Vapour-liquid equilibrium data for binary mixtures of some new refrigerants

ELSEVIER

Fluid Phase Equilibria

Vapour-liquid M.H. Barley

140 (1997) IX-206

equilibrium data for binary mixtures of some new refrigerants

* , J.D. Morrison, A. O’Donnel, I.B. Parker, S. Petherbridge,

R.W. Wheelhouse

ICI Group R and T Centre, PO Box 8, Thr Heath, Runcorn, Cheshire WA7 4QD. UK Received 22 May 1996; accepted 2 1 April 1097

Abstract The acceptance of the Montreal Protocol has led to a timetable for the phasing out of chlorine-containing refrigerants and their replacement by new chlorine-free materials. For many applications a pure alternative refrigerant can not be found with the appropriate properties and refrigerant mixtures have been considered. In order to model the properties of these refrigerant blends accurate vapour-liquid equilibrium (VLE) data are required over the range of temperature and pressure of interest to the refrigeration engineer. In this paper we report VLE data for six binary mixtures of the new hydrofluorocarbon refrigerants over a wide range of temperatures and pressures. The six mixtures are: R32/R125, R32/R143a, R32/Rl34a. R12S/R143a, R125/R134a and R143a/R134a. Results for R32/R125 and R32/R134a were obtained down to at least - 30°C and were done in duplicate. The raw data were correlated to two models using Maximum Likelihood techniques. One of the models was then used to predict azeotropic compositions for three of the mixtures (R32/R125, R32/R143a and R125/R143a) and the approximate composition of a ternary saddle point azeotrope. 0 1997 Elsevier Science B.V. Kewortis:

Experimental

method; Data; Vapour-liquid

equilibria;

Mixture: Refrigerants:

Correlation

1. Introduction Concern conventional

about the depletion of the stratospheric ozone layer by chlorine atoms derived from refrigerants has led to the Montreal Protocol and subsequent international agreements

[ 1,2]. These require the phase out of chlorine

and their replacement HFC’s).

* Corresponding

containing refrigerants (known as CFC’s and HCFC’s) by non-chlorine-containing molecules (such as the hydrofluorocarbons or

author.

037X-3X12/97/$17.00 0 1997 Elsevier Science B.V. All rights reserved. PI1 SO378-38 I2(97)00 146-5

184

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Phase Equilibria

140 Cl 997) 183-206

At the present time the replacement of R12 (CCI,F,-a CFC) by the chlorine-free, ozone-friendly HFC-134a (CF,-CFH,-referred to in this paper as R134a) is being rapidly implemented. R134a is a very acceptable replacement for R12 and it is now being used commercially for domestic refrigeration and automobile air-conditioning. However replacement of other refrigerants such as R22 (CClF,H-an HCFC), R500 and R502 (both azeotropic mixtures of CFC’s and R22) is less straightforward since no single fluid could be found that has all the desired properties as a replacement for any one of these three refrigerants. Hence, to obtain a fluid with the best possible combination of properties, blends of refrigerants are being used. There is a great deal of interest in binary and ternary refrigerant blends based upon the hydrofluorocarbon refrigerants R134a, R32 (CF,H *), R 125 (CF,-CF, H), and R 143a (CF,-CH X>. Some of the mixtures of interest are non-azeotropic (also known as zeotropic) and compositions can be chosen that match the properties of each of the traditional refrigerants to be replaced. To be able to predict the behaviour of a non-azeotropic refrigerant mixture an accurate thermodynamic model is required. A number of models suitable for refrigerant mixtures have appeared in the literature and, to give two examples, Morrison and colleagues at NIST have developed the Carnahan-Starling-DeSantis equation of state (EOS) [3] while we at ICI [4] have used the Modified Huron-Vidal method with the Redlich-Kwong EOS [5] to represent the vapour phase, and for the non-ideality of the liquid phase we have used the Wilson model [6]. The thermodynamic models need to be fitted to accurate VLE data for all binary mixtures of interest in order to be able to correctly predict the properties of refrigerant blends. In this paper we report detailed VLE data for six binary mixtures. At the time of writing this paper some small data sets have been reported for all these mixtures by other workers (e.g., Fujiwara et al. [7]; Kubota and Matsumoto [8]; Widiatmo et al. [9]; and Higashi, pers. comm., 1995) and two mixtures (R32/R134a and R32/R125) have recently been intensively investigated ([lo- 121; C.D. 1995; D.R. Defibaugh and G. Morrison (NIST), pers. comm., 1995; Holcomb (NIST), pers. comm., M. Kleemiss, (Univ. of Hannover), pers. comm., 1995). This work includes a comprehensive set of experimental VLE results (about 480 data points in total) over the temperature range of about 240 to 320 K. The VLE of binary and ternary mixtures comprising these four principal fluids at low temperatures is a particularly important consideration for the engineer wishing to design the evaporator of a refrigeration system since they typically have bubble point temperatures in the range 223-233 K ( - 50 to - 40°C). Many of the reported VLE data sets for these mixtures do not include data at these low temperatures. In addition to reporting the experimental VLE data we also report the results of fitting the data to two thermodynamic models using a fitting program based upon Maximum Likelihood techniques [ 13,141. The models are based upon the Redlich-Kwong EOS (RK EOS) [ 151 with the data fitted by adjusting an expression for the excess Gibbs energy (Wilson parameters). The first model (Wilson/RK) uses the unmodified RK EOS to describe the vapour phase and the Wilson equation to describe the liquid nonideality. In the second model (Wilson/RK/MHV-2) the RK EOS is modified to incorporate an MHV-2 mixing rule with the excess Gibbs energies matched to those of the Wilson activity coefficient model at zero pressure [ 161. This work presents the results of fitting these data to these models with some comparisons of the quality of the fits and a tabulation of the fitting coefficients. We also report calculated azeotropic compositions as a function of temperature for three mixtures (R32/R125, R32/R143a, and R125/R143a) and the approximate composition of a ternary saddle point azeotrope as predicted by the Wilson/RK model.

M.H. Barlq

2. Experimental

et ol./

Fluid Phosr Equilibritr

I40 CIYY71 IX-206

IX5

method

All refrigerant samples were analysed prior to use by gas chromatography. Samples were better than 99.5% pure with less than 20 ppm of water. The major impurity in R134a was the isomer R134 (CF, H-CF, H, about 600 ppm); in R125 it was RI 15 (C,ClF,l, and in R32 it was CFH,. The experimental method used in this work to collect the raw VLE data is a static method. The experimental measurements are: pressure, temperature, the mass of each refrigerant in the fixed volume apparatus and the temperature of the apparatus. This is an accepted method of measuring VLE data [ 171 in which the liquid compositions are calculated as part of the data reduction process. A schematic diagram of the apparatus is shown in Fig. 1. The rig was constructed from l/4 in. (0.635 cm) 316 stainless steel pipework, and standard stainless steel ‘Whitey’ bombs and valves. It consisted of a filling manifold, where pure refrigerant materials were thoroughly degassed and weighed prior to transfer via vacuum distillation to the tempemture-controlled measuring manifold (Fig. 1). Rig operation and data collection were automated allowing the rig to run and collect data overnight. The measuring manifold, consisting of a sample vessel (C in Fig. 11, pipework and transducer, is of an accurately known volume. The transducer and pipework are maintained at least 20 K above the maximum temperature to which the sample is to be subjected to avoid condensation. The transducer is typically contained within a circulating air oven whilst the pipework connecting it to the sample vessel is lagged with heating tape. The MKS 25000 Baratron pressure transducer is calibrated using a Degranges et Huot DH21000 Force Transducer under these conditions. The temperature of the transducer and the connecting pipework are measured by three thermocouples (one on the transducer (t 11, the other two on different parts of the connecting pipework (t2, t3). The attached sample vessel (C) is immersed in a glycol/water bath whose temperature can be maintained to within i 0.15 K anywhere within the

Fig. I. A schematic diagram of the apparatus used to obtain the pure component of the refrigerant mixtures.

Lapour pressures and the vapour pressures

Abbreviations: A: Sample reservoir; B: Weighing bomb; C: Sample vessel: PRT: Calibrated platinum resistance thermometer attached to the wall of C; t I, t2, t3: Thermocouples monitoring the temperature of the apparatus containing sample vapour

186

M.H. Barley et al./ Fluid Phase Equilibria 140 (19971 183-206

range 233 to 333 K. The contents of C are stirred by a magnetic follower and the sample temperature is continuously monitored by a platinum resistance thermometer (PRT), positioned in or at the vessel wall and previously calibrated against an ASL F25 standard PRT probe. The sample reservoir (A) on the filling manifold contains a pure sample of the first refrigerant. This material is degassed using repeated freeze-pump-thaw cycles under vacuum until all volatiles/inerts have been removed and no vapour pressure at liquid nitrogen temperatures can be detected by the Pirani gauge. A sample of this degassed refrigerant is then vacuum distilled into a weighing bomb (B) which is detached and weighed. The refrigerant is then vacuum distilled from (B) into the sample vessel (C) on the measurement manifold. Reweighing bomb (B) gives an accurate mass of the refrigerant transferred to the equilibrium vessel. The procedure is repeated for the second refrigerant. Once the required refrigerant mixture has been distilled into the sample vessel then the run can start at the lowest temperature (about 233 K). The whole system is allowed to equilibrate while the temperature and pressure are both continuously monitored by the computer. Equilibrium has been reached when both the sample vessel temperature (to within 0.15 K) and the measured pressure (to within f0.02 bar) have been stable for at least 20 min. Once this condition has been reached the computer monitors the measured pressure and temperature for a further 20 min and providing neither value strays beyond the above limits the average temperature and pressure is recorded as a data point. If either measurement strays outside the above limits then the system is re-equilibrated for a further 20 min. Once a data point has been obtained the temperature of the glycol/water bath is then raised to a new value whereupon the equilibrium procedure is repeated prior to obtaining a further data point. In this way over a period of 8- 16 h between 6 and 8 data points (P and T) are collected (typical temperature range 233 to 333 K) for a sample of known initial composition. The procedure is then repeated for refrigerant mixtures of different compositions (typically 3 to 5 mixtures). The two experimental pure component vapour pressures and raw VLE data recorded using the above methods are reported in Appendix A. The experimentally observed refrigerant masses, pressures and sample vessel temperatures are given. Measurements of pressure, temperature and mass were obtained according to NAMAS procedures and the temperatures are quoted on the ITS90 scale. The temperature and pressure data in Appendix A are given after correction for the calibration of the transducer and the platinum resistance thermometer. All quoted pressures are absolute.

3. Experimental errors Error in measured pressure ( P) = + 0.005 bar f 0.1% of observed pressure. Error in measured temperature (T) = + 0.2 K for T < 273K and + 0.1 K for T > 273K. Error in vapour temperature: Vapour temperature is measured by the two or three thermocouples (tl, t2, t3) placed upon the manifold and transducer. The transducer contains the bulk of the vapour volume and is kept at a constant temperature (typically close to 390 K) so tl will probably have an error < 0.5 K. However the other thermocouples (t2 and t3) are difficult to position and their estimated error is f4 K. The vapour temperature quoted in Appendix A is the volume weighted average of these temperatures with an overall estimated error of +2 K. By varying the vapour temperature for a given data set and repeating the fit we were able to confirm that a change in vapour temperature of several degrees has a negligible effect on the quality of the data fit.

M. H. Barley et al. /Fluid

Phase Equilibriu

LX7

140 (I 997) 18%206

Error in refrigerant masses (M,, M,): The error in each individual weighing is estimated to be f0.002 g. Hence for the first binary reported in Appendix A the weight of R32 placed in the sample vessel would be 8.793 + 0.002fi g as this is derived from a difference of two absolute weighings. However, mixtures for this binary (R32/R134a) are made by the sequential addition of small samples of the second refrigerant thereby giving an associated accumulation of errors. Thus the first mixture would contain 1.758 f 0.002fi g of R134a but the second mixture would contain 5.747 + 0.002& g of R134a (four absolute weighings). Statistical methods showed that it was important to account for this accumulation of errors in the refrigerant masses when fitting the data. Error in the total volume of the apparatus: This is not an experimental error but a constant value was used in the fits. The total volume of the apparatus is known to within 0.1 ml and is reported in Appendix A for each data set.

4. Results and correlation

of the data

In Table 1 some pure component data for the four refrigerants are presented. The data are taken from the literature at the time this work was undertaken and the sources are referenced within the table. The critical properties are used in the EOS model, the molar volumes are nominal values

Table I Properties

of the pure refrigerants

used in this work

R134a Molecular weight Th (K)

102.13 247. I 374.2 [ 181 40.55 [181

T, (K) PC(Bar) Vupour

prtwures

B D E

densit)! (kg m

A B

c D E Vm (cm3 mole-

RI25 52.024 221.4 351.6 [19] 58.16 [19]

120.02 225. I 339.4 [I91 36.43 [ 191

Rl43a 84.00 225.8 346.0 [20] 37.9 [20]

(bar)

A

Liquid

R32

’)

[2 1,221

1231 92.681 I3 -4461.955 0.025 169 - 14.46098

l24.251 129.4138 - 5383.943 0.036158 - 20.9763 1

1201

137.7291 -6152.14 0.0343508 -22.01648

[26,271

t231

[28,291

[20.301

94.6649 - 4459.92 0.025856 - 14.8973

-iJ

509.15 902.402 637.097 -504.31 415.766 82.5

429.76 610.355 1372.037 - 1962.64 1370.34 I 48.39

57 I .o 979.403 768.93 I - 750.58 607.143 107.54

45.5.0 71X.382 0.0 I I 13.79 - 75 I.576 87.05

The vapour pressure (P) correlations are of the form ln( P) = A + B/T + DT + /Z%(T) where the temperature (T) is in K. The liquid density correlation is of the form DENSITY = A + BX + CX* + DX3 + EX4 where X = (I - T/q.,“’ with T and T, in K.

188

M.H. Barley et al. / Fluid Phuse Equilibria 140 (1997) 183-206

-2

L----&-__.-.a

220

240

260

300

320

340

TEMPERATURE (K)

Fig. 2. A comparison of the pure component vapour pressures as measured in this work (see Appendix) and the calculated pressures from the standard vapour pressure correlations in Table 1. % Error in vapour pressure = (P(exp)P(calc)). lOO/ P(exp).

(obtained by dividing the molecular weight by a calculated liquid density at 298.15 K) used in fitting the liquid non-ideality to the Wilson equation. In some cases the vapour pressure and liquid density coefficients presented here have been obtained by reconciling several sources of data available at the time. In these cases only literature sources having a dominant effect on the correlation coefficients are listed in the footnotes. The ‘standard’ vapour pressure correlations in Table 1 were not used in the correlation of the VLE data and are presented here because they were used later to generate phase envelopes (see below), to calculate azeotropic compositions, and for comparison against the pure component vapour pressures for each mixture (see Fig. 2). The correlation for R32 is based upon the rather old data of Malbrunot [23] but is in excellent agreement (within 0.5%) with the more recent experimental data from NIST (e.g., Defibaugh et al. [31]; Weber and Goodwin [32]). In a similar way the correlation for R143a is in good agreement with the experimental data of Fukushima [33]. Normal boiling points for each refrigerant were calculated from the vapour pressure curves. As a check on both the purity of the refrigerant samples and the accuracy of our vapour pressure measurements the pure component vapour pressures reported in Appendix A were compared to the standard correlations from Table 1 (see Fig. 2). For all but one sample of R125 the experimental vapour pressures agreed with the correlations above 233 K to within 1%. The problems of obtaining pure samples of R125 for physical property measurements have been discussed in the literature [25]. In our work the R125 sample used for the R32/R125 and R125/R134a was known to be of very high purity but the sample used in the R125/R143a mixture came from a different source and may have had a small degree of contamination, No attempt was made to correct the data for the presence of any impurities.

5. Methods used to correlate the data The chosen thermodynamic models were fitted to each of the binary mixture VLE data sets using Maximum Likelihood techniques [ 13,141. The use of these methods to correlate thermodynamic data for pure refrigerants have been described elsewhere [34,35]. The experimental apparatus is modelled via a constant volume, temperature flash calculation where we have at least two zones in the apparatus at different temperatures but the same pressure. All the liquid present is at the temperature

M.H. Barley et al. /Fluid

Phase Equilibrist

140 (1997)

IX-206

I x9

measured by the PRT. The remaining volume in the equilibrium vessel contains vapour and is also at this temperature and the pipework and transducer are at the temperatures measured by the thermocouples tl, t2 and t3 (see Fig. 1). The volume occupied by the liquid mixture was calculated using simple mixing rules [36], p. 75. In this work the volume of the apparatus was assumed to be exact throughout the experiment. All other measured quantities (temperature (T), pressure (PI and masses (M) of each refrigerant) were assumed to have associated errors. When using the method of Maximum Likelihood the objective function (OF in Eq. (1)) is minimised by moving the co-ordinates of an experimental point a minimum amount in multi-dimensional space such that the thermodynamic model is satisfied.

OF=Lljr;;l-/l+[!$r+ L”;;* 1’1

(1)

In Eq. (I) T, P, M are the experimental values while T *. P . M * are the associated model values; up, CT~ and CT, are the standard deviations associated with pressure, temperature and refrigerant mass measurements. In this technique the objective function is the sum over all the experimental points of the deviation of the experimental point from the prediction of the model (i.c., T-T * , P-P * etc.) normalised by the appropriate standard deviation that is derived for each measured property from the accuracy of the experimental measurements. The parameters of the model are then varied until the objective function is minimised. The VLE data were correlated to two models: In the first thermodynamic model the vapour phase fugacity was represented by the RK EOS [ 151 and the liquid phase fugacity (via activity coefficients) by the Wilson equation. This equation is a thermodynamically consistent representation of the excess Gibbs energy of the liquid phase as a function of composition and temperature ([37] section 4.9). For this work a temperature dependent form of the Wilson equation was used in which the usual Wilson energy term is replaced by a simple expression linear in temperature: ln( ‘I,?) = ln( VI/V,)

- (~5;~ + Ey2T)/RT

(2)

- E,,/RT

(-9

rather than the usual form: 1n( *I,?) = ln(Vz/V,)

In Eqs. (2) and (3) A,* is one of the Wilson parameters for an isotherm; V,, V, are the nominal molar volumes for components 1 and 2 respectively (see Table I) and E,?, Ei2, and Ey2 are Wilson Energy values. The possible use of four adjustable parameters allows an improved fit to VLE data over a wider range of temperatures (typically 230 to 330 K) and pressures. For the VLE experiment the vapour fugacity must equal the liquid fugacity for each component at every experimental point. This is represented by Eq. (4) ([37], p. 167): .v,4, P = y; P,~x;@;(PNTcorr)

(4)

where x,, yi, 4; are the liquid mole fraction, the vapour mole fraction, and the vapour fugacity coefficient respectively of component i in the mixture. PO and CD, are the vapour pressure and the fugacity coefficient for saturated pure component i at temperature T respectively. P is the system pressure, (PNTcoIx~ is the Poynting correction and yi is the activity coefficient of component i in the liquid mixture. The fugacity coefficients are calculated from the EOS and the Poynting correction from the liquid density correlation. The coefficients in the correlation for the pure component vapour pressures are included in the set of parameters to be fitted. This requires that experimental pure

190

M.H. Barley et al./ Fluid Phase Equilibria 140 (1997) 183-206

Table 2 The Wilson coefficients

and binary interaction

W/RK R32/

parameters

mixtures W/RK-MHV-2

W/RK-MHV-2

W/RK

1406.101 - 3232.64 - 6.29320 16.26376 1.52

Ei2 = 4401.365

R32/R143a

R134a

E;* = 1712.828 E;, = - 3492.24

,?$ = - 6.768006 Ey;, = 15.69194 QFIT 1.35 R32/R125

E;, = - 2269.478 Ey2 = -9.34447 E;, = 3.74354 QFIT 1.4 1 R125/

E& = - 358.5024

- 925.0552 2179.257 7.266844 - 9.20768 1.18

E;, = 2110.644

l?;* = 4.7 17848 E”;, = - 8.35962 QFIT 0.95 R125/

for the six refrigerant

R134a

E;2 = 551.2677

6532.03 -3184.54 - 18.8157 9.247696 1.49

E;, = - 1167.539 Ey2 = - 4.62098 ,!?;, = 7.303572 QFIT 1.25

3692.472 - 1996.209 - 6.27389 2.5 1644

1.62

R143a

Ei2 = - 544.03 1 E;, = - 359.508 E’,f2 = 1.52811 E;, = 1.17455

QFIT 0.62 R143a/R134a Ei2 = 29.14804 E;, = - 141.0725 E& = -3.17771 E;, = 4.214665 QFIT 1.23

1247.905 -915.013 0 0 0.97

499.2832 5.987741 -0.850391 - 0.799846 1.34

A,,=(I:/~)~x~[-(E~~+~~T)/RT], R=8.31451 Pam3 K-’ mole ‘. W/RK refers to liquid fugacity described by the Wilson model with the vapour fugacity described by the Redlich-Kwong EOS. W/RK-MHV-2 refers to an RK EOS model using the MHV-2 mixing rule and the Wilson equation to describe the excess free energy. QFIT is the quality of fit parameter referred to in the text.

component vapour pressure data for the specific samples used to make the mixtures are included in the data set to be fitted. This eliminates many potential systematic errors associated with sample purity, the equipment, operator etc. The fitting program determines the activity coefficients that best fit the experimental pressure at the observed temperature for each experimental point and are thermodynamically consistent with the Wilson/RK equation across the range of compositions and temperatures. To ensure this, the calculated liquid phase activity coefficients ( ri) for both components (1 and 2) must obey the following equations: Along any isotherm for a binary mixture: ln( rr) = - In( x1 + AI2 x2) + Bx, ln(y,)

= -ln(A,,x,

B=

42 + 42x2

-Bx,

+x2)

(5) (6)

where

Xl

A21 A21-5

(7)

+x2

and: 42

=

(V2/V,kXP[

-

4,

=

wV2kxP[

- 6

Pi2

+

J?,W~~]

+ J%w~~I

(8)

(9)

M.H. Barley et al./ Fluid Phase Equilibria 140 (lY47) 183-206

141

where xi, Vi have the definitions explained above, A,*, AZ, are the Wilson parameters for the appropriate isotherm, Ei2, E;12, E;,, and Ei, are the Wilson Energy values for the temperature dependent Wilson correlation describing the binary mixture (Eq. (2)). The Wilson coefficients for all six binary mixtures are presented in Table 2. Using similar methods, and with some modification of Eq. (41, the VLE data can be fitted to the Wilson/RK/MHV-2 model. The mathematical basis of this model and its use in modelling refrigerant mixtures has already been described [4]. In this work the VLE data were fitted directly to the Wilson/RK/MHV-2 model. Mathias-Copeman parameters [38] were determined by including the experimental pure component vapour pressures in the fitting exercise as described above for the Wilson/RK model. The RK/MHV-2 EOS was used to describe the fugacity in both phases. This ensured consistency of activity coefficients near the critical point when the resulting Wilson parameters were used in a Wilson/RK/MHV-2 model to generate thermodynamic property predictions for a specific mixture. In our earlier work [4] Wilson coefficients obtained by fitting VLE data to a Wilson/RK model were adapted for use in a Wilson/RK/MHV-2 model with some small inconsistencies in the model. This was originally done to get a useful model developed and in-use as quickly as possible and the small inconsistencies were of minimal practical significance. However this does mean that the Wilson coefficients (for the Wilson/RK/MHV-2 model) quoted in Table 2 are different from those reported in our earlier paper and the values quoted in this paper should be used in preference.

6. The quality of the fits The quality of the deviation between the deviation derived from a normal distribution

fits can be presented using statistical plots of the type seen in Fig. 3. The model and each experimental data point is shown relative to the standard the expected errors in the measurement of temperature and pressure. Assuming of errors and a reasonable assessment of the likely experimental errors in

LJ-31

0

10

5

3’ -40

I

20

15 PREDICTED

PRESSURE

25

30

35

(bar)

I -20

0 TEMPERATURE

20

40

60

(DEG. C)

Fig. 3. An analysis of the residual errors in the fit of the R32/R125 experimental data to the RK/Wilson model: (a) the distribution of pressure errors, (b) the distribution of temperature errors. The different symbols refer to different mixture compositions and include the pure component vapour pressures.

M.H. Barley et al. / Fluid Phase Equilibria 140 (1997) 183-206

192

measuring temperature and pressure then greater than 99% of the experimental values should be within +3 standard deviations of the values predicted by the model and no systematic trends should be apparent. In Table 2 the best fit for each binary mixture to each of the three models is given along with a parameter (QFIT) indicating the quality of the fit. QFIT relates the sum-of-the-squares of the normalised errors (see Eq. (1)) to the number of degrees of freedom in the fit. The number of degrees of freedom (NDF) equals the number of independent data points minus the sum of the number of parameters to be fitted and the number of independent constraints in the thermodynamic model. This is described mathematically by:

QFIT =

NDF

(10)

From statistical theory the ideal value for QFIT is unity. A value between 1 and 1.5 is considered a good fit and a value below 1.7 suggested an adequate fit. A QFIT value below unity indicates that the model is overfitting the data; that is, some of the features of the errors are being fitted to the model. A more exact inference can be gained from x’ tables but as the estimated experimental errors are not themselves accurate a more sophisticated analysis was not tried. As can be seen from Table 2 good fits were obtained using both the Wilson/RK and the Wilson/RK/MHV-2 models. All six sets of VLE data were fitted to the two models with QFIT values below 1.7 suggesting that almost all the data in each set are represented by the model to within experimental error. In Table 2 it is apparent that the QFIT value for the Wilson/RK model is consistently smaller than that for the Wilson/RK/MHV-2 model suggesting that the former model is slightly more flexible in representing VLE data. The statistical plots of the type shown in Fig. 3 provide information on the experimental errors used in the fits. Providing the model is sufficiently flexible to fit the data then if the experimental errors are underestimated many of the residual errors will be greater than + 3 standard deviations but the distribution of the errors will still be random. If the experimental errors have been overestimated then the errors will be grouped around the x-axis well within +3 standard deviations: in an extreme case of overestimated experimental errors the points will be almost coincident with the x-axis in graphs such as Fig. 3. The graphs obtained during this work suggest we have been about right in assessing our temperature and pressure errors (perhaps they have been very slightly overestimated). In all these fits the errors in the refrigerant masses were distributed within 3 standard deviations. With so few data points per binary VLE set (5-7 mixtures) it is impossible to comment on their degree of randomness but the distribution of errors was consistent with the derived experimental errors for refrigerant mass described above.

7. Comparisons

with literature

data

7.1. R32/R125 Six data sets (either PTx or PTxy data) on the VLE of this binary were available at the time of writing. A Wilson/RK model using pure component properties from Table 1 and the Wilson

M.H. Barley et al./ Fluid Phase

Eyuilihritr

IJO (lYY7J

IX--206

19:

coefficients from Table 2 was used to predict bubble point pressures and vapour compositions for the experimental temperatures and liquid compositions. The predicted pressures were then compared to the experimental pressures by expressing the difference as a percentage error as shown in Fig. 4. As can be seen from the graph our predictions lie quite centrally among the scatter of available experimental data. We show excellent agreement (within + 1%) with the data from NIST (set 31, and the data from the labs of Watanabe (set 41. Only two data sets (sets 1 and 3) provided experimental vapour compositions for comparison with the model values. Most of the experimental points agreed with our model to within 0.03 mole fraction (Fig. 4b) although a few of the points showed disagreement up to 0.07 mole fraction. 7.2. R32 / R134a Five data sets for the VLE of this binary were available at the time of writing. Similar graphs to those described above were done for this mixture (see Fig. 5). The experimental data for this binary were quite scattered (to within 15%) around the vapour pressures predicted by our model. However our model was in good agreement with some of the NIST data (set 21, and also the data from Kleemis (set 6). When the errors in vapour composition (_v(exp) - y(calc) for component 1 in mole fractions) were analysed in the same way then most of the data were within kO.02 in mole fraction from our predicted values, the main exception being some of the data of set 3. 7.3. R125/

R134u

For this binary only one set of data was available (Y. Higashi, pers. comm., 1995). Comparisons between these data and our model showed good agreement on vapour compositions even if our model ??

10

20

PREDICTED

30

PRESSURE

40

(bar)

o.‘7--pp

1

---

,J

-0.1 i__-_;;;_ 10 5

20 PREDICTED

PRESSURE

25

30

35

(bar)

Fig. 4. A comparison between the predictions of our Wilson/RK model for R32/R 125 and the experimental data of various authors: (I ) Fujiwara et al. [7], (2) D.R. Defibaugh and G. Morrison (NIST), pers. comm.. 1995. (3) C.D. Holcomb (NISI?, pers. comm., 1995, (4) Widiatmo et al. [12], (5) Higashi [IO], (6) M. Kleemiss (Univ. of Hannover). pers. comm.. 1995. (a) Bubble point pressures with o/c error in bubble point pressure = (P(exp)P(calc)). lOO/P(expl, (b) vapour composition. The data sets are identified by the key in (a>. The Wilson/RK model used the appropriate data from Tables I and 2.

194

M.H. Barley et al. /Fluid Phase Equilibria 140 (1997) 183-206

IO

20 PREDICTED

-0.05

IO

0

(bar)

30

20

PREDICTED

Fig. 5. A comparison between various authors: (1) Fujiwara et (NIST), pers. comm., 1995, (4) pressures with % error in bubble model used the appropriate data

30

PRESSURE

PRESSURE

40

(bar)

the predictions of our Wilson/RK model for R32/R134a and the experimental data of al. [7], (2) D.R. Defibaugh and G. Morrison (NET), pers. comm., 1995, (3) C.D. Holcomb Higashi [ill, (5) M. Kl eemiss (Univ. of Hannover), pers. comm., 1995. (a) Bubble point point pressure = (P(exp)P(calc)). lOO/P(exp), (b) vapour composition. The Wilson/RK from Table 1 and 2.

was consistently predicting more R125 in the vapour than was experimentally observed. agreement for pressures was not so good with a couple of points disagreeing to between 510% Fig. 6).

The (see

7.4. R143a/R134a A set of PTq data has been reported in the literature [8]. Our model shows reasonable agreement in pressure with these experimental data. However the agreement between our predicted vapour compositions and those observed is surprisingly good (within f0.02 mole fraction - see Fig. 7).

E

0.01

a 0 0 5 2 -0.0,

E 0.02 10 PREDICTED

15 PRESSURE

20 (bar)

25

: w

-o.030

u=u==u

=

::

0

0 008 O 0 0” 0” ? ? 5 10 15 20 PREDICTED PRESSURE (bar)

25

Fig. 6. A comparison between the predictions of our Wilson/RK model and the experimental data for R125/R134a: Y. Higashi, pers comm., 1995. (a) bubble point pressures with % error in bubble point pressure = (P(exp)P(calc)). 1OO/ P(exp), (b) vapour composition.

M.H. Barley et al. /Fluid

Phase Equilibria 140 f 1997) 183-206

E

0.03

d

0.02

E ~

0.01

5

0 B 0: 4.01 z 4.02

10 15 20 25 30 PREDICTED PRESSURE (bar)

35

PREDICTED PRESSURE (bar)

Fig. 7. A comparison between the predictions of our Wilson/RK model and the experimental data for R143a/Rl34a (Kubota and Matsumoto [81X (a) Bubble point pressures with % error in bubble point pressure = (P(exp)Ptcalc)). IOO/P(exp), (b) vapour composition.

s > 32

s

0

0

0

8

LI

0

65-0.01 p 0.01 i_, 65

Fig. 8. A comparison between the predictions R32/R I43a (Fujiwara et al. [7]).

,n~ 0

0

-t-PREDICTED

7.5 8 PRESSURE (bar)

of our Wilson/RK

model and the experimental

vapour compositions

for

7.5. R32 / R143a For this mixture the only VLE data available were the single isotherm of Fujiwara [7]. Our model agrees with this data to within 3% in pressure, but agreement on vapour composition was much better (within 0.01 mole fraction - see Fig. 8). 7.6. R125/

R143a

A set of bubble-point pressures and liquid densities for this mixture are available in the literature [9]. These data show some scatter in bubble-points (especially along the 280K isotherm) and our model agrees with these pressures to within 2.5% (see Fig. 9). No vapour compositions were reported.

15 PREDICTED

20 PRESSURE

25 (bar)

Fig. 9. A comparison between the predicted bubble point pressures of our Wilson/RK model and experimental RI 25/R143a (Widiatmo et al. [9]). % Error in bubble point pressure = ( P(exp) - P(calc)). lOOP(exp).

data for

196 8.

The predictions

M.H. Barley et al. /Fluid Phase Equilibria 140 (1997) 183-206

of the model

The models can be used to predict the behaviour of the six binary mixtures. Thus the Wilson/RK model was used to generate the phase envelopes seen in Fig. 10. Of the six mixtures, three (R32/R125, R125/R134a and R143a/R134a) are very close to ideal and hence the phase envelopes for R32/R134a and R143a/R134a are very similar in shape to the phase envelopes for R125/R134a seen in Fig. 10. The other three mixtures were much less ideal and due to the close boiling points of R32, R125 and R143a, all of them showed azeotropic behaviour for all or part of the temperature range studied here. The Wilson/RK model can be extrapolated to lower temperatures to estimate the normal boiling point of the azeotropic mixtures as well as predicting the compositions and associated vapour pressures of the azeotropic mixtures over a range of temperatures. The results of these calculations are seen in Table 3. As can be seen from Fig. 10, of the three non-ideal mixtures two (R32/R125 and R32/R143a) show positive deviations from ideality while the third (R125/R143a) shows negative deviations. This makes the ternary mixture (R32/R125/R143a) of great interest and the Wilson/RK model does predict a saddle point azeotrope at about 84 mole% R32, 4.5 mole% R125 and 11.5 mole% R143a.

R125-R134A

R32-R125ATOC 8.5,~

0

0.2

0.4

0.6

0.8

6.5 0

1

X(R125) R32-R143AAT4OC

0.2 0.4 0.6 0.8 X(R32)

R125-R143AAT-20C 3.4,-

1.8

d 2

3.3

; 3.25 I F

3.2

3.15

1.3 0

0.2

0.4

0.6

0.8

1

3.10

0.2 0.4 0.6 0.8 X(R125)

Fig. 10. Phase envelopes for four mixtures: The bubble-point and dew-point pressures were calculated using the Wilson/RK model and the appropriate data from Tables I and 2. The phase envelopes for R32/R134a and R143a/R134a are very similar to those illustrated for R125/R134a.

M.H. Barley et al./ Fluid Phase Equilibria

Table 3 Compositions

and vapour pressures of the azeotropic Rl25/Rl43a

140 (19971 IX-206

197

mixtures R32/R143a

R32/R 125

T(K)

P (bar)

X CR1251

P (bar)

X (R32)

P(bar)

X (R32)

223.15 233.15 243.15 253.15 263.15 273.15 2X3.15 293.15 303. IS 313.15

0.864 I .3X6 2.124 3.131 4.466 6.1X9 X.366 I 1.066

0.392 0.343 0.295 0.247 0.19x 0.146 0.087 0.019

1.116 I.790 2.749 4.066 5.X26 X.122

0.792 0.826 0.861 0.89X 0.936 0.978

I.125 I .X04 2.76X 4.094 1 5.856 8.153 I I.082 14.753 19.288

0.83 I 0.840 0.853 0.X69 0.88X 0.90X 0.930 O.YS4 0.97x

_

_

Predicted normal boiling points of the azeotropes: R125/R143a at 226.45 K ( - 46.7”C) and 0.376 mole fraction Rl2S. R32/Rl43a at 22 I .25 K t - 5 I .9”C) and 0.782 mole fraction R32. R32/R125 at 221.05 K ( - 52.I”C) and 0.X29 mole fraction R32. Mixture compositions and pressures calculated using the Wilson/RK model with the pure component data given in Table I and the interaction coefficients given in Table 2.

The calculations to find this azeotrope were done at 221 K which corresponded to a pressure of about 1 bar. No attempt was made to find how the azeotrope changed in composition as the temperature was raised.

9. Conclusion In this paper we present extensive VLE data for six refrigerant mixtures. The data for each mixture cover several compositions and an extended temperature range (usually 230 to 330 K). The experimental data for all the mixtures were analysed and fitted to two models derived from the Wilson equation. In each case the data were consistent with these models to within experimental error and the appropriate coefficients have been presented. Comparisons of the predictions of one of these models with literature experimental data showed good general agreement, particularly for the more intensively studied mixtures (R32/R134a and R32/R125). Using the thermodynamic models, phase envelopes and azeotropic behaviour could be predicted over a range of temperatures and pressures.

10. List of symbols E

M P P” R

Wilson energy values experimental mass of refrigerant experimental pressure pure component vapour pressure (calculated) gas constant

M.H. Barley et al./ Fluid Phase Equilibria 140 (1997) 183-206

198

T V

experimental temperature molar volume liquid mole fraction vapour mole fraction

X

Y Greek letters

activity coefficient vapour fugacity coefficient of a mixture component vapour fugacity coefficient of a pure component Wilson parameter for an isotherm standard deviation

S @ A (T Subscripts i j 1, 2

ith component jth component component identification

Superscripts *

value calculated

by a model

Acknowledgements The contributions

of colleagues

within ICI are gratefully

acknowledged.

Appendix A A. 1. Data for CF, Hz / CF,CFH, Manifold and transducer Manifold temperature (K)

(R32 / RI 34a)

volume (ml) = 22.65; Sample vessel (S/V)

s/v T(K)

cpbar,

235.68 242.69 253.06 263.28 273.54 283.61 293.84 313.96 323.06

&n-j

1.9807

2.6894 4.0346 5.8333 8.1998 11.1658 14.943 1 25.0907 31.0731

235.74 242.77 253.01 263.23 273.52 283.60 293.83 313.96 323.06

s/v T (K)

cpbar,

8.793 g R32, 1.758 g R134a

8.793 g R32

8.793 g R32 378 378 378 378 379 379 380 382 383

S/V T(K)

volume (ml) = 16.99.

1.9820 2.6947 4.0275 5.8232 8.1944 11.1653 14.9430 25.0915 3 1.0700

242.83 244.24 252.83 263.09 273.3 1 283.39 293.6 313.72 323.73

(2.43 17) (2.5633) 3.8100 5.5267 7.7493 10.5456 14.0877 23.6232 29.8349

M.H. Barley et al./ Fluid Phase Equilibria 140 (1997) 183-206

8.793 g R32, 1.758 g R134a 377 378 378 378 379 380 381 382

244.18 252.86 263.05 273.28 283.38 293.59 313.71 323.72

8.793 g R32, 5.747 g R134a (2.5589) 3.8183 5.5206 7.7389 10.5393 14.0840 23.6232 29.8315

4.554 g R32, 8.869 g R134a 376 376 377 378 379 380 380 382 383

233.56 239.37 25 1.08 262.59 273.24 283.32 293.55 313.58 323.72

230.9 1 238.98 252.44 262.99 273.22 283.30 293.52 313.61 323.72

1.1578 1.5375 2.5309 3.8455 5.4950 7.5257 10.1021 16.9382 2 1.4707

(2.2501) 3.348 1 4.8867 6.8611 9.348 I 12.5043 20.9976 26.5199

233.16 239.47 250.76 262.73 273.22 283.27 293.54 313.58 323.72

0.7711 1.0981 1.9652 2.9170 4.1722 5.7689 7.8096 13.3028 16.9562

235.91 243.88 252.8 263.05 273.29 283.36 293.57 313.59 323.72

1.1357 1.5388 2.4355 3.8557 5.4896 7.5170 10.0950 16.9354 2 1.4707

273.5 283.59 293.81

(2.2487) 3.3623 4.8789 6.8534 9.3410 12.4999 20.994 I 26.5 108

23 1.68 239.01 252.07 263.02 273.25 283.3 1 293.53 313.6 323.72

0.8105 1.097 1 I .936 1 2.9195 4.1769 5.771 7.8 130 13.2975 16.9587

0.530 g R32, 8.869 g R134a (0.5718) (0.8784) 1.5413 2.3754 3.4444 4.8216 6.6047 11.4522 14.7354

8.869 g R134a 0.5066 0.8308 1.3165 2.0011

244.13 252.76 263.05 273.28 283.35 293.54 313.72 323.73 1.504 g R32, 8.869 g R134a

0.530 g R32, 8.869 g R134a

8.869 g R134a 232.76 242.79 253.01 263.22

8.793 g R32, 5.747 g RI 34a

4.554 g R32, 8.869 g R134a

1.504 g R32, 8.869 g R134a 376 377 377 378 378 379 379 381 381

244.05 252.73 263.08 273.3 1 283.37 293.55 313.72 323.74

199

235.96 243.99 252.83 263.00 273.25 283.35 293.56 313.59 323.72

(0.5717) (0.8781) 1.5442 2.3709 3.440 1 4.8204 6.6043 1 I .4498 14.7373

8.869 g Rl34a 2.9462 4.1826 5.7979

313.84 323.98 333.99

10.2734 13.3332 16.9595

200

M.H. Barley et al./ Fluid Phase Equilibria 140 (1997) 183-206

A.2. Data for CF, H2 / C, F5 H CR32 / RI51 Manifold and transducer Manifold temperature (K)

volume (ml> = 22.65; Sample vessel (S/V)

s/v T (K)

cpbar,

8.02 g R32 378 377 378 378 379 379 380 382 382

235.68 242.69 253.06 263.28 273.54 283.61 293.84 313.96 323.06

1.9807 2.6894 4.0346 5.8333 8.1998 11.1658 14.943 I 25.0907 31.0731

8.02 g R32, 2.022 g R125 378 377 378 379 379 380 380 382 383

377 378 378 378 378 379 379 379 381 381

237.06 242.38 253.02 263.18 273.5 1 283.6 293.83 313.96 323.06

s/v T (K)

s/v T (K)

8.02 g R32

8.02 g R32, 2.022 g R125

235.74 242.77 253.01 263.23 273.52 283.6 293.83 313.96 323.06

1.982 2.6947 4.0275 5.8232 8.1944 11.1653 14.943 25.0915 3 1.07

238.4 240.57 253.11 263.28 273.55 283.62 293.84 313.97 323.07

237.22 242.25 253.07 263.28 273.54 283.61 293.84 313.96 323.06

Pbar1

2.1625 2.6879 4.0784 5.8779 8.2464 11.2139 14.9932 25.115 31.0619

8.02 g R32, 6.188 g R125

8.02 g R32, 6.188 g R125 2.1432 2.7052 4.0724 5.8575 8.2402 11.2134 14.9897 25.1131 3 1.0599

volume (ml> = 16.99.

2.2825 2.5073 4.0278 5.7997 8.148 11.0891 14.806 24.7826 30.6562

238.4 240.59 253.08 263.22 273.52 283.61 293.83 313.97 323.07

2.2823 2.5102 4.0217 5.7873 8.1461 11.0867 14.8043 24.78 1 30.654

0.479 g R32, 10.051 g R125

0.479 g R32, 10.051 g R125

1.457 g R32, 10.051 g R125

236.18 238.76 253.09 263.27 273.54 283.61 293.84

1.8841 2.0907 3.4959 5.0218 7.0266 9.5182 12.665 1

236.1 239.15 253.07 263.18 273.51 283.61 293.83

1.881 2.115 3.4947 5.0073 7.0206 9.5169 12.6617

313.83 323.06

20.9796 25.9348

313.83 323.06

20.978 1 25.9404

241.75 245.66 252.78 263.05 273.29 283.36 293.58 303.62 313.58 322.82

(2.1709) (2.5632) 3.6357 5.2427 7.3348 9.9506 13.2488 17.2204 2 1.9894 27.1976

M.H. Barley et al. / Fluid Phase Equilibria

1.457 g R32, 10.051 g RI25 379 380 380 380 380 380 381 382 382 381

241.77 245.56 252.79 262.96 273.28 283.35 293.57 303.61 3 13.58 322.8 1 10.051 g RI25

1.4937 2.28 11 3.3672

A.3. Data C,F,H/CF,CFH, Manifold and transducer Manifold temperature (K)

4.366 g R32, 10.051 g R25

241.36 243.56 252.77 262.97 273.27 283.39 293.6

(2.2936) (2.4832) 3.8712 5.5659 7.8114 10.6153 14.1503

241.36 243.33 252.78 262.96 273.27 283.35 293.59

(2.2737) (2.4746) 3.8735 5.5642 7.81 1 10.606 14.1493

313.71 322.8

23.6357 29.1925

313.71 322.79

23.6387 29.1927

10.051 g R12.5

10.051 g RI25

263.23 273.5 283.59

293.8 313.85 323.96

4.8339 6.7689 9.1817

12.2194 20.2624 25.5719

(R125/R134a) volume (ml) = 22.65; Sample vessel (S/V)

s/v T (K)

(Pbar)

14.13 g R125 379 379 379 379 380

20

183-206

4.366 g R32, 10.051 g R125 (2.1507) (2.5618) 3.6362 5.2255 7.3328 9.9490 13.2478 17.2178 21.9839 27.1882

232.79 242.94 253.04

140 (1997)

232.79 242.94 253.04 263.23 273.5

1.4937 2.28 11 3.3672 4.8339 6.7689

14.13 g R125

volume (ml) = 16.99. P (bar)

s/v T(K)

s/v T (K)

14.13 g R125, 1.275 g R134a

14.13 g R125, 4.036 g R134a

232.83 242.9 253.07 263.22 273.49

232.34 242.99 253.04 263.21 273.49

1.3949 2.1439 3.1802 4.5585 6.3864

1.1961 1.901 I 2.833 1 4.0924 5.7652

14.13 g R125, 4.036 g R134cr

14.13 g R125, 1.275 g Rl34a

381

283.59

9.1817

283.6 1

8.6719

283.57

7.8564

382 383 384

293.8 313.85 323.96

12.2194 20.2624 25.5719

293.82 313.82 323.95 333.95

11.55 19.1539 24.1894 30.0904

293.77 313.82 323.95 339.95

10.4947 17.5077 22.141 I 27.5483

I

202

M.H. Barley et al./Fluid

7.304 g R125, 6.305 g R134a 378 379 379 379 379 380 380 382 383 384

232.8 242.96 253.05 263.24 273.49 283.58 293.79 313.97 323.99 333.99

0.971 g R125, 9.46 g R134a

3.501 g R125, 9.46 g R134a 0.9795 1.5418 2.323 1 3.386 4.7972 6.589 8.8615 14.9889 18.9692 23.6499

232.89 242.99 253.01 263.21 273.44 283.54 293.78 313.97 323.99 333.99

0.7627 1.2045 1.8272 2.6958 3.8639 5.3589 7.2739 12.5053 15.9523 20.0353

0.5066 0.8308 1.3165 2.0011

273.5 283.59 293.8 1

232.84 242.98 253.04 263.22 273.48 283.59 293.83 313.97 323.99 333.98

0.6115 0.9814 1.5041 2.2467 3.269 4.5876 6.2943 11.0236 14.1757 17.934

9.36 g R134a

9.46 g R134a

9.44 g R134a 232.76 242.79 253.01 263.22

Phase Equilibria 140 (1997) 183-206

2.9462 4.1826 5.7979

313.84 323.98 333.99

10.2734 13.3332 16.9595

A.4. Data CF, Hz / CF, CH, (R32 / RI 43a) Manifold and transducer volume (ml) = 22.65; Sample vessel (S/V) Manifold temperature (K)

379 379 379 380 380 380 381 382 382

379 379 379

volume (ml) = 16.99.

s/v T (K)

s/v T (K)

S/V T (K)

5.762 g RI 43a

5.762 g R143a, 0.6500 g R32

5.672 g R143a, 1.183 g R32

234.61 242.42 252.47 262.66 272.83 282.91 293.14 313.27 323.33

1.5024 2.0845 3.0691 4.3983 6.1185 8.2816 11.0156 18.2818 23.0226

234.68 242.45 252.37 262.5 1 272.74 282.87 293.09 313.18 323.29

1.6666 2.3018 3.3735 4.8138 6.6844 9.0286 11.9619 19.6812 24.7013

235.61 242.46 252.35 262.59 272.67 282.88 293.11 313.24 323.3 1

cpbar,

1.8197 2.4150 3.5295 5.0603 7.0067 9.4915 12.5808 20.6842 25.8835

5.762 g R143a, 2.405 g R32

4.258 g R143a, 5.937 g R32

1.047 g R143a, 5.937 g R32

235.21 242.47 252.43

234.5 1 242.50 252.44

235.85 242.45 252.32

1.9242 2.6109 3.845 1

1.9016 2.6640 3.9222

2.0462 2.7043 3.9644

203

M.H. Barley et al./ Fluid Phase Equilibriu 140 (1997) IRS-206

380 380 380 381 382 382

262.63 272.7 282.92 293.10 313.26 323.3 I

5.5184 7.6585 10.4124 13.8289 22.9247 28.83 17

5.937 g R32 234.15 242.43 252.38

1.8583 2.6595 3.9506

Manifold and transducer

384 385 386 386 387 388 389 390 391

384 385 386 386 387 388 389 390 391

5.6383 7.8324 10.6654 14.2098 23.6565 29.8340

5.937 g R32

A.5. Data C,F,H/CF,CH,

Manifold temperature (K)

262.63 272.68 282.89 293.12 313.23 323.29

262.66 272.75 282.9 1

262.46 272.70 282.85 293.08 313.26 323.3 1

5.7089 8.0192 10.9501 14.6315 24.5360 3 1.0368

5.937 g R32 5.7338 8.0207 10.9704

293.12 313.27 323.34

14.6907 24.7132 31.3217

(R125/R143a) volume (ml) = 22.65; Sample vessel (S/V)

volume (ml) = 16.99.

s/v T (K)

s/v T (K)

s/v T (K)

6.0920 g RI25

6.920 g R125, 0.547 g R143a

6.920 g R125, 1.645 g R143a

253.65 263.56 273.51 283.55 293.67 303.40 313.72 323.91 334.11

255.13 265.08 275.04 285.13 295.25 305.05 315.39 325.56 335.63

252.52 262.43 272.42 282.59 292.53 302.54 312.49 322.66 332.79

3.3548 4.7718 6.6182 9.0028 11.9019 15.4794 19.7979 25.1111 29.9348

3.428 4.866 6.725 9.098 12.057 15.548 20.035 25.389 31.365

6.920 g R125, 4.803 g R143a

1.123 g R125, 7.012 g R143a

253.44 263.42 273.42 283.50 293.62 303.21 313.53 323.66 333.84

253.99 263.19 273.15 283.18 293.2 303.6 313.5 323.56 333.7

3.226 4.592 6.365 8.629 11.462 14.730 19.005 24.057 30.164

3.249 4.487 6.192 8.365 11.061 14.490 18.447 23.266 29.065

3.539 5.008 6.899 9.315 12.324 15.893 20.44 1 25.839 32.26

7.012 g R143a 252.80 262.67 272.64 282.77 293.06 302.78 312.71 322.77 332.91

3.1196 4.413 1 6.0953 8.2607 10.996 1 14.1671 18.0580 22.779 28.4658

204

M.H. Barley et al/Fluid

A.6. Data CF,CH,/

CF,CFH,

Manifold and transducer Manifold temperature (K)

365 365 366 366 366 366 367 367 367

(R143a/

R134a)

volume (ml) = 22.65; Sample vessel (S/V)

volume (ml) = 16.99.

S/V T (K)

S/V T (K)

s/v T (K)

10.246 g R143a

10.246 g R143a, 1.3925 g R134a

10.246 g R143a, 4.1494 g R134a

253.27 263.11 273.00 282.99 293.62 303.43 313.54 323.68 333.84

253.36 263.21 273.05 283.04 293.64 303.44 313.53 323.66 333.82

253.2 263.14 273.07 283.02 296.11 303.45 313.55 323.67 333.83

3.162 4.467 6.155 8.309 11.148 14.385 18.393 23.223 29.018

7.4788 g R143a, 9.0279 g R134a 365 365 366 366 366 366 367 367 367

Phase Equilibria 140 (1997) 183-206

253.18 263.08 273.05 283.04 293.58 303.46 3 13.55 323.68 333.84

2.240 3.216 4.510 6.161 8.342 10.887 14.039 17.839 22.398

2.4841 g R143a, 9.0279 g R134a 253.2 263.09 273.01 282.97 295.53 303.44 313.55 323.68 333.81

1.778 2.594 3.679 5.086 7.398 9.219 11.996 15.374 19.427

9.0279 g R134a

9.0279 g R134a 253.29 263.19 273.10

2.982 4.220 5.818 7.864 10.560 13.627 17.424 21.997 27.467

1.336 2.0073 2.9176

28.303 293.55 303.44

4.126 5.771 7.742

2.684 3.823 5.301 7.186 10.347 12.533 16.065 20.312 25.388

0.8294 g R143a, 8.9412 g R134a 253.2 263.08 273.02 283.02 293.65 303.43 3 13.53 323.67 333.83

1.504 2.227 3.206 4.494 6.247 8.289 10.879 14.044 17.879

9.0279 g R134a 313.52 323.65 333.82

10.226 13.279 16.998

S/V stands for sample or equilibrium vessel (C in Fig. 1). The manifold temperature is a volume weighted average of the readings of the two or more thermocouples placed on the transducer and connecting pipework.

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