0.0610 0.4 0.175.00 1.0222 0.1522 0.0444 0.3 0.294.00 1.0120 0.1489 0.0392 0.2 0.3

a Parameters for the Legendre polynomial [21] used in consistency.b Average absolute deviation in vapor phase mole fractions �y =

(1/N)∑N

i=1|yexpi − ycal

i | (N: number of data points). ∑P

Fe

26 A. Mejıa et al. / Fluid Phase

s in agreement with Wrewki’s law [19], according to whichpositive azeotrope becomes impoverished in the component

haracterized by the lowest heat of vaporization as pressure (oremperature) increases.

The VLE data reported in Tables 2–4 were found to be ther-odynamically consistent by the point-to-point method of Vaness et al. [20] as modified by Fredenslund et al. [21]. For eachixture, consistency criterion (�y < 0.01) was met by fitting

he equilibrium vapor pressure according to the Barker’s [22]eduction method. Statistical analysis reveals that a three param-ter Legendre polynomial is adequate for fitting the equilibriumapor pressure in each case. Pertinent consistency statistics andegendre polynomial parameters are presented in Table 7. In

ddition, as it can be seen in Fig. 5, vapor phase residuals doot scatter randomly about the zero line. As explained in [23],his trend of vapor phase residuals may be attributed to the facthat the present consistency analysis neglects the contribution of

tG

W

ig. 5. Residuals for vapor phase mole fractions and equilibrium pressure, as calculatquilibrium pressure residuals δP.

c Average absolute deviation in vapor pressure δP = (1/N)N

i=1|Pexpi −

cali |.

he heat of mixing when approximating numerically the excessibbs energy.The VLE data reported in Tables 2–4 were correlated with the

ohl [24], NRTL [25], Wilson [26] and UNIQUAC [27] equa-

ed from consistency analysis. (�) Vapor phase mole fraction residuals δy, (©)

A. Mejıa et al. / Fluid Phase Equilibria 255 (2007) 121–130 127

Table 8Parameters and prediction statistics for different GE models

Model P (kPa) A12 A21 α12 Bubble-point pressures Dew-point pressures

�P (%)a 100 × �yi �P (%) 100 × �xi

Wohl 50.00 0.885 1.200 0.629b 0.45 0.3 0.75 0.475.00 0.884 1.173 0.754b 0.34 0.2 0.51 0.394.00 0.884 1.133 0.780b 0.43 0.1 0.42 0.1

NRTLc 50.00 2485.84 934.31 0.470 0.31 0.5 0.83 0.775.00 2550.14 937.24 0.470 0.30 0.3 0.60 0.494.00 2509.14 981.70 0.470 0.38 0.1 0.38 0.1

Wilsonc,d 50.00 −939.70 4469.38 0.28 0.4 0.76 0.475.00 −926.23 4487.66 0.30 0.3 0.58 0.494.00 −880.02 4418.23 0.35 0.1 0.39 0.1

UNIQUACc,e 50.00 1890.82 −768.42 0.39 0.3 0.80 0.475.00 1891.08 −768.20 0.32 0.3 0.54 0.394.00 1805.05 −716.50 0.40 0.1 0.39 0.2

a �P = 100N

N∑i

|Pexpi

−Pcali

|P

expi

.

b “q” parameter for the Wohl’s model.c Parameters in J mol−1.

tt

O

dtWtpf

mdpwbs

3

Twl

V

wc

newenegative, behavior that may be explained in terms of cross associ-ation between components, as expected for specific interactionsbetween the polar alkanol and the aprotic aliphatic ether. Theexcess volume data have been correlated using a three parameter

d Liquid volumes have been estimated from the Rackett equation [37].e Molecular parameters are those calculated from UNIFAC [21].

ions, whose adjustable parameters were obtained by minimizinghe following objective function (OF):

F =N∑

i=1

(|Pexp

i − Pcali |

Pexpi

+ |yexpi − ycal

i |)2

(6)

From the results presented in Table 8, it is possible toeduce that all the fitted models gave a reasonable correla-ion of the binary systems, the best fit corresponding to the

ilson model. The capability of predicting simultaneouslyhe bubble- and dew-point pressures and the vapor and liquidhase mole fractions, respectively, has been used as the rankingactor.

In order to compare our experimental results with theeasurements by Oh and Park [7], we predicted the VLE

ata reported by these authors at 333.15 K using the Wilsonarameters indicated in Table 8. Results are shown in Fig. 6,here we can observe a good agreement both in the predictedubble-point (�P = 0.89%, �yi = 1.08%) and dew-point pres-ures (�P = 1.79%, �xi = 1.01%).

.2. Excess volume data

The density of the mixture and its pure constituents at= 298.15 K and P = 101.3 kPa are reported in Table 9, togetherith the excess volumes VE that were calculated from the fol-

owing equation

1 2∑ 2∑ Mi

E =ρ

i=1

xiMi −i=1

xiρi

(7)

here ρ is the density of the mixture, ρi the density of the pureomponents, and Mi is the molecular weight of pure compo-

F3i

ents whose values were taken from DIPPR [28]. The calculatedxcess volumes reported in Table 9 are estimated accurate toithin ±10−3 cm3 mol−1. Table 9 and Fig. 7 indicate that the

xcess volumes of the system ETBE (1) + propan-1-ol (2) are

ig. 6. Isothermal phase diagram for the system ETBE (1) + propan-1-ol (2) at33.15 K. (—) Predicted from the Wilson model with the parameters indicatedn Table 8. (©) Experimental data from Oh and Park [7].

128 A. Mejıa et al. / Fluid Phase Equilibria 255 (2007) 121–130

Table 9Densities and excess volumes for the binary system ETBE (1) + propan-1-ol (2)at 298.15 K and 101.3 kPa

xi ρ (g cm−3) 103 × VE (cm3 mol−1)

0.0572 0.79504 −1990.1043 0.79143 −3400.1752 0.78611 −5140.2085 0.78367 −5820.2700 0.77929 −6880.3317 0.77503 −7680.3732 0.77224 −8080.4205 0.76920 −8490.4756 0.76572 −8750.5326 0.76215 −8710.5639 0.76027 −8670.6210 0.75693 −8470.6820 0.75359 −8260.7188 0.75155 −7920.7551 0.74937 −7180.8334 0.74506 −5800.8737 0.74266 −4540.9142 0.74045 −3440.9567 0.73798 −185

Fig. 7. Excess volume for the system ETBE (1) + propan-1-ol (2) at 298.15 K and101.3 kPa. (�) Experimental data; (—) smoothed by a Redlich–Kister expansionwith the parameters shown in Table 10. (©) Experimental data from Oh and Park[7].

Table 11Interfacial tensions for the binary system ETBE (1) + propan-1-ol (2) at 298.15 Kand 101.3 kPa

x1 σ (mN m−1)

0.067 23.580.171 22.810.315 22.120.412 21.480.579 20.810.763 20.25

Fae

R

V

waVto

3

P

Table 10Coefficients and deviations (maximum, average and standard) obtained in correlati298.15 K and 101.3 kPa

c0 (cm3 mol−1) c1 (cm3 mol−1) c2 (cm3 mol−1) Maxim(103 ×

−3.5055 0.4087 −0.7674 0.29

ig. 8. Interfacial tension for the system ETBE (1) + propan-1-ol (2) at 298.15 Knd 101.3 kPa. (�) Experimental data; (—) smoothed by a Redlich–Kisterxpansion with the parameters shown in Table 12. (- - -) Linear behavior.

edlich–Kister expansion [29]

E = x1x2

m∑k=0

ck(x1 − x2)k (8)

here the ck parameters, together with the correlation statistics,re reported in Table 10. Eq. (8) was also used for predicting theE data reported by Oh and Park [7] at 298.15 K and, according

o results, a fair agreement is obtained with maximum deviationf 8% (see Fig. 7).

.3. Interfacial tension data

The interfacial tension measurements at T = 298.15 K and= 101.3 kPa are reported in Table 11 and depicted in Fig. 8.

on of excess volumes, Eq. (8), for the ETBE (1) + propan-1-ol (2) system at

um deviationcm3 mol−1)

Average deviation(103 × cm3 mol−1)

Standard deviation(103 × cm3 mol−1)

0.06 0.08

A. Mejıa et al. / Fluid Phase Equilibria 255 (2007) 121–130 129

Table 12Coefficients and deviations (maximum, average and standard) obtained in correlation of interfacial tension, Eq. (9), for the ETBE (1) + propan-1-ol (2) system at298.15 K and 101.3 kPa

c

−

TR

σ

Iioicdid

4

dfEfaaAOurtameOwea

LAcddgHL

MnPP

RSTVx

Gγ

ρ

σ

SceEL0

Sbimm

A

wdS

R

[

0 (mN m−1) c1 (mN m−1) Maximum deviation (103 × mN m−1)

2.6106 0.4381 11.18

hese experimental data were also correlated using the followingedlich–Kister expansion [30]

= x1x2

m∑k=0

ck(x1 − x2)k + x1σ1 + x2σ2 (9)

n Eq. (9) σ is the interfacial tension of the mixture while σi is thenterfacial tension of the pure components. The ck parametersf Eq. (9), together with the correlation statistics, are reportedn Table 12. From Fig. 8 it is possible to observe that the interfa-ial tensions of the mixture ETBE + propan-l-ol exhibit negativeeviation from the lineal behavior (x1σ1 + x2σ2). In addition, its possible to observe that the interfacial tension of the mixtureecreases as the ETBE concentration increases.

. Concluding remarks

In this work, consistent isobaric vapor–liquid equilibriumata (at 50, 75, and 94 kPa), excess molar volumes and inter-acial tension data at 298.15 K were measured for the systemTBE + propan-1-ol. The system exhibits positive deviation

rom ideal behavior and azeotropic behavior is present at 75nd 94 kPa. The azeotrope, when present, follows Wrewki’s lawnd converges to pure ETBE as pressure decreases to 70 kPa.t lower pressures, and in coherency with the data reported byh and Park [7], no azeotrope can be detected. The excess vol-mes of the system are negative over the whole mole fractionange, and the interfacial tensions of the system exhibits nega-ive deviation from the lineal behavior. The activity coefficientsnd boiling points of the solutions were well correlated with theole fraction using the Wohl, NRTL, Wilson, and UNIQUAC,

quations whose parameters also predicted the data reported byh and Park [7]. Excess volume data and interfacial tensionsere correlated using the Redlich–Kister expansion. The present

xcess volume data compare well with the data reported by Ohnd Park [7].

ist of symbolsi, Bi, Ci Antoine’s equation parameters, Eq. (4)k Redlich–Kister parameter, Eqs. (8) and (9)e equatorial diameters horizontal diameter of the dropc local gravitational constant (≈981 cm seg−2),

auxiliary function, Eq. (2)i Parameters for the Legendre polynomial used in con-

sistency

molecular weight

D refractive indexabsolute pressure (kPa)

0 pure component vapor pressure (kPa)

[

[

Average deviation (103 × mN m−1) Standard deviation (103 × mN m−1)

3.34 3.72

universal gas constant (J mol K−1)shape ratio (=ds/de)absolute temperature (K)volume (cm3 mol−1)

, y fractions of the liquid and vapor phases

reek lettersactivity coefficientdensity (g cm−3)interfacial tension (mN m−1)

uperscriptsal calculatexp experimental

excess propertypertaining to the liquid phasereference state (pure component)

ubscriptsnormal boiling points

, j component i, j, respectivelyix, L mixture in liquid phaseix, V mixture in vapor phase

cknowledgements

We would like to thank Mr. P.R. Brauer from Temco Inc.,ho kindly assisted us in the preparation and setup of the pen-ant drop tensiometer. This work was financed by FONDECYT,antiago, Chile (Project 1050157).

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