NIATX INST. OF STAND & TECH R.I.C.

AlllDS 3SS3S3

NISI

PUBLICATIONS

fMisr United States Department of Commerce

Technology Administration

National Institute of Standards and Technology

NIST Technical Note 1397

Selected Thermodynamic Properties for

IVIixtures of R-32 (Difluoromethane),

R-125 (Pentafluoroethane), R-134A

(1 ,1 ,1 ,2-Tetrafluoroethane),

R-143A (1,1,1-Trifluoroethane),

R-41 (Fluorometliane), R-290 (Propane),

and R-744 (Carbon Dioxide)

QC

100

U5753

NO. 1397

1997

Cynthia D. Holcomb

Joseph W. Magee

Jennifer L. Scott

Stephanie L. Outcalt

William M. Haynes

NIST Technical Note 1397

Selected Thermodynamic Properties for

Mixtures of R-32 (Difluoromethane),

R-125 (Pentafluoroethane), R-134A

(1 ,1,1 ,2-Tetrafluoroethane),

R-143A (1,1J-Trifluoroethane),

R-41 (Fluoromethane), R-290 (Propane),

and R-744 (Carbon Dioxide)

Cynthia D. Holcomb

Joseph W. Magee

Jennifer L. Scott

Stephanie L. Outcalt

William M. Haynes

Physical and Chemical Properties Division

Chemical Science and Technology Laboratory

National Institute of Standards and Technology

325 Broadway

Boulder, Colorado 80303-3328

December 1997

f/ \

''^ATSS O*

U.S. DEPARTMENT OF COMMERCE, William M. Daley, Secretary

TECHNOLOGY ADMINISTRATION, Gcry R. Bachula, Acting Under Secretary for Technology

NATIONAL INSTITUTE OF STANDARDS AND TECHNOLOGY, Raymond G. Kammer, Director
National Institute of Standards and Technology Technical Note

Natl. Inst. Stand. Technol., Tech. Note 1397, 200 pages (December 1997)

CODEN:NTNOEF

U.S. GOVERNMENT PRINTING OFFICE

WASHINGTON: 1997

For sale by the Superintendent of Documents, U.S. Government Printing Office, Washington, DC 20402-9325
TABLE OF CONTENTS

Page

LIST OF FIGURES iv LIST OF TABLES vii L INTRODUCTION 1

2. SIGNIRCANT RESULTS 2

2.1 Thermophysical Properties of R-32/125/134a and Its Constituent Binaries 2

2.1.1 R-32/134a 3

2.1.2 R-32/125 4

2.1.3 R-125/134a 5

2.1.4 R-32/125/134a 5

2.2 Mixtures withR-143a 6

2.2.1 R-32/143a 6

2.2.2 R-125/143a 7

2.2.3 R-143a/134a 8 2.3 Mixtures with R-290 8 2.3.1 R-32/290 '. 9

2.3.2 R-125/290 10

2.3.3 R-134a/290 1

2.4 R-41 and R-41/744 11

2.4.1 R-41 12

2.4.2 R-41/744 13

3. REFERENCES 15

4. FIGURES 18

APPENDIX A. TABLES OF THERMOPHYSICAL PROPERTY DATA A-1

APPENDIX B. EXPERIMENTAL APPARATUS B-1

B.l Vapor-Liquid Equilibrium Apparatus B-2

B.2 Bubble-Point and Near-Saturation (p,pT) Apparatus B-2

B.3 Isochoric (p,pT) Apparatus B-3

B.4 Adiabatic Constant Volume Calorimeter , B-4

B.5 Spherical Resonator Speed of Sound Apparatus B-4

APPENDIX C. ESTIMATES OF STATE-POINT UNCERTAINTIES C-1

APPENDIX D. GAS CHROMATOGRAPH CALIBRATION PROCEDURES D

APPENDIX E. REPRINTS OF PAPERS RELATED TO THE LEMMON-JACOBSEN

MODEL [1] AND REFPROP 6.0 [2] E-1

111
LIST OF FIGURES

Page

Figure 1. Types of Phase Behavior for the Refrigerant Mixtures Studied in This Project.... 19

Figure 2. Comparison of Bubble-Point Pressures for the R-32/134a System; the Baseline

is from the Lemmon-Jacobsen Model in REFPROP 6.0. 20

Figure 3. Comparison of Vapor Compositions for the R-32/134a System; the Baseline

is from the Lemmon-Jacobsen Model in REFPROP 6.0. 21

Figure 4. Comparison of Densities for the R-32/134a System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0. 22

Figure 5. Comparison of Bubble-Point Pressures for the R-32/125 System; the Baseline

is from the Lemmon-Jacobsen Model in REFPROP 6.0. 23

Figure 6. Comparison of Vapor Compositions for the R-32/125 System; the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0 24

Figure 7. Comparison of Densities for the R-32/125 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.............. 25

Figure 8. Comparison of Bubble-Point Pressures for the R-125/134a System; the Baseline

is from the Lemmon-Jacobsen Model in REFPROP 6.0. 26

Figure 9. Comparison of Vapor Compositions for the R-125/134a System; the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0..... 27

Figure 10. Comparison of Densities for the R- 125/1 34a System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0....... 28

Figure 11. Comparison of Bubble-Point Pressures for the R-32/125/134a System; the Baseline

is from the Lemmon-Jacobsen Model in REFPROP 6.0... 29

Figure 12. Comparison of Vapor Compositions for the R-32/125/134a System; the Baseline

is from the Lemmon-Jacobsen Model in REFPROP 6.0. 30

Figure 13. Comparison of Densities for the R-32/125/134a System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0 3

Figure 14. Comparison of Bubble-Point Pressures for the R-32/143a System; the Baseline is from the Lenmion-Jacobsen Model in REFPROP 6.0.............. 32

Figure 15. Comparison of Vapor Compositions for the R-32/143a System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0. 33

Figure 16. Comparison of Densities for the R-32/143a System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0 34

Figure 17. Comparison of Bubble-Point Pressures for the R-125/143a System; the Baseline

is from the Lemmon-Jacobsen Model in REFPROP 6.0............................... 35

Figure 18. Comparison of Vapor Compositions for the R-125/143a System; the Baseline is from the Lemmon Jacobsen Model in REFPROP 6.0 36

Figure 19. Range of Measured Temperatures and Pressures for Isochoric (p,p,T) Data for a Mixture of R-125/143a with x(R-125) = 0.49996 Mole Fraction

(0.58812 Mass Fraction) 37

IV
Page

Figure 20. Comparison of Densities for the R-125/143a System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.Q 38

Figure 21. Comparison of Bubble-Point Pressures for the R-143a/134a System; the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0 39

Figure 22. Comparison of Vapor Compositions for the R-143a/134a System; the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0 40

Figure 23. Comparison of Densities for the R-143a/134a System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0 41

Figure 24. Comparison of Bubble-Point Pressures for the R-32/290 System; the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0 42

Figure 25. Comparison of Vapor Compositions for the R-32/290 System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0. 43

Figure 26. Comparison of Densities for the R-32/290 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0 44

Figure 27. Comparison of Bubble-Point Pressures for the R-125/290 System; the Baseline is from the Lenmion-Jacobsen Model in REFPROP 6.0 45

Figure 28. Comparison of Vapor Compositions for the R-125/290 System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0L 46

Figure 29. Comparison of Densities for the R-125/290 System; the Baseline is from the

Lenmion-Jacobsen Model in REFPROP 6.0 47

Figure 30. Comparison of Bubble-Point Pressures for the R-134a/290 System; the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0 48

Figure 31. Comparison of Vapor Compositions for the R-134a/290 System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0 49

Figure 32. Comparison of Densities for the R-134a/290 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0 50

Figure 33. Comparison of Vapor Pressures for R-41; the Baseline is from Equation (1) 51

Figure 34. Range of Measured Temperatures and Pressures for Isochoric (p,p,T) Data

for R-41 52

Figure 35. Comparison of Densities for R-41; the Baseline is from the MBWR Equation of State

in REFPROP 6.0 53

Figure 36. Comparison of the Isochoric Heat Capacity Data at Saturation and in the SinglePhase Liquid Region for R-41; the Baseline is from the MBWR Equation of State

in REFPROP 6.0 54

Figure 37. Comparison of the Vapor Speed of Sound Data for R-41; the Baseline is from the

MBWR Equation of State in REFPROP 6.0 55

Figure 38. Comparison of Bubble-Point Pressures for the R-4 1/744 System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0. 56
Page

Figure 39. Comparison of Vapor Compositions for the R-4 1/744 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.Q.... 57

Figm^e 40. Range of Measured Temperatures and Pressures for Isochoric (p,p,T) Data for a Mixmre of R-4 1/744 with x(R-41) = 0.49982 Mole Fraction (0.43591 Mass Fraction). 58

Figure 41. Comparison of Isochoric (p,p,T) Data for the Equimolar Mixture of R-4 1/744;

the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.Q 59

Figure Bl. Schematic Diagram of the Dynamic Phase Equilibrium Apparatus B-6

Figure B2. Schematic Diagram of the Bubble-Point and Near-Saturation (p,p,T) Apparatus B-7

Figure B3. Schematic Diagram of the Isochoric (p,p,T) Apparatus B-8

Figure B4. Schematic Diagram of the Calorimeter B-9

Figure B5. Schematic Diagram of the Spherical Resonator Speed of Sound Apparatus..... B-10

VI
LIST OF TABLES

Page

Table 1. Summary of the Refrigerant Systems Studied, the Properties Measured for Each

System, and the Molecular Weights of the I*ure Components A-2

Table 2. Summary of the Experimental Uncertainties for the Apparatus Used in This

Project A-3

Table 3. Vapor-Liquid Equilibrium Data for R-32/134a Mixtures from 280 to 340 K

(44 to 152°F) A-4

Table 4. Near-Saturation (p,p,T) Data for R-32/134a Mixtures from 279 to 340 K

(43 to 152*'F) A-6

Table 5. Vapor-Liquid Equilibrium Data for R-32/125 Mixtures from 280 to 340 K

(44 to 152°F) A-8

Table 6. Near-Saturation (p,p,T) Data for R-32/125 Mixtures from 279 to 341 K

(43 to 154"?) A-9

Table 7. Vapor-Liquid Equilibrium Data for R-125/134a Mixtures from 280 to 340 K

(44 to 152°F) A-11

Table 8. Bubble-Point Pressures for R-125/134a Mixtures from 280 to 340 K

(45 to 153°F) A-12

Table 9. Near-Saturation (p,p,T) Data for R-125/134a Mixtures from 280 to 342 K

(44 to 15TF) A-13

Table 10. Vapor-Liquid Equilibrium Data for R-32/125/134a Mixtures from 280 to 340 K

(44 to 152°?)..... A-14

Table 11. Bubble-Point Pressures for R-32/125/134a Mixtures from 221 to 345 K

(-62 to 162°F) A-15

Table 12. Near-Saturation (p,p,T) Data for R-32/125/134a Mixtures from 244 to 346 K

(-21 to 163°F) A-16

Table 13. Vapor-Liquid Equilibrium Data for R-32/143a Mixtures from 280 to 340 K

(44 to 152°F) A-18

Table 14. Near-Saturation (p,p,T) Data for R-32/143a Mixtures from 279 to 340 K

(43 to 152°F) A-I9

Table 15. Vapor-Liquid EquUibrium Data for R-125/143a Mixtures from 280 to 326 K

(44 to UTF) A-21

Table 16. Bubble-Point Pressures for R-125/143a Mixtures from 280 to 325 K

(45 to 125°F) A-22

Table 17. Near-Saturation (p,p,T) Data for R-125/143a Mixtures from 280 to 328 K

(44 to 13rF) A-23

Table 18. Isochoric (p,p,T) Data from 200 to 400 K (-100 to 260°F) for a Mixture of

R-125/143a with x(R-125) = 0.49996 Mole Fraction (0.58812 Mass Fraction) A-24

vu
Page

Table 19. Isochoric Heat Capacities (Q,) from 205 to 344 K (-90 to 160°F) for a Mixture of R-125/143a with x(R-125) = 0.49996 Mole Fraction

(0.58812 Mass Fraction) A-33

Table 20. Vapor-Liquid Equilibrium Data for R-143a/134a Mixtures from 280 to 340 K

(45 to 152°F) A-37

Table 21. Bubble-Point Pressures for R-143a/134a Mixtures from 281 to 340 K

(46 to 153°F) A-38

Table 22. Near-Saturation (p,p,T) Data for R-143a/134a Mixtures from 280 to 343 K

(45 to 158'='F) A-39

Table 23. Vapor-Liquid Equilibrium Data for R-32/290 Mixtures from 280 to 341 K

(44 to 154°F). ("LL" indicates possible liquid-liquid inmiiscibility). A-40

Table 24. Near-Saturation (p,p,T) Data for R-32/290 Mixtures from 278 to 341 K

(42 to 153°F) .A-43

Table 25. Vapor-Liquid Equilibrium Data for R-125/290 Mixtures from 280 to 364 K

(44 to 195°F) A-45

Table 26. Near-Saturation (p,p,T) Data for R-125/290 Mixtures from 280 to 326 K

(44 to 128°F) A-47

Table 27. Vapor-Liquid Equilibrium Data for R-134a/290 Mixtures from 279 to 357 K

(43 to 182''F). ("LL" indicates possible liquid-liquid immiscibility). ....................... A-48

Table 28. Near-Saturation (p,p,T) Data for R-134a/290 Mixtures from 278 to 357 K

(40 to 183°F)...... A-51

Table 29. Vapor Pressures for R-41 from 252 to 312 K (-6 to 102*'F) Measured Using the

Vapor-Liquid Equilibrium Apparatus.................. A-53

Table 30. Vapor Pressures for R-41 from 170 to 317 K (-154 to 111°F) Measured Using the

Isochoric (p,p,T) Apparatus................................ A-54

Table 31. Near-Saturation (p,p,T) Data for R-41 from 274 to 296 K

(34 to 73''F)..,..................................................... A-55

Table 32. Comparisons of Critical Point Parameters for R-4L........... A-56

Table 33. Isochoric (p,p,T) Data for R-41 from 132 to 400 K (-222 to 260°F).. A-57

Table 34. Saturated Liquid Densities for R-41 from 131 to 309 K (-224 to 96°F)

Extrapolated from Isochoric (p,p,T) Data A-70

Table 35. Isochoric Heat Capacities (Cv) for R-41 from 148 to 343 K (-193 to 157"'F).. .......... A-71

Table 36. Two-Phase Heat Capacities for R-41 from 136 to 314 K (-215 to 106°F). ....... ......... A-75

Table 37. Vapor Phase Speed of Sound Data for R-41 from 249.5 to 350 K

(-1 1 to 170°F) A-79

Table 38. Vapor-Liquid Equihbrium Data for R-4 1/744 Mixtures from 218 to 290 K

(-68 to 62°F) A-80

vui
Page

Table 39. Isochoric (p,p,T) Data from 192 to 400 K (-1 14 to 260°F) for a Mixture of

R41/744 with x(R^l) = 0.49982 Mole Fraction (0.43591 Mass Fraction) A-81

Table Dl. Summary of the Gas Chromatograph Settings, the Separation Columns, and the

Standard Reference Mixtures Used for Each Refrigerant System D-4

IX

SELECTED THERMODYNAMIC PROPERTIES FOR MIXTURES OF R-32

(DIFLUOROMETHANE), R-125 (PENTAFLUOROETHANE), R-134A (1,1,1,2-

TETRAFLUOROETHANE), R-143A (1,1,1-TRIFLUOROETHANE), R-41

(FLUOROMETHANE), R-290 (PROPANE), AND R-744 (CARBON DIOXIDE)

Cynthia D. Holcomb

Joseph W. Magee

Jennifer L. Scott

Stephanie L. Outcalt

William M. Haynes

Physical and Chemical Properties Division

Chemical Science and Technology Laboratory

National Institute of Standards and Technology

Boulder, Colorado 80303

Numerous fluids and fluid mixtures have been identified as promising

alternatives to the HCFC and CFG refrigerants R-22, R-22/115, and R-13, but, for

many of them, reliable thermodynamic data are not available. Thermophysical

property measurements were carried out on one pure fluid, ten binary mixtures, and one temary mixture which are potential HCFC altematives. Vapor-liquid equilibrium,

near-saturation density, (p,p,T), heat capacity, and vapor-phase speed of sound data were measured over wide ranges of temperature and pressure. The systems studied

include R-41, R-32/134a; R-32/125; R-125/134a; R-32/143a; R-125/143a; R143a/134a; R-32/290; R-125/290; R-134a/290; R-41/744; and R-32/1 25/1 34a. The

new data compliment the existing data and help to resolve problems and differences

that exist in and between existing data sets. The measured data were compared to selected literature results, where available, and to predictions from REFPROP 6.0.

Keywords: alternative refrigerants; vapor-liquid equilibria; pressure-volumetemperature; vapor speed of sounds; refrigerant mixtures

1. INTRODUCTION

Altemative refrigerants have become increasingly important to industrial applications as the

ozone-depleting refrigerants are being phased out under the requirements of the Montreal Protocol.

The main objective of this project was to measure the data needed to refine the models used by

engineers for evaluating new refrigerants and refrigerant mixtures and for designing refrigeration

cycles with these new fluids. To meet this objective, several tasks were addressed: (1) measure

properties of pure fluids and mixture properties identified as replacements for specific fluids, (2)

investigate the behavior of propane-containing mixtures (propane has been suggested as an additive

to enhance oil solubility), and 3) provide data needed to test the ability of models to predict

multicomponent mixture properties and complex phase behavior. The data from this project provide

property information on previously unstudied systems, supplement existing data sets, resolve
problems and differences between existing data sets, and refine models used in equipment design.

Ten binary mixtures, one ternary, and one pure fluid were studied in this project.

Knowledge of the phase behavior of a fluid is important in determining if that fluid will be a

suitable replacement for a pure refrigerant or refrigerant mixture. Azeotropic mixtures are important

because they behave more like pure fluids which require simpler refrigeration cycle designs. A

system exhibiting liquid-liquid immiscibility and azeotropic behavior may also be exploited in a

refrigeration cycle. It was important to gather information on systems that might exhibit hquid-liquid

immiscibility for two reasons. The first was to provide the data needed to examine the behavior of

these systems in a refrigeration cycle. The second was to provide test data for implementing

algorithms in predictive packages that can adequately represent liquid-liquid immiscibility.

The phase behavior for the systems measured in this report are classified as either Type I or Type II systems. Figure 1 shows generic examples of these two types of phase behavior. Type I systems have a continuous critical Une and can have an azeotrope; the azeotrope can be either a

positive pressure or a negative pressure azeotrope. All three subclasses of Type I systems were

observed in this study. R-32/134a, R-125/1 34a, R-32/143a, and R-4 1/744 are nonazeotropic systems

with the simplest type of phase behavior. R-32/125 and R-143a/134a are Type I systems with positive

pressure azeotropes. R-125/143a is a Type I system with a negative pressure azeotrope. Type II systems have a continuous critical line with liquid-hquid immiscibility. Type n systems also can have

azeotropes, and the liquid-liquid region can either encompass the azeotrope or the azeotrope can be

outside the Uquid-liquid region. R-32/290, R- 125/290, and R-134a/290 are aH beUeved to be Type II systems where the liquid-liquid immiscibility region encompasses the azeotrope. The azeotrope

persists into the critical region, but the liquid-hquid region disappears around 310 K (98°F).

A full summary of the data taken on each system is presented in Table 1; tables of the

thermophysical property data are presented in Appendix A, The experimental uncertainties of these

measurements are reported in Table 2. The temperature scale used in all of the measurements was

ITS-90. The descriptions of the apparatus used for the measurements in this project are presented in

Appendix B. Appendix C describes how the state-point uncertainties were estimated. Appendix D

summarizes the gas chromatograph calibration procedures and settings. Appendix E contains reprints

of the papers describing the models [1] and predictive package [2] used for comparing the data sets.

2. SIGNIFICANT RESULTS

2.1 Thermophysical Properties of R-32/125/134a and Its Constituent Binaries

Property measurements of the ternary R-32/125/134a system are important in meeting the

objectives of this project. This ternary is a possible replacement for R-22 and/or R-22/115, and the

ternary results can be used to test the capability of the Helmholtz energy model of Lemmon and

Jacobsen [1]. This model is used to calculate the thermodynamic properties of mixtures in REFPROP
6.0 [2] and to predict the thermophysical properties for a ternary system based on the interaction

parameters of its constituent binaries. The constituent binaries of R-32/125/134a are also potential

replacements for R-22 and/or R-22/115.

2.1.1 R-32/134a

R-32/134a is a Type I mixture with no azeotrope. It is one of the least complicated systems to model, but it is an important mixture for ternary calculations. The data for the R-32/134a system

cover five isotherms from 279 to 340 K (43 to 152**?). Forty-eight vapor-hquid equilibrium and 44

near-saturation (p,p,T) measurements were made. These data, presented in Tables 3 and 4, have been

compared to five other data sets available in the literature. The deviations between the data and the

values predicted from the Lemmon-Jacobsen model are presented in Figures 2 through 4. In Figure

2, the deviations between the bubble-point pressures and the predicted values are presented. The data

from this woric agree with the data of Nagel and Bier [3] within ±1%. Most of the data of Higashi [4]

and of Widiatmo et al. [5] also agree with the data of this work within ±1% except for a few points.

The data of Fujiwara et al. [6] are 3% lower than any of the other data sets. The data of Deifbaugh

and Morrison [7] that are at similar compositions agree within ±0.5% with the data of Uiis work.

The scatter in the data might appear large compared to the experimental uncertainties quoted in

Table 2. A distinction must be made between the experimental uncertainty of a specific parameter

measurement (such as temperature, pressure, or composition) and the state-point uncertainty. The

state point is system dependent and is defined as the equilibrium bubble-point pressure and vapor

composition at a given liquid composition and temperature. Each of the parameters (P,T,x,y) has an

experimental uncertainty associated with the measurement of that property. However, the state-point

uncertainty is a function of the uncertainties of each of these four parameters as weU as the

dependences between them. The experimental uncertainty of the bubble-point pressure state point is a function of the uncertainty of the temperature measurement, the Uquid composition measurement,

the pressure measurement, and the relation of the bubble-point pressure with temperature and liquid

composition for the specific system under study. The state-point uncertainties are different for each

system and vary with temperature and composition. A detailed discussion of estimates for the state- point uncertainties is presented in Appendix C. The bubble-point pressure state-point uncertainty for

the R-32/134a system at temperatures from 280 to 340 K (44 to 152°F) ranges from ±0.22 to ±0.30%.

In Figure 3, the vapor compositions of the data sets are compared. In general, the four data sets agree within ±0.015 mole fraction R-32. The data of Fujiwara et al. [6] are systematically 0.015 mole

fraction R-32 lower than the predictions of the Lemmon-Jacobsen model in REFPROP 6.0. The

vapor-composition state-point uncertainty ranges from ±0.006 to ±0.013 mole fraction R-32 for the

R-32/134a system at temperatures from 280 to 340 K (44 to 152°F). As with the bubble-point

pressure state-point uncertainty, a similar estimate of the state-point uncertainty of the vapor
composition can be calculated. A detailed discussion of the estimate of the vapor-composition state- point uncertainty is presented in Appendix C.

The deviations for the near-saturation (p,p,T) data for R-32/134a are presented in Rgure 4. The

liquid densities agree within ±1.0% of the predicted values from the Lemmon-Jacobsen model in

REFPROP 6.0. The vapor densities agree within ±2% of the predicted values and most of the data of

Weber and Defibaugh [8]. The liquid-density state-point uncertainty for the R-32/134a system at temperatures from 279 to 340 K (43 to 152°F) ranges from ±0.22 to ±0.26%. The vapor-density

state-point uncertainty for the R-32/134a system at temperatures from 309 to 340 K (97 to 152°F)

ranges from ±0.32 to ±0.79%. As with the bubble-point pressure state-point uncertainty, a similar

estimate of the state-point uncertainty of the liquid and vapor densities can be calculated. A detailed

discussion for the estimate of the density state-point uncertainties is presented in Appendix C.

2.1^ R-32/125

R-32/125 is a Type I mixture with a slight positive pressure azeotrope. The data for the R32/125 system cover five isotherms from 279 to 341 K (43 to 154°F). Thirty vapor-liquid

equilibrium and 45 near-saturation (p,p,T) measurements were made. These data, presented in Tables

5 and 6, were compared to five other data sets available in the literature. The deviations between the

data and the values predicted from the Lemmon-Jacobsen model in REFPROP 6.0 are presented in

Figures 5 through 7. In Figure 5, the deviations between the bubble-point pressures and the predicted

values are presented. The data from this woik agree with the data of Nagel and Bier [3] and that of

Widiatmo et al. [9] within ±1% except for one point on the 325 K (125''F) isotherm. Except for two

points, the data of Higashi [10] also agree with the data of this work within ±1%. The data of

Deifbaugh and Morrison [7] agree within ±0.5% with the data of this work that is at a similar

composition. The data of Fujiwara et al. [6] are 2% higher than any of the other data sets. The

bubble-point pressure state-point uncertainty for the R-32/125 system at temperatures from 280 to 340 K (44 to 152°F) ranges from ±0.15 to ±0.19%.

In Figure 6, the vapor compositions of the data sets are compared. In general, the four data sets agree within ±0.013 mole fraction R-32. The vapor-composition state-point uncertainty ranges from

±0.008 to i:0.009 mole fraction R-32 for the R-32/125 system at temperatures from 280 to 340 K

(44 to 152^F).

The deviations for the near-saturation (p,p,T) data for R-32/125 are presented in Figure 7. The

liquid densities agree within ±1% of the predicted values from the Lemmon-Jacobsen model in

REFPROP 6.0. The vapor densities agree within ±2% of the predicted values and the data of Weber

and Defibaugh [8]. The liquid-density state-point uncertainty for the R-32/125 system at temperatures from 279 to 340 K (43 to 152°F) ranges from ±0.22 to 0.25%. The vapor-density statepoint uncertainty for the R-32/125 system at temperatures from 294 to 341 K (70 to 154*F) ranges

from ±0.34 to ±0.90%.
2.1.3 R-125/134a

R-125/134a is a Type I mixture with no azeotrope. The data for the R-125/134a system cover

five isotherms from 280 to 342 K (44 to 157°F). Thirty vapor-liquid equilibrium and 17 nearsaturation (p,p,T) measurements were made. Ten bubble-point pressure measurements on standard

mixtures were also made. The data are presented in Tables 7 through 9. Bubble-point pressure

measurements on standard mixtures were included for this system because of difficulties involved

with the gas chromatograph (GC) calibration. The bubble-point pressure measurements eliminate the

uncertainty introduced by sampling and analysis to determine the liquid composition. These

measurements represent an independent verification of the bubble-point measurements from the

vapor-liquid equilibrium measurements, and can be used to check the validity of the GC calibration.

The data were compared to two other data sets available in the literature. The deviations between the

data and the predicted values from the Lemmon-Jacobsen model are presented hi Figures 8 through

10. In Figure 8, the deviations between the bubble-point pressures and the predicted values are

presented. The data from this work agree with the data of Nagel and Bier [3] within ±2%. Except for

two points, the data of Higuchi and Higashi [11] agree with the data of this work within ±2%. The

bubble-point pressure state-point uncertainty for the R- 125/1 34a system at temperatures from 280 to 340 K (44 to 152°F) ranges from ±0.23 to ±0.30%.

In Figure 9, the vapor compositions of the data sets are compared. In general, the three data sets agree Avithin ±0,018 mole fraction R- 125. The vapor-composition state-point uncertainty ranges from

±0.006 to ±0.012 mole fraction R-125 for the R-125/134a system at temperatures from 280 to 340 K

(44 to 152°F).

The deviations for the near-saturation (p,p,T) data for R-125/134a are presented in Figure 10. The liquid densities agree within ±1% of the predicted values from the Lemmon-Jacobsen model in

REFPROP 6.0. The vapor densities agree within ±2% of the predicted values and most of the data of

Weber and Defibaugh [8]. The liquid-density state-point uncertainty for the R-125/134a system at temperatures from 280 to 340 K (44 to 153°F) ranges from ±0.21 to ±0.24%. The vapor-density

state-point uncertainty for the R-125/134a system at temperatures from 296 to 342 K (73 to 157°F)

ranges from ±0.51 to ±1.08%.

2.1.4 R-32/125/134a

The data for the R-32/125/134a system cover five isotherms from 280 to 340 K (44 to 152°I^)

with some additional measurements at temperatures from 221 to 345 K(-62 to 162°F) that were

obtained under a project funded by ICI. Twenty-four vapor-liquid equilibrium and 43 nearsaturation (p,p,T) measurements were made. Thirty-four bubble-point measurements on standard

mixtures were performed to determine the reliability of the GC caUbration procedure for a ternary

mixture. The data are presented in Tables 10 through 12. The data were compared to the other data

set available in the literature. The deviations between the data and the predicted values from the
Lemmon-Jacobsen model are presented in Figures 11 through 13. In Figure 10, the deviations

between the experimental bubble-point pressures and predicted values are presented. The data from

this woric agree with the data of Nagel and Bier [3] and of Higashi [12] within ±3%. The bubblepoint pressure state-point uncertainty for the R-32/125/134a system at temperatures from 280 to 340

K (44 to 152^F) ranges from ±0.15 to ±0.30%.

In Figure 12, the vapor compositions of the data sets are compared. In general, the three data

sets agree within ±0.018 mole fraction R-32. The vapor-composition state-point uncertainty ranges

from ±0.006 to ±0.013 mole fraction R-32 for the R-32/125/134a system at temperatures from 280

to 340 K (44 to 152°F).

The deviations for the near-saturation (p,p,T) data for R-32/1 25/1 34a are presented in Figure

13. The liquid densities agree within ±1% of the predicted values from the Lemmon-Jacobsen model

in REFPROP 6.0. The vapor densities agree within ±2% of the predicted values and most of the data

of Gillis et al. [13]. The liquid-density state-point uncertainty for the R-32/125/134a system at temperatures from 244 to 346 K (-21 to 163°F) ranges from ±0.21 to ±0.26%. The vapor-density

state-point uncertainty for the R-32/125/134a system at temperatures from 313 to 343 K (103 to 158°F) ranges from ±0.32 to ±1.08%.

The Lemmon-Jacobsen model accurately predicted the phase behavior and densities of the

ternary using only the binary interaction parameters. The model did not require any additional

parameters specifically for the ternary mixture. This result confirms that, for this system, the

Lemmon-Jacobsen model can be used to predict multicomponent mixture properties when the model

uses mixture parameters for the constituent binaries which have been fit to experimental data.

22 Mixtures with R-143a

Three mixtures containing R-143a were studied. These mixtures are potential substitutes for R22 and /or R-22/115, and the data for these systems with different types of phase behavior can be

used to test mixture models.

2^.1 R-32/143a

R-32/143a is a Type I system with no azeotrope. The data for the R-32/143a system cover five

isotherms from 279 to 340 K (43 to 152°F). Twenty-nine vapor-liquid equilibrium and 49 nearsaturation (p,p,T) measurements were made. The data are presented in Tables 13 and 14. In Figure

14, the bubble-point pressures are compared. The bubble-point pressures of this work agree with the

predicted values from the Lemmon-Jacobsen model within ±0.7%. The data of Fujiwara et al. [6] are 2 to 3% higher than the data from this work. The bubble-point pressure state-point uncertainty for R32/143a at temperatures from 280 to 340 K (44 to 152°F) ranges from ±0.14 to ±0.20%.

In Figure 15, the vapor compositions are compared. The data of Fujiwara et al. [6] and this work agree within ±0.025 mole fraction R-32 of the predicted values of the Lemmon-Jacobsen model
in REFPROP 6.0. The vapor-composition state-point uncertainty for R-32/143a at temperatures from

280 to 340 K (44 to 152°F) ranges from ±0.007 to ±0.010 mole fraction R-32.

In Figure 16, the liquid and vapor near-saturation (p,p,T) data are compared to the predicted

values from the Lemmon-Jacobsen model in REFPROP 6.0. The liquid densities agree within ±1%

except for the 340 K (152°F) isotherm which is close to the critical point of R-143a. The vapor

densities for the two isotherms closest to the critical region also show larger deviations. The liquiddensity state-point uncertainty for the R-32/143a system at temperatures from 279 to 340 K (43 to 152°F) ranges from ±0.21 to ±0.24%. The vapor-density state-point uncertainty for tiie R-32/143a

system at temperatures from 294 to 340 K (70 to 152°F) ranges from ±0.27 to ±0.69%.

2.2.2 R-125/143a

R-125/143a is a Type I system with a weak negative pressure azeotrope. The data for the R125/1 43a system cover four isotherms from 280 to 328 K (44 to 131°F). Twenty-five vapor-hquid

equilibrium and 14 near-saturation (p,p,T) measurements were made. Eleven bubble-point

measurements on standard mixtures were performed to determine the reliability of the GC calibration

procedure for this binary mixture. The data are presented in Tables 15 through 17. The single-phase

data for the R-125/143a system cover 15 isochores from 200 to 400 K (-100 to 260°F) at pressures

up to 35 MPa (5100 psi). Both gas- and liquid-phase (p,p,T) data were measured, as weU as hquidphase heat capacity data. The data are presented in Tables 18 and 19. Two-hundred eighty-one

isochoric (p,p,T) data points and 120 isochoric heat capacity measurements were made.

Rgure 17 shows the comparison of the bubble-point pressures for R-125/143a to the predicted

values from the Lemmon-Jacobsen model. The compositions were difficult to measure because the

vapor pressure curves of these two components are very similar. The experimental uncertainty of the

composition measurements for this system is approximately twice that of the other systems in this

study. The experimental uncertainty of the measured composition is ±0.008 mole fraction R-143a.

The data from this woric faU between the data of Takashima and Higashi [14] and those of Nagel and

Bier [15]. The bubble-point pressure state-point uncertainty for R-125/143a at temperatures from

280 to 326 K (44 to 127°F) ranges from ±0.15 to ±0.18%.

Figure 18 shows the comparison of the vapor compositions for R-125/143a to the values

predicted from the Lemmon-Jacobsen model. The vapor compositions for the three data sets agree

within ±0.011 mole fraction R-125. The vapor-composition state-point uncertainty for the R125/143a system at temperatures from 280 to 326 K (44 to 127°F) ranges from ±0.011 to ±0.012

mole fraction R-143a.

The ranges of temperature and pressure for the isochoric (p,p,T) data are shown in Figure 19. Figure 20 shows the comparison of the liquid and vapor near-saturation (p,p,T) and the isochoric

(p,p,T) data to the predicted values from the Lemmon-Jacobsen model. The near-saturation (p.p.T)

liquid densities agree with the model within ±0.8%, and the near-saturation (p.p.T) vapor densities

agree with the model within ±0.75%. The liquid-phase isochoric (p.p.T) data agree with the model
within ±0.25%, and the vapor-phase isochoric (p,p,T) data agree with the model and the data of

Weber and Defibaugh [8] within ±0.85%. The liquid-density state-point uncertainty for the nearsaturation (p,p,T) data for the R-125/143a system at temperatures from 280 to 325 K (44 to 125°F)

ranges from ±0.24 to ±0.27%. The vapor-density state-point uncertainty for the near-saturation

(p,p,T) data for the R-125/143a system at temperatures from 296 to 328 K (74 to 131°F) ranges from

±0.48 to ±0.60%.

2^.3 R-143a/134a

R-143a/134a is a Type I system with no azeotrope. The data for the R-143a/134a system cover

five isotherms from 280 to 343 K (45 to 158°F). Twenty-nine vapor-liquid equilibrium and 17 nearsaturation (p,p,T) measurements were made. Eleven bubble-point measurements on standard mixtures were performed to determine the reliability of the GC calibration procedure for this binary mixture.

The data are presented in Tables 20 through 22.

Figure 21 shows comparisons of the bubble-point pressures for R-143a/134a with the predicted

values from the Lemmon-Jacobsen model. The bubble-point pressures agree with those of Nagel and

Bier [15] within ±2%. The bubble-point pressure state-point uncertainty for the R-143a/134a system

at temperatures from 280 to 340 K (44 to 152°F) ranges from ±0.20 to ±0.25%.

Figure 22 shows comparisons of the vapor compositions for R-125/143a with the predicted

values from the Lemmon-Jacobsen modeL The vapor compositions agree with those of Nagel and

Bier [15] within ±0.018 mole fraction R-143a. The vapor-composition state-point uncertainty for the

R-143a/134a system at temperatures from 280 to 340 K (45 to 152°F) ranges from ±0.006 to ±0.011

mole fraction R-143a.

Figure 23 shows comparisons of the liquid and vapor near-saturation (p,p,T) data with the

values predicted from the Lemmon-Jacobsen model. The liquid densities agree with the predicted

values within ±1%. The liquid-density state-point uncertainty for the R-143a/134a system at temperatures from 280 to 340 K (45 to 153°F) ranges from ±0.22 to ±0.24%. The vapor-density

state-point uncertainty for the R-143a/134a system at temperatures from 312 to 343 K (102 to 158°F) ranges from ±0.38 to ±0.73%. Additional data in the moderate temperature and pressure

range are needed to determine the accuracy of the near-saturation vapor-phase (p,p,T) data and to

better optimize the fit.

2.3 Mixtures with R-290

There are two main reasons for studying refrigerant mixtures containing R-290 (propane). One

of the difficulties with the chlorine-free alternative refrigerants is their immiscibility with the oils used

in refrigeration cycles. The addition of R-290 to the mixture to enhance oil solubility is under

investigation. In order to model multicomponent systems that contain R-290, the mixture parameters

for binary systems containing R-290 are required. Another reason for studying R-290 is the complex
phase behavior exhibited by R-290 mixtures allows the models to be rigorously tested. These systems

have very strong positive azeotropes and possible hquid-liquid immiscibility. They provide a very

stringent test for the REFPROP 6.0 model and its ability to calculate properties for polar/nonpolar

mixtures with three phases in equilibrium.

Although the Lemmon-Jacobsen model used in REFPROP 6.0 is capable of predicting liquidliquid immiscibility, the algorithms implementing it in REFPROP 6.0 do not consider this behavior.

Because of the complexity of this phase behavior, more experimental data are required to model the

two separate liquid phases. Accurate liquid- and gas-phase (p,p,T) data and a knowledge of the

compositions of the two liquid phases and gas phase in equilibrium are necessary for accurately

determining the interaction parameters that can be used to model this phase behavior. Unfortunately,

the measurements scheduled for these systems included near-saturation (p,p,T) data, but not isochoric

(p,p,T) data. The vapor-liquid equilibrium apparatus was not designed to separate and to measure two

distinct liquid phases. In the region where the liquid-liquid immiscibility occurs, no stable liquid

densities or compositions can be recorded. The data available in the literature for these systems are

limited, possibly because all three systems are in the process of being patented. The overall

uncertainty of the measurements and the lack of data to optimize the model have resulted in

deviations between the model and the data that are larger than for the other refrigerant systems.

2.3.1 R-32/290

R-32/290 is a Type II system with a strong positive pressure azeotrope. It shows the most

extreme phase behavior of the three propane systems studied. It is a mixture of a very polar molecule

with a nonpolar molecule. The two molecules have strong repulsive interactions and exhibit a strong

positive pressure azeotrope. At temperatures below 310 K (98°F), the Uquid separates into two phases

that were observed visually. The vapor-liquid equilibrium apparatus was not designed to measure two

separate liquid phases, but the presence of two liquid phases is indicated by an unstable liquid

composition and density for a constant vapor composition, bubble-point pressure, and vapor density.

Measurements were made outside the liquid-liquid region where the liquid composition could be

measured. A few measurements were taken to determine the vapor composition in equilibrium with

the two Uquid phases.

The data for the R-32/290 system cover five isotherms from 278 to 341 K (42 to 154°F).

Seventy-five vapor-liquid equilibrium and 50 near-saturation (p,p,T) measurements were made. The

data are presented in Tables 23 and 24. There were no other data available for comparison for any of

the properties measured for this system.

Figure 24 shows the comparison of the bubble-point pressures to the predicted values of the

Lemmon-Jacobsen model in REFPROP 6.0. The data agree with the model within ±5%. The bubblepoint pressure state-point uncertainty for the R-32/290 system at temperatures from 280 to 340 K (44

to 153°F) ranges from ±0.14 to ±0.63%. The uncertainties for the systems with propane are greater

than for the other refrigerant systems studied because of the greater sensitivity of the bubble-point
pressure to changes in composition. This sensitivity means that small uncertainties in the composition

measurement correspond to much larger uncertainties in the bubble-point pressure.

Figure 25 shows the comparison of the vapor compositions to the predicted values from the

Lemmon-Jacobsen model in REFPROP 6.0. The vapor compositions agree with the model within

±0.05 mole fraction R-32. The vapor-composition state-point uncertainty for the R-32/290 system at temperatures from 280 to 340 K (44 to 153°F) ranges from ±0.015 to ±0.025 mole fraction R-32.

The state-point uncertainties are much larger because of the sensitivity of the equilibrium vapor

composition to both temperature and liquid composition.

Figure 26 shows the comparison of the liquid and vapor near-saturation (p,p,T) data to the

predicted values from the Lemmon-Jacobsen model in REFPROP 6.0. The liquid and vapor densities

agree with the predicted values within ±7%. The liquid-density state-point uncertainty for the

R-32/290 system at temperatures from 278 to 340 K (42 to 153°F) ranges from ±0.23 to ±0.33%.

The vapor-density state-point uncertainty for the R-32/290 system at temperatures from 296 to 341 K

(73 to 153°F) ranges from ±0.28 to ±3.10%.

2.3.2 R-125/290

R- 125/290 is a Type II system with a strong positive pressure azeotrope that may exhibit liquidliquid immiscibility at temperatures below 280 K (44°F), The data for the R-125/290 system cover

eight isotherms from 280 to 364 K (44 to 195°F). Sixty-three vapor-liquid equilibrium and 26 nearsaturation (p,p,T) measurements were made. The data are presented in Tables 25 and 26. In Figure 27, comparisons of the bubble-point pressures for the R-125/290 system with the

predicted values from the Lemmon-Jacobsen model in REFPROP 6.0 are shovra. The bubble-point

pressures agree with values from the model within ±4.5%. The bubble-point pressure state-point

uncertainty for the R-125/290 system at temperatures from 280 to 364 K (44 to 195°F) ranges from

±0.15 to ±0.42%.

In Figure 28, comparisons of the vapor compositions for the R-125/290 system to the predicted

values of the Lemmon-Jacobsen model in REFPROP 6.0 are shown. The vapor compositions agree

with the predicted values within ±0.04 mole fraction R-125. The vapor-composition state-point

uncertainty for the R-125/290 system at temperatures from 280 to 364 K (44 to 195°F) ranges from

±0.006 to ±0.015 mole fraction R-125.

In Figure 29, comparisons of the liquid and vapor near-saturation (p,p,T) data to the predicted

values of the Lemmon-Jacobsen model in REFPROP 6.0 are shown. The liquid densities agree with

the model within ±2%, and the vapor densities agree within ±3.5%. The liquid-density state-point

uncertainty for the R-125/290 system at temperamres from 280 to 325 K (44 to 125°F) ranges from

±0.26 to ±0.34%. The vapor-density state-point uncertainty for the R-125/290 system at temperatures

from 294 to 326 K (69 to 128°F) ranges from ±0.26 to ±2.30%.

10
2J.3 R-134a/290

R-134a/290 is a Type 11 system with a strong positive pressure azeotrope. The data for the R134a/290 system cover six isotherms from 278 to 357 K (42 to 183°F). Seventy-two vapor-liquid

equilibrium and 52 near-saturation (p,p,T) measurements were made. The data are presented in

Tables 27 and 28. Because of the liquid-hquid immiscibility, only qualitative comparisons can be

made with the Lemmon-Jacobsen model.

Figure 30 shows the comparison of the bubble-point pressures to the predicted values of the

Lemmon-Jacobsen model in REFPROP 6.0. The data of Kleiber [16] agree with the data from this work within a few percent. The data of Jadot and Frere [17] appears to be quahtatively different from

our data and from the data of Kleiber [16]. The bubble-point pressure state-point uncertainty for the

R-134a/290 system at temperatures from 280 to 357 K (44 to 183''F) ranges from ±0.13 to ±0.57%.

Figure 31 shows the comparison of the vapor compositions to the predicted values of the

Lemmon-Jacobsen model in REFPROP 6.0. The three data sets agree within ±0.09 mole fraction R134a. The vapor-composition state-point uncertainty for the R-134a/290 system at temperatures from

280 to 357 K (44 to 183°F) ranges from ±0.005 to ±0.019 mole fraction R-134a.

Figure 32 shows the comparison of the liquid and vapor near-saturation (p,p,T) data to the

predicted values from the Lemmon-Jacobsen model in REFPROP 6.0. The liquid densities agree with

the model within ±5%, and the vapor densities show increasing deviations as the temperature

increases. The liquid-density state-point uncertainty for the R-134a/290 system at temperatures from

278 to 356 K (40 to 181"'F) ranges from ±0.29 to ±0.42%. The vapor-density state-point uncertainty

for the R-125/290 system at temperatures from 311 to 357 K (100 to 183°F) ranges from ±0.24 to ±2.06%.

2.4 R-41 and R-41/744

R-41 is a possible substitute for R-13. While R-23 is most often considered for these

applications, the long atmospheric lifetime and high global warming potential of R-23 leaves open

the search for a more environmentally desirable fluid. Unfortunately, R-41 is a flammable fluid. R744 (carbon dioxide) is a nonflammable fluid, but its freezing point is too high to be used as a replacement for R-13. The expectations that a mixture of R^1/744 would have reduced flammability

relative to pure R-41 and a lower freezing point than pure R-744 make this mixture an attractive

possibihty. An accurate equation of state exists for R-744 and was incorporated in REFPROP 6.0. A

preliminary MBWR equation of state for R-41 was incorporated in REFPROP 6.0. Measurements on

R-41 and R-41/744 were performed. Comparisons with predictions from REFPROP 6.0 were used to validate the data.

II
2.4.1 R-41

A variety of single-phase and two-phase data was collected for R-4L Twenty vapor pressure

measurements were made in the vapor-liquid equilibrium apparatus. Vapor pressure data were also

measured in the isochoric (p,p,T) apparatus. Thirty-one vapor pressure measurements were obtained

with this apparatus. The vapor pressure data are presented in Tables 29 and 30. The vapor pressure

data from this work were combined with the data of Oi et al. [18], and a vapor pressure correlation

that is valid from the triple point to the critical point was developed. This equation is of the fonn,

P = P a ^ c

ajT + ^2^^ + a^T^ + a^x^ iTx

(1)

where

T = T =

Tc =

Pa =

Pc =

1 - T/Tc,

temperature,

critical temperature,

vapor pressure, and

critical pressure.

The critical temperature used in the fit is 317.28 K (111.4°F) (This equation requires the absolute

scale temperatures. Therefore, degrees Fahrenheit must be converted to degrees Rankine by adding

459.7). The critical pressure used in the fit is 5.897 MPa (855.3 psia). The coefficients ai are

dimensionless and are as follows:

ai =

a2 =

a3 =

a4 =

-7.01707106

1.33436478

-1.78330695

-1.84394112

Figure 33 shows comparisons of the two data sets from this work and the vapor pressures of Oi et al.

[18] with the calculated values from eq. (1). The data agree within ±0.3%.

The critical point parameters used in eq. (1) were determined from the near-saturation vapor

and liquid (p,p,T) data and the vapor pressure data collected in this work. Twenty-four nearsaturation vapor and liquid (p,p,T) data points were obtained. The data are presented in Table 31. The

vapor pressure data and near-saturation (p,p,T) data were used to estimate the critical point parameters

of R-41 using the method of Van Poolen et al. [19]. The critical point temperature is estimated to be

317.28 ± 0.08 K (111.43 ± 0.14°F). This temperature agrees with those of Bominaar et al. [20] and

12
of Biwas et al. [21] within ±0.12 K (±0.22°F). The critical pressure is estimated to be 5.897 ± 0.01

MPa (855.3 ± 1.4 psia). This critical pressure agrees within ±0.27 MPa (±4.3 psia) with the critical

pressures of Bominaar et al. [20] and of Biwas et al. [21]. The critical density is estimated to be 316.5

± 1.5 kg/m3 (19.74 ± 0.9 Ib/ft^). The critical density agrees within ±5.1 kg/m3 (±0.34 Ib/ft^) of the

critical densities of Bominaar et al. [20] and of Biwas et al. [21]. A summary of the critical point

parameters is presented in Table 32.

Liquid- and gas-phase (p,p,T) measurements for R-41 were taken along 17 isochores at temperatures from 132 to 400 K (-222 to 260°F) at pressures up to 35 MPa (5100 psi). Fourhundred forty-five isochoric (p,p,T) measurements were obtained. The data are presented in Table

33. Saturated liquid densities for R-41 were estimated by extrapolating the isochoric (p,p,T) data to the vapor pressures calculated from eq. (1). These calculated saturated liquid densities are presented

in Table 34. Figure 34 shows the range of temperatures and pressures covered by the isochoric

(p,p,T) data. Figure 35 compares the near-saturation (p,p,T) hquid densities, the isochoric (p,p,T) data,

and the calculated saturated liquid densities from the isochoric (p,p,T) data with calculated values

from the MBWR equation in REFPROP 6.0. All of the data agree with the predicted values within

±0.5%, and the majority was fit to within approximately 0.2% except in close proximity to the triple

point temperature and the critical temperature. The near-saturation densities agree with the predicted

values within ±0.3%. The overall fit of the (p,p,T) data has an average absolute deviation of ±0.048%

and a bias of 0.002%.

Liquid- and gas-phase heat capacity measurements were taken for R-41 along 17 isochores at temperatures from 148 to 343 K (-193 to 157°F) at pressures up to 33 MPa (4800 psi). Onehundred twenty-two isochoric heat capacity measurements and 133 saturated-liquid heat capacity

measurements were obtained. The data are presented in Tables 35 and 36. Figure 36 compares the

isochoric heat capacity data at saturation and in the single-phase liquid region with calculated values

from the MBWR equation in REFPROP 6.0. The liquid-phase heat capacity data agree with the

predictions within ±2.5 except near the critical point. The saturated heat capacity data agree with the

predictions within ±1.5% except near the critical point and near the triple point.

Vapor phase speed of sound measurements were made from 249.5 to 350 K (-10.6 to 170.3°F). Thirty-seven measurements were obtained. The data are presented in Table 37. The

experimental data agreed with the predictions within ±0.82% with a bias of 0.17% and an average

deviation of ±0.25%. Figure 37 compares the vapor speed of sound data with calculated values from

the MBWR equation in REFPROP 6.0.

2.4.2 R-41/744

R-4 1/744 is a Type I system with no azeotrope. The data for the R-41/744 systems cover six

isotherms from 218 to 290 K (-68 to 62°F). Thirty-seven vapor-liquid equihbrium points were

recorded. The data are presented in Table 38. In Figure 38, the deviations between the bubble-point

pressures and the predicted values are presented. The bubble-point pressure data agreed with the

13
predicted values from the Lemmon-Jacobsen model within ±1.5%. The bubble-point pressure state- point uncertainty for R-4 1/744 at temperatures from 218 to 290 K (-68 to 62°F) ranges from ±0.19

to ±0.22%. In Figure 39, the vapor compositions are compared to the predicted values of the

Lemmon-Jacobsen model. No other data sets were available for comparison. The vapor compositions

agree with the predicted values from the Lemmon-Jacobsen model within ±0.018 mole fraction of R41. The vapor-composition state-point uncertainty for the R-41/744 system at temperatures from 218

to 290 K (-68 to 62°F) ranges from ±0.006 to ±0.011 mole fraction R-41.

Isochoric (p,p,T) data were also obtained for the R-41/744 system. Measurements were taken in

both the gas and liquid phases. The data are reported in Table 39. Figure 40 shows the range of

temperatures and pressures covered by the isochoric (p,p,T) data. Figure 41 shows the comparison of

the isochoric (p,p,T) data to the predicted values from the Lemmon-Jacobsen model in REFPROP

6.0. The isochoric (p,p,T) data agreed with the predicted values from the Lemmon-Jacobsen model

within ±1.0% except for two vapor-phase points.

The R^1/744 system is a possible replacement for R-13 if the freezing point of the mixture is

significantly lower than the freezing point of pure R-744. Prior to this study, no pubhshed data were

available for the sohd-liquid equilibria of R-41/744 mixtures. Based on known behavior of similar

mixtures, we expected the freezing points of R-41/744 mixtures to fall below that of pure R-744

whose triple point temperature is 216.59 K (-69.83°F). The triple point temperature of R-41 was

estimated to be 129.82 ± 0.04 K (-226.01 ± 0.07''F), as observed in the adiabatic calorimeter. The

calculated freezing point of our equimolar mixture is approximately 186 K(-125°F), which is 13 K

(23.4°F) above the average of the triple point temperatures of R-41 and R-744.

To determine some bounds on the freezing point temperature for the equimolar R-41/744

mixture (0.499823 mole fraction R-41 (0.43591 mass fraction R-41)), we condensed the sample into

a precooled PVT ceU. In such experiments, the sample would condense to either a liquid or to a solid

after it enters the cell, depending on the initial temperature of the PVT cell. This test required only

minimal additional effort, since the PVT ceU had to be fiUed to carry out PVT experiments. We

carried out two filling experiments. The initial cell temperatures were 180 K and 170 K (-135.6 and

-153,7*F), both at a final pressure of 2 MPa (290 psia). We determined that both samples were hquid

phase by heating the samples and observing a steady rise of pressure of approximately 2 MPa/K (161

psia/°F). This technique has estabhshed that the fi^eezing point of an equimolar mixture is less than

182.5 K (-131.2°F).

We then employed an adiabatic calorimeter to accurately measure the temperature at which

sohd precipitates from a liquid sample of the R-41/744 mixture. The calorimeter bomb was filled with

a liquid sample (57.907 g) (0.1277 lb) at a final temperature and pressure of (292.68 K, 7.50 MPa

(67.14°F, 1088 psia)). This quantity of material leaves only a small vapor space above the liquid

when it was subsequently cooled. The sample was cooled rapidly to 225 K (-55°F), then slowly to 95

K (-289°F). We observed the formation of a solid phase at a temperature of 125 K (-235''F), as

indicated by a sharp break in a graph of the recorded temperatures versus times. We then heated the

sample from 95 K (-289°F) to about 115 K (-253°F), then allowed the sample to equilibrate and

14
measured the small heat leak. The heater power was then lowered to about 15% of the normal level,

and we continued to heat the sample to 135 K (-217°F). The freezing temperature was observed

where the temperature reached a plateau value for the elapsed time needed to melt the solid phase.

Three experiments were carried out. The elapsed time at the plateau temperature was about 640

minutes. The observed freezing temperatures were 124.98 ± 0.02 K (-234.72 ± 0.036°F), 124.97 ±

0.02 K (-234.74 ± 0.036°F), and 125.00 ± 0.02 K (-234.69 ± 0.036°F). Though two heater power

levels were selected, no dependence on the applied power was seen. The observed enthalpies of fusion were 2.069 ± 0.01 kJ-mol-i (22.79 ±0.11 BTU-lb-i) and 2.074 ± 0.01 kJ-mol-i (22.84 ±0.11

BTU»lb-i), after corrections for energy to heat the empty calorimeter and for energy lost to parasitic

heat leaks were applied.

The average freezing temperature of 124.98 K (-234.72°F) is 48 K (86.4°F) below the mole

fraction average of the pure component triple point temperatures, and also is 5 K (9°F) lower than

that of the lowest-melting pure substance in the mixture. This evidence implies that this binary

mixture will have at least one eutectic composition. More extensive experimental and theoretical

studies of solid-liquid transitions are needed to further elucidate this phenomenon.

The measurements on R-125/134a, R-32/125/134a, R-32/143a, R-125/143a, R-143a/134a, R32/290, R-125/290, R-134a/290, R-41, and R-41/744 were supported in part by a grant from the U.S.

Department of Energy, Office of Building Technology (Grant number DE-FG02-91CE23810:

Materials Compatibilities and Lubricants Research on CFC-Refrigerant Substitutes) administered by

the Air-Conditioning and Refrigeration Technology Institute (MCLR Project Number 660-50800).

Federal funding supporting this project constitutes 93.57% of allowable costs. Funding from the air conditioning and refrigeration industry supporting the project consists of direct cost sharing of

6.43% of allowable costs; and significant in-kind contributions. The measurements on R-32/134a and

R-32/125 were supported by a grant from the U.S. Department of Energy, Office of Building

Technology (Grant number 93CE23808.001). We would like to thank the following companies for

supplying high purity refrigerant samples: ICI for supplying R-32, DuPont and ICI for supplying R125, and AlliedSignal for supplying R-143a. We would also like to thank Michael J. Hiza for

preparing the standard mixtures used in this work and Mark O. McLinden for determining the vapor

pressure equation for R-41,

3. REFERENCES

[1] Lemmon, E.W.; Jacobsen, R.T.; Thermodynamic properties of mixtures of refrigerants R-32, R125, R-134a, and R-152a. Conference Preprint, 13th Symposium on Thermophysical

Properties, Boulder, CO, June 22-27, 1997, submitted for publication in Int. J. Thermophys.;

1997.

15
2] McLinden, M.O.; Klein, S.A. A next generation refrigerant properties database. Proc. 6th Int. Refrig. Conf. at Purdue, West Lafayette, IN, July 23-26, 1996, 409; 1996.

3] Nagel M.; Bier, K. Vapour-liquid equilibrium of ternary mixtures of the refrigerants R32, R125

and R134a. Int. J. Refrig. 18: 534; 1995.

4] Higashi, Y. Vapor-liquid equilibrium, coexistence curve and critical locus for binary HFC32/HFC-134a. Int. J. Thermophys. 16(5): 1175; 1995.

5] Widiatmo, J.V.; Sato, H.; Watanabe, K. Measurements of liquid densities of the binary HFC-32

+HFC-134a System. Huid Phase Equilibria 99: 199; 1994.

6] Fujiwara, K.; Momota, H.; Noguchi, M. Vapor-liquid equilibria of HFC-32 mixtures. Proc. 13th

Japanese Sym. Thermophysical F*roperties A116: 61; 1992.

7] Defibaugh, D.R.; Morrison, G. Interaction coefficients for 15 mixtures of flammable and non- flammable components. Int. J. Refr. 18(8): 518; 1995.

8] Weber, L.; Defibaugh, D. R. The virial coefficients of five binary mixtures of fluorinated

methanes and ethanes. Int. J. Thermophys. 15(5): 863; 1994.

9] Widiatmo, J.V.; Sato, H.; and Watanabe, K. Bubble-point pressures and saturated-liquid densities

for binary HFC-32 + HFC-125 system. High Temperatures - High Pressures 25: 677; 1993.

10

11

12

13

14

15

16

17

19

20

Higashi, Y. Vapor-liquid equilibrium, coexistence curve and critical locus for binary HFC32/125. Proc. 19th Int. Congress of Refrigeration, Netiieriands IVa: 297; 1995.

Higuchi, M.; Higashi, Y. Measurements of the vapor-liquid equilibrium for binary R125/134a

mixture. Proc. 16th Japanese Symp. Thermophysical Properties, 16: 5; 1995.

Higashi, Y. Vapor-liquid equihbrium and critical point for the R32/125/134a mixture. Presented

at the AIChE Spring Meeting, New Orleans, LA; February 1996.

GiUis, K.A.; Hurly, J.J.; Schmidt, J.W. Equation of state and ideal-gas heat capacity of a gaseous

mixture of 1,1,1,2-tetrafluoroethane, pentafluoroethane, and difluoromethane. Int. J. Thermophys. 18(3): 655; 1997.

Takashima, H.; Higashi, Y. Vapor-liquid equilibrium, coexistence curve and critical locus for

binary HFC-125/143a. Proc. 16th Japanese Symp. Thermophysical Properties, 16: 9; 1995.

Nagel M.; Bier, K. Vapour-liquid equilibrium of ternary mixtures of the refrigerants R125,

R143a and R134a. hit. J. Reftig. 19, 264 (1996).

Kleiber, M., Vapor-Uquid equilibria of binary refrigerant mixtures containing propylene or R134a, Fluid Phase Equflibria 92: 149; 1994.

Jadot, R.; Frere, M. Simulation of liquid-vapor equilibria of hydrofluorocarbon (HFC) mixtures

by the UNIFAC model. High Temperatures - High Pressures 25: 491; 1993.

Oi, T.; Shulman, J.; Popowicz, A.; Ishida, J. Vapor pressure isotope effects in liquid methyl

fluoride. J. Phys. Chem. 87: 3153; 1983.

Van Poolen, L.J.; Niesen, V.G.; Holcomb, CD.; Outcalt, S.L. Critical densities from coexisting

density data: application to refrigerants R22, R134a, and R124. Fluid Phase Equilibria 97: 97;

1994.

Bominaar, S. A. R. C; Biswas, S. N.; Trappeniers, N. J.; Ten Seldam, C.A. (p,Vm,T) Properties of

mclhyl fluoride in the (gas+liquid) critical region. J. Chem. Thermodynami. 19: 959; 1987.

16
[21] Biswas, S. N.; Ten Seldam, C. A.; Bominaar, S. A. R. C; Trappeniers, N. J. Liquid-vapor

coexistence curve of methyl fluoride in the critical region. Fluid Phase Equilibria 49: 1; 1989.

17
6. FIGURES

18
a>Q.Oh. *<oNnc

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cu

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ro '^ ^^ ^ CO L.

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Q.

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to 2 JZ

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CD

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19
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Temperature - ** F

80.3 170.3

_l

260.3

u

1o

a

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9

a.

mo

4-

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+

X

V

o

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-5

H K

++ +

+ +

200

Nagel and Bier [3] Higashi [4] Widiatmo et al, [5] Fujiwara et al. [6]

+ .

+

+

+

250

9

4 +

S ? * 7 K „

+ + <<

< Defibaugh and Morrison [7] This Work, Measured Composition Data

'

1 300

Temperature - K

350 400

Rgure 2. Comparison of Bubble-Point Pressures for the R-32/134a System; the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0.

20
Temperature - °F

-99.7

0.05

CM

CO

I O

o

(0

Ia.X

>

0.04-

0.03-

0.02-

0.01-

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-0.03 -

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-0.05

-9.7 _J_

80.3

I

170.3

I

260.3

X Higashi [4] O Fujiwara et al. [6] This Work, Measured Composition Data + Nagel and Bier [3]

*++ * *

+

200

o *x^

—

I—

250

1 300

Temperature - K

1 350 400

Figure 3. Comparison of Vapor Compositions for the R-32/1 34a System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0.

21
-9.7

(0o

COo

'ca. X

Q.

4-

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° -2

-3-

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Temperature

80.3

X

This Work, Liquid Densities

Tliis Work, Vapor Densities

< Weber and Defibaugh [8]

2 4

<

<

<

300

I 350

Temperature - K

170.3

400

Figure 4. Connparison of Densities for the R-32/134a System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

22
Temperature -*'F

(0o

(0o

CL.'q. X

9

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X

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++

200

-9.7 _J

80.3

_J

170.3

I Nagel and Bier [3] Fujiwara et al. [6] Widiatmo et al. [9] O

Higashi[10] g

-I- +

-I- Ttt

260.3

< Defibaugh and Morrison [7] + This Work, Measured Composition Data —

I—

250 300

Temperature - K

—]—

350 400

Figure 5. Comparison of Bubble-Point Pressures for the R-32/1 25 System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0.

23
-99.7

CNJ CO6

(QU

>-

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Figure 6.

-9.7

Temperature - °f

80.3 170.3 260.3

X Higashi[10]

O Fujiwara et al. [6] + Nagel and Bier [3] This Work, Measured Composition Data

—

I—

250 300

Temperature - K

350 400

Comparison of Vapor Compositions for the R-32/125 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

24
-9.7

Temperature -°F

80.3

_J

170.3

I

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too

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<

This Work, Liquid Densities

This Work, Vapor Densities Weber and Defibaugh [8] ^-4

g

a

f

D

—I—

300 350 400

Temperature - K

Figure 7. Comparison of Densities for the R-32/1 25 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

25
Temperature - °F

-99.7 -9.7 J

-80.3

I

170.3

I

260.3

4-

3-

+ This Work, Measured Composition Data

+ Nagel and Bier [3] • This Work, Standard Mixture Data X Higuchi and Higashi [11] x

X

X

400

Temperature - K

Figure 8. Comparison of Bubble-Point Pressures for the R-1 25/1 34a System; the

Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0.

26
-99.7

0.05

CM

I O

c

o1

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+

200

-9.7

I

Temperature - °F

80.3

I

170.3

_J

^' This Work, Measured Composition Data X Higuchi and Higashi [11] + Nagei and Bier [3]

+ +

+ + -I- +

^--^

^ X

X

-I250

1 300

Temperature - K

350

260.3

400

Figure 9. Comparison of Vapor Compositions for the R-1 25/1 34a System; tfie Baseline

is from the Lemmon-Jacobsen Model in REFPROP 6.0.

27
_o

(0o

a.

COu

9-

Q.X

a> Q. O

-9.7

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<

3-

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1 -

0- —

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Temperature - °F

-80.3 170.3

J. This Work, Liquid Densities

Tiiis Work, Vapor Densities Weber and Defibaugli [8]

M—Q.

D

^B

IFd

—

r300 350 400

Temperature - K

Figure 1 0. Connparison of Densities for the R-1 25/1 34a System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

28
-99.7 -9.7 _L_

Temperature -°F

80.3

I

170.3 260.3

4-

3-

2-

+ This Work, Measured Composition Data

+ Nagel and Bier [3] X Higashi [12] • This Work, Standard Mixture Data

300

Temperature - K

r

350 400

Figure 1 1 . Comparison of Bubble-Point Pressures for the R-32/1 25/1 34a System; the

Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0.

29
-99.7

0.05

CO6

X

c

o

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>

0.04 H

0.03

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0.01 -

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+

X

200

-9.7L

Temperature - °F

80.3

I

170.3

l_

This Work, Measured Composition Data Nagel and Bier [3] Higashi [12]

+

• \*

+

+ X +

*

X

JI 250 300

Temperature

—I—

350

260.3

400

Figure 1 2. Comparison of Vapor Compositions for the R-32/1 25/1 34a System; the

Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0.

30
au

a

o

9.

a.X

•o

-99.7

4-

3-

2-

1 -

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200

-9.7

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Temperature -"F

80.3

I

170.3

This Work, Liquid Densities

D This Work, Vapor Densities

< Gillis et al. [13] B

B

n

*^* t *^'^i^*^

250

1 300

Temperature - K

350

260.3

400

Figure 13. Comparison of Densities for the R-32/1 25/1 34a System; the Baseline is from the

Lemmon-Jacobsen Model In REFPROP 6.0.

31
-54.7 -9.7

I

Temperature -"F

35.3 80.3

± ±

125.3

L_

170.3

(0u

(0 a."

'q. X

o

•o

4-

3-

2-

1

-1-

-2-

-3-

-4-

-5-

O Fujiwara et al. [6]

This Work, Measured Composition Data O

e

t I

—

T275

"1—

300

1

225 —r250

Temperature - K

325 350

Figure 14. Comparison of Bubble-Point Pressures for the R-32/143a System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0.

32
Temperature - °F

8.3 0.05

0.04 H ^ 0.03 H

X 0.02 H c

oU 0.01 ^

-2 0.00

s -0-01 H

o

'q. -0.02 H X

(D ^ -0.03 H

-0.04 -

-0.05 260

44.3

_JI

8

-8-

—

I—

280

80.3

I 116.3 152.3

I

188.3

O Fujiwara et al. [6]

This Work, Measured Composition Data

I 300 320 340 360

Temperature - K

Figure 1 5. Comparison of Vapor Compositions for the R-32/1 43a System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

33
-54.7 -9.7

Temperature -°F

35.3 80.3

__J L_

125.3

l_

170.3

a

u

4-

3-

2-

2: 1

(0o

Q.

Ia.

Q. O

o^ -2-

-1 -

-3-

-4-

-5

This Work, Liquid Densities

D Tliis Work, Vapor Densities

% $

"TT

n

JU_l

P

D

325

I

t

g

225 250 275 300

Temperature - K

350

Figure 16. Comparison of Densities for the R-32/143a Systenn; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

34
Temperature -°F

170.3

COo

Q."5

(0o

a.

Ia.X

o

Q.O^ -2-

-1-

-4H

-5

Takashima and Higashi [14] This Work, Standard Mixture Data

This Work, Measured Con^iposition Data

Nagel and Bier [1 5] —

225 250 275 300 325 350

Temperature - K

Figure 1 7. Comparison of Bubble-Point Pressures for the R-1 25/1 43a System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0.

35
8.3 0.05

6u_X

c

onto

o

a.X

<s

0.04-

0.03-

0.02-

0.01 -

0.00

-0.01 -

-0.02 -

-0.03 -

-0.04-

-0.05

-f

260

44.3

_JI

Temperature - °F

80.3 116.3

J. J. X Takashima and Higashi [14] This Work, Measured Composition Data

-f- Nagel and Bier [15]

X

+

280

<

*

-*

"~T~" 300

1

+

152.3

L_

188.3

320

Temperature - K

—

r340 360

Figure 18. Connparison of Vapor Compositions for the R-1 25/1 43a System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0.

36
Temperature -°F

-99.7 -9.7

40'

30-

0.

2 20-

Q.

80.3

_l

170.3

I ^ This Work, Isochoric (p.p.T) Liquid Density Data

260.3

- 4351

This Work, Isochoric (p, p,T) Gas Density Data mHHmm mfflfflfflfflBBEB

200 250 300

5802

350 400

Temperature - K

Figure 19. Range of Measured Tennperatures and Pressures for Isochoric (p, p,T) Data for a Mixture

of R-125/143a with x(R-125) = 0.49996 Mole Fraction (0.58812 Mass Fraction).

37
(0

o

'q. X

(D O

o

-99.7

1.4- 1.2-

1.0-

0.8-

0.6-

0.4-

0.2-

0.0 -0.2-

-0.4-

-0.6-

-0.8-

-1.0-

-1.2-

-1.4200

<

m

-9.7

Temperature -"F

80.3 170.3 260.3

This Work, Near Saturation (p,p,T) Liquid Density Data

This Work, Near Saturation (pp,T) Vapor Density Data Weber and Defibaugh [8]

This Work, Isochoric (p,p,T) Liquid Density Data

This Work, Isochoric (p,p,T) Gas Density Data

1 1 250 300

Temperature - K

350 400

Figure 20. Comparison of Densities for the R-1 25/143a System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

38
(0

Ia.X

a

•oo

-99.7

4-

3-

2-

a

," n

1 -

-2-

-3-

-4-

200

-9.7

_l

Temperature - °F

80.3

I

170.3 260.3

+

This Work, Measured Composition Data

This Work, Standard Mixture Data Nagel and Bier [15]

t H

+ •

—

I—

250

1 300

Temperature - K

—

I—

350 400

Figure 21 . Connparison of Bubble-Point Pressures for the R-1 43a/1 34a System; the Baseline

is from the Lemmon-Jacobsen Model in REFPROP 6.0.

39
-99.7

0.05

CO 6

c

o1

(0

a.X

a> >-

0.04-

0.03-

0.02-

0.01 -

0.00

-0.01 -

-0.02-

-0.03 -

-0.04-

-0.05200

-9.7

Temperature - "F

80.3

I

170.3

«

This Work, Measured Composition Data + NagelandBier[15]

—

I—

250

+

*.+ +

*

1 300

Temperature - K

350

260.3

400

Figure 22. Comparison of Vapor Compositions for the R-1 43a/1 34a System; the

Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0.

40
-9.7 5

4'

3-

2-

1-

Temperature - °F

80.3

I

170.3

(0o

(0o

Io.X

<a a.

•o

This Work, Liquid Densities

n This Work, Vapor Densities

-4-&.

-1 -

-2-

-3-

-4-

-5 1 300

Temperature - K

250 350

Figure 23. Comparison of Densities for the R-143a/134a System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

41
Temperature -"F

m

u

a

o

a.

IX

o

•o

-99.7 -9.7 80.3 170.3 26

5- 1 1 1 «

4- t

3- +

+ 4

2-

1-

*

+

+*

\j - +

+

-1 - * t

*

-2-

-3- *

t

+

-4- + This Work Measured Composition Data

-5-

1 *1 1 200 250 300

Temperature - K

350 400

Figure 24. Comparison of Bubble-Point Pressures for the R-32/290 System; the Baseline is from

the Lemmon-Jacobsen Model in REFPROP 6.0.

42
Temperature - °F

CMOU-X

c

o

(0

-99.7 -9.7 80.3 170.3 26

n f\K _ 1 1 1 U.UO ~ + This Work, Measured Composition Data 0.04-

0.03- ^

0.02- % *

0.01-

n nt\ _ +

*

* *

* u.uu - * t

-0.01 -

-0.02 -

-0.03 - +

** <

+

-0.04 - +

n fm - +

-u.uo 1 1 1 200 250 300 350 400

Temperature - K

Figure 25. Comparison of Vapor Compositions for the R-32/290 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

43
O

(QO

fQ.X

<D

•O

o

-9.7 10

6-

4-

2-

-2-

o -4-

-6-

-8-

-10250

Temperature - °F

-80.3 170.3

This Work, Liquid Densities

This Work, Vapor Densities

Ji 4

300

Temperature - K

350

Figure 26. Comparison of Densities for the R-32/290 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

44
Temperature -"F

c _

>9.7 -9.7 80.3

1

170.3 26

r

4- * This Work, Measured Composition Data

3- +

o

ns

Ij

(0u

2-

1-

0-

-1-

-2-

-3-

1

t -M. +

1exX

(D

•1

1

+ *

* t

* *

-4-

-5-

1

+

200 250 300

Temperature - K

350 400

Figure 27. Comparison of Bubble-Point Pressures for the R-1 25/290 System; the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0.

45
CMOLLX

c

oI

o

(0

IQ.X

>-

-99.7 0.05

0.04-

0.03-

0.02-

0.01-

0.00

-0.01 -

-0.02 -

-0.03 -

-0.04 -

-0.05 200

-9.7

Temperature - °F

80.3

L

170.3

_I

260.3

This Work, Measured Composition Data

* * * * ^ • *

t I * *

250 __ ^ _„

300

Temperature - K

350 400

Figure 28. Comparison of Vapor Compositions for the R-1 25/290 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

46
Temperature - °F

tao

'~oa

o

Q.X

o

•o

c _

9.7 -80.3

1

17C D

~

This Work, Liquid Densities

4- n This Work, Vapor Densities

3- fe

2-

D

1 -

U

"

-1-

-2-

D

>

-3- D

-4-

-5-

1 250 300 350

Temperature - K

Figure 29. Comparison of Densities for the R-1 25/290 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

47
-99.7 15

12-

9-

u

a,

CO a,

a.X

9

•o

o

-3-

-6-

-9-

-12-

-15

200

+

>

-9.7 _1_

Temperature -°F

80.3

I

170.3

I This Work, Measured Composition Data

Jadot and Frere [1 7] Kleiber[16]

^

^

>

> *

t i li

-T250

T"

300

Temperature - K

—

I—

350

260.3

400

Figure 30. Comparison of Bubble-Point Pressures for the R-134a/290 System; the Baseline Is from the Lemmon-Jacobsen Model In REFPROP 6.0.

48
Temperature - °F

-99.7 -9.7

CO

I OU-X

c

g

XS

CO

III _2o

(0

IOlX

0) >

0.10'

0.08-

0.06-

0.04-

0.02-

0.00

-0.02 -

-0.04 -

-0.06 -

-0.08 -

-0.10 200

Figure 31

+

>

80.3

I

170.3

_J

260.3

^ ^

*

M.

*

%

I

g M n

This Worl<, Measured Composition Data

Kleiber[16]

Jadot and Frere [1 7] —

I—

250

I 300

I 350

Temperature - K

400

Comparison of Vapor Compositions for the R-134a/290 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

49
Temperature - °F

(0o

«

O

Q.

'q. X

a.

•o

o

-9.7 20'

15-

10-

5-

-5 250

-80.3 _J

This Work, Liquid Densities

n This Work, Vapor Densities

D

D

n

>

I

300

Temperature - K

D

D

D

D

D

B

D

$

$

170.3

350

Figure 32. Comparison of Densities for the R-134a/290 System; the Baseline is from the

Lemmon-Jacobsen Model in REFPROP 6.0.

50
Temperature -"F

-279.7 -189.7

1-

nu

(0 Q."Q.X

a>

•o

-1 -

150

-99.7 -9.7

I

80.3 170.3

** O**

t^S^^^<>^'^<^^^'(^ii>%*^^t^^^

This Work, Vapor-liquid Equilibrium Apparatus

> Oietal.[18] * This Work, Isochoric (p, p,T) Apparatus

I 200 250

Temperature - K

300 350

Figure 33. Comparison of Vapor Pressures for R-41 ; the Baseline is from Equation (1 ).

51
(0 CL

-279.7

40

30-

20-

10-

100

-189.7

I

-99.7 _l

Temperature -°F

-9.7 80.3

L_

170.3

_l

260.3

This Work, Isochoric (p, p,T) Liquid Density Data ^ ^

5802

A^ ^ ^

A ^ Z

^

^ ^

V V V V V V

ffl This Work, Isochoric (p, p,T) Gas Density Data

- 4351

2901

- 1450

a.

150 200 250

Temperature - K

300 350 400

Figure 34. Range of Measured Temperatures and Pressures for Isochoric (pp.T) Data for R-41.

52
_oa

o

(0u

o.

IX

a>

•oo

-279.7

0.5'

0.4-

0.3-

0.2-

0.1-

0.0

-0.1 -

-0.2-

-0.3-

-0.4-

-0.5

m

100

-189.7 -99.7

I

Temperature -°F

-9.7 80.3

L_

170.3 260.3

This Work, Extrapolation of Isochoric (p,p,T) Data to Saturation

This Work, Near Saturation (p,p,T) Liquid Density Data

This Work, Isochoric (p,p,T) Liquid Density Data

This Work, Isochoric (p,p,T) Gas Density Data

400

Tennperature - K

Figure 35. Comparison of Densities for R-41 ; the Baseline is from the MBWR Equation of State in REFPROP 6.0.

53
I

i

9a.X

oO

•o

o

-279.7 10

5-

-5-

10-

-15'100

V

-189.7 __J

Temperature -°F

-99.7 -9.7 J. J.

This Work, Cy

This Work, C^at

T

150

7 V V

200

"1—

250

80.3 170.3

V V V

vV

T

300 350

Temperature - K

Figure 36. Comparison of the Isochoric Heat Capacity Data at Saturation and in the Single-Phase

Liquid Region for R-41 ; the Baseline is from the MBWR Equation of State in REFPROP

6.0.

54
Temperature -°F

80.3

Temperature - K

170.3

350

Figure 37. Comparison of the Vapor Speed of Sound Data for R-41 ; the Baseline is from the

MBWR Equation of State in REFPROP 6.0.

55
Temperature -"F

-99.7

(0u

a.

(0u

'a. X

Oo

80.3

300

Temperature - K

Figure 38. Comparison of Bubble-Point Pressures for the R-41/744 System; the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0.

56
Temperature - °F

-99.7 0.05

6

j2o

0.04-

0.03-

0.02-

0.01-

0.00 o

a -0.01

X

>-^ -0.02 H

-0.03 -

-0.04 -

-0.05 200

-9.7 J

80.3

1 t

This Work, Measured Composition Data

Figure 39.

250 300

Temperature - K

Comparison of Vapor Compositions for the R-41/744 System; the Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0.

57
Temperature -"F

-189.7

40'

sons a.

2 20-

10-

150

m

-99.7 -9.7 J

80.3

-J

170.3

This Work, Isochoric (p, p,T) Liquid Density Data

This Work, Isochoric (p, p,T) Gas Density Data

1 ' 200

—r250 300 —

r350

260.3

- 4351

5802

400

Temperature - K

Figure 40. Range of Measured Temperatures and Pressures for Isochoric (p, p,T) Data for a

Mixture of R-41/744 with x(R-41 ) = 0.49982 Mole Fraction (0.43591 Mass Fraction).

58
Temperature -°F

-189.7

1 - a

o

(0

IQ.X

0)

•o

-1-

-2 S

150

-99.7 -9.7 80.3

I 170.3 260.3

This Work, Isochoric (p, p,T) Liquid Density Data

*^'''Si^i«^<P»^»*^'^^^^!^^

This Work, Isochoric (p, p,T) Gas Density Data —

T—

200

I 250I 300 350 400

Temperature - K

Figure 41 . Comparison of Isochoric (pp,T) Data for the Equimolar Mixture of R-41/744; the

Baseline is from the Lemmon-Jacobsen Model in REFPROP 6.0.

59

APPENDIX A:

TABLES OF THERMOPHYSICAL PROPERTY DATA

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APPENDIX B:

EXPERIMENTAL APPARATUS

B-l
Five apparatus were used to acquire the data for this project. These apparatus will be described

in this section.

B.l Vapor-Liquid Equilibrium Apparatus

A schematic diagram of the vapor-liquid equilibrium apparatus used in this project is presented

in Figure Bl. The apparatus is designed to allow the liquid and vapor of a fluid to come to equilibrium. Then the equilibrium pressure and temperature are measured. Also, the liquid and vapor

phase compositions are measured by flashing a small sample of either the liquid or the vapor into a low pressure helium carrier gas and using a gas chromatograph to analyze the sample. The main part

of the apparatus consists of an equilibrium cell, two recirculating pumps, and a platinum resistance

thermometer that are contained in a Dewar filled with silicone oil. The platinum resistance

thermometer has been calibrated on the ITS-90. The pressure transducer and feed system are contained in an insulated box with an independent temperature control system to maintain it at a temperature above the temperature of the Dewar. The pressure transducer is an oscillating quartz

crystal pressure transducer calibrated in our laboratory.

One or more mixture components are introduced into the system using the feed system. The

liquid and the vapor are continuously circulated to ensure that the system is well mixed. The liquid

and vapor samples are withdrawn through capillary tubing to gas chromatograph sampling valves that

are used to inject a smaU quantity of the sample into a helium stream. The helium stream carries the

sample to the gas chromatograph where the mixture components are separated. The procedure for

calibrating the gas chromatograph is outlined in Appendix C,

B2 Bubble-Point and Near-Saturation (p,p,T) Apparatus

The apparatus used to measure the bubble-point pressures and near-saturation (p,p,T) data is

similar to the vapor-liquid equilibrium apparatus described above. The main part of the apparatus

consists of an equilibrium ceU, two pumps, and two vibrating tube densimeters. These components are housed in an oven which is used to control the temperature of the system. Figure B2 shows a schematic diagram of this system. The liquid pump is located outside of the system and is used when

the system is above room temperature.

The system can either be filled with a standard mixture of known composition, or the liquid

composition can be measured. The liquid and the vapor are continuously circulated to ensure that the

system is weU mixed. Once the system is equilibrated, the temperatures of the main cell and both

densimeters are measured, the pressure of the system is recorded, and both the liquid and vapor

densities are recorded. The composition of the vapor phase is calculated from the vapor-liquid

equilibrium data from the VLB apparatus described above. It is assumed that the vapor composition

is that of the vapor in equilibrium with the liquid which is at the temperature and pressure of the main

ceU. The reason the vapor composition is not measured directly is because the temperature gradients

in this apparatus adversely affect the vapor composition determination.

B-2
B.3 Isochoric (p,p,T) Apparatus

An isochoric technique was employed to measure the single-phase densities in this study. In

this method, a sample of fixed mass in confined in a container of nearly fixed volume. The volume

of the container is accurately known as a function of pressure and temperature. The temperature is changed in selected increments, and the pressure is measured at each temperature, until the upper

limit of either temperature (400 K (260°F)) or pressure (35 MPa (5100 psia)) was attained. When an

isochore is completed, that is, after the upper temperature or pressure limit of the run has been

reached, the sample is condensed into a light-weight stainless steel cylinder which is immersed in

liquid nitrogen. When the (p,p,T) cell and its connecting capillary have been heated to about 20 K

(36°F) above the critical temperature of the sample, the cylinder is sealed, warmed to ambient

temperature, and weighed. The density of the test fluid is then determined from a knowledge of the

cell volume and of the mass difference of the steel cylinder before and after trapping the sample.

Allowances to account for the noxious volumes in the system, such as those of the capillaries and the

pressure gauge, are made for each point. A small adjustment to the apparent sample mass was made

to account for the change in the atmospheric buoyant force acting on the steel cylinder. The density

of the sample fluid is then the quotient of the mass of the sample and the volume of the cell at each

pressure and temperature.

The sample cell, shown in Figure B3, is a cylindrical piece of electrolytic tough pitch copper

containing a cavity with a volume of approximately 28.5 cm^ (1.74 in3). It is suspended inside an

evacuated cryostat from a thin-walled stainless steel tube used for reflux cooling. High resistance wire

wound tightly around the cell is used to heat the cell. The cell temperature is determined with a platinum resistance thermometer (calibrated at NIST relative to the IPTS-68, with temperatures

converted to the ITS-90) embedded in a small well at the top of the cell. An ultrastable current source

supplies the thermometer with a current of 2 mA. Errors caused by steady-state thermal and contact EMF's are minimized by averaging voltages measured for opposite directions of current flow. The

temperatures were controlled and reproduced within 1 mK (0.(X)18°F). The total uncertainty in the

temperature ranged from 10 mK at 100 K to 30 mK at 400 K (0.018°F at -279.7°F to 0.054°F at 260°F).

Pressures are measured by reading the period of vibration, averaged over 10 s, of an oscillating

quartz crystal transducer which is cormected to the sample cell through a fine diameter (0.2 mm i.d. (0.008 in i.d.)) capillary. Since the frequency of the transducer varies with temperature, the

transducer has been anchored in an insulated aluminum block controlled at 333.15 ± 0.05 K (139.98

± 0.09°F). The transducer has been calibrated with an oil-lubricated piston gauge, accurate within

±0.01%. Calibrations have demonstrated that the transducer is extremely stable over long periods of

time. Changes of less than 0.003% were observed over one year. The expanded uncertainty in the

pressure measurements is approximately 0.01% for pressure greater than 3 MPa (435 psia). but

increases to 0.05% at low pressures (1 MPa (145 psia) and lower) as a result of the transducer

resolution, of the fluctuations in the temperature of the pressure transducer, and of the occasional

hysteresis in the vibrational frequency of the quartz element.

B-3
B.4 Adiabatic Constant Volume Calorimeter

The principal components of the calorimeter used for these measurements are shown in Figure

B4. A spherical bomb contains a sample of well-established mass. The volume of the bomb,

approximately 73 cm3 (4.6 in3), is known as a function of temperature and pressure. A platinum

resistance thermometer is attached to the bomb for the temperature measurement. Temperatures are reported on the ITS-90, after conversions from the original calibration on the IPTS-68. Pressures are measured with an oscillating quartz crystal pressure transducer with a to 70 MPa (0 to 10000 psia)

range. Adiabatic conditions are ensured by a high vacuum (3 x 10-3 pa (4.4 x 10-^ psia)) in the can

surrounding the bomb, by a temperature-controlled radiation shield, and by a temperature-controlled

guard ring which thermally anchors the filling capillary and the lead wires to the bomb.

For heat capacity measurement, a precisely determined electrical energy (Q) is applied and the

resulting temperature rise (AT = T2 - Ti) is measured. We obtain the heat capacity fi-om

where U is the internal energy, Qo is the energy required to heat the empty calorimeter, Wpv is the

change-of-volume work due to the slight dilation of the bomb, and n is the number of moles

contained in the bomb. In this work, the bomb was charged with sample up to the (p,T) conditions of

the highest-density isochore. The bomb and its contents were cooled to a starting temperature in the

single-phase liquid region. The measurements were performed with increasing temperature until

either the upper temperature (345 K (161°F)) or pressure limit (35 MPa (5076 psia)) was attained. At

the completion of a run, a small part of the sample was cryopumped into a lightweight cylinder for

weighing. The next run was started with a smaller density. A maximum of eight runs was measured

with one filling of the bomb. When the runs were completed, the remaining sample was discharged

and weighed. A series of such nms from different fillings completes the investigation of the (p,T,Cv)

surface.

B^ Spherical Resonator Speed of Sound Apparatus

Accurate data of the speed of sound of a gas are obtained from measurements of the resonant

frequencies produced in a stainless steel spherical cavity containing the gas. Measurements of several

radial modes are taken at a very stable temperature and pressure. The speed of sound is computed

from the frequency at each of the modes, and an average is taken. The sphere radius used in the

computation of the speed of sound is found from similar measurements of the radial modes using

argon. Highly accurate properties of argon gas, including the speed of sound, are found in the work

of M. R. Moldover, et al. [Moldover, M.R.; Mehl, J.B.; Greenspan, M.J. Accoust. Soc. Am. 79: 2;

B4
1986]. Also included in this reference is a detailed description of the theory of operation of the

spherical resonator. The spherical steel cavity is mounted in a pressure vessel which also contains the gas. This

vessel is surrounded by a temperature shield and is housed in a cryostat which provides cooling for

lower temperatures. The pressure vessel has been designed for operation at pressures up to 7 MPa

(1000 psia) and temperatures from 249 K to 350 K (-11.5 to 170.3°F). Figure B5 shows a schematic

diagram of this apparatus.

B-5
Temperature

Controlled __

Enclosure

Refrigerator

Heater Power

Supply

Gas

Chromatograph

Figure B1. Schematic Diagram of the Dynamic Phase Equilibrium Apparatus.

B-6
Gas

Chromatograph

Figure B2. Schematic Diagram of Bubble-Point and Near-Saturation

(p,p,T) Apparatus.

B-7
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Cold Wall

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Tube

Cold Ring

Reflux Tube

Outer Refrigerant

Dewar

Shield

Vacuum Can

Calorimeter

Case

Thermometer

Figure B4. Schematic Diagram of the Calorimeter.

B-9
M NjGas

DATA ACQUISITION AND TEMPERATURE CONTROL INSTRUMENTS

Receiving microphone

Vacuum

Radiation shields

Pressure/Fill line

Transmitting microphone

Heaters

Figure B5. Schematic Diagram of Spherical Resonator Speed

of Sound Apparatus.

B-io
APPENDIX C:

ESTIMATES OF STATE-POINT UNCERTAINTIES

c-i
A state point is a group of parameters that uniquely specifies a point in a system such that there

are no degrees of freedom. For a binary system, if there are two phases in equilibrium and the

temperature and liquid composition (the independent parameters) are known, the state point is specified. The bubble-point pressure, vapor composition, densities, viscosity, surface tension, and

many other dependent properties are functions only of the temperature and liquid composition along

the two-phase boundary. Therefore, any uncertainties in the temperature and liquid composition

directly affect the uncertainty of the dependent properties. Unlike the measurement of an individual

parameter where the uncertainty is a function of only the calibration and measurement technique, the

state-point uncertainty also depends on the relations among the measured parameters.

The bubble-point pressure state-point uncertainty is defined as the uncertainty of the bubblepoint pressure given the uncertainty of the liquid composition measurement, the uncertainty of the

temperature measurement, the change of the bubble-point pressure with temperature and

composition, and the uncertainty in the pressure measurement. The bubble-point pressure state-point

uncertainty can be estimated using eq. (CI).

±AP„= ±

Bx )t

Ax -^1 -ATI-H |AP|

100

P

(CI)

where

P(j = bubble-point pressure,

A P(y = bubble-point state-point pressure uncertainty in percent,

X = measured liquid composition.

Ax = liquid-composition uncertainty,

T = measured temperature,

AT = temperature measurement uncertainty,

P = measured pressure, and

A P = pressure measurement uncertainty.

The partial derivatives can be calculated using either the experimental data or an equation of state. The uncertainty of the vapor-composition state point is a function of temperature, liquid

composition, and the interdependence of the parameters. It can be determined with a calculation

similar to that of the bubble-point pressure state-point uncertainty. The expression for the vaporcomposition state-point uncertainty is shown in eq. (C2).

±Ay,= ±

5yo

6x )t

Ax +

6y^

Ut

AT + |Ay|

(C2)

C-2
where

y(j = equilibrium vapor composition,

A Yct = vapor-composition state-point pressure uncertainty,

X = measured liquid composition,

Ax = liquid-composition uncertainty,

T = measured temperature,

AT = temperature measurement uncertainty,

y = measured vapor composition, and

Ay = vapor-composition measurement uncertainty.

The uncertainty of the density state point is a function of temperature, composition, and the

interdependence of the parameters. It can be determined with a calculation similar to that of the

bubble-point pressure state-point uncertainty. The expression for the density state-point uncertainty

is shown in eq. (C3).

±Ap.= ±

'5po Ac

5T

AT + |Ap|

100 (C3)

where

p<j = state-point density,

A p<j = density state-point pressure uncertainty in percent,

c = measured liquid or vapor composition,

Ac = liquid or vapor composition measurement uncertainty,

T = measured temperature,

AT = temperature measurement uncertainty,

p = measured density, and

A p = density measurement uncertainty.

C-3

APPENDIX D:

GAS CHROMATOGRAPH CALIBRATION PROCEDURES

D-l
The gas chromatograph calibration procedure involves six steps. The first step is the

preparation of two or three standard mixtures for each system. These mixtures are used to determine

which GC column to use, to refine the calibration, and to check the sampling system for adsorption or mixing problems. The number of mixtures needed is dependent on the type of GC column used to separate the two components. A short-packed column usually requires two standard mixtures. A

capillary column or long-packed column requires three or more mixtures due to the complex

behavior of adsorption and desorption along the length of the column.

In tlie second step, the standard mixtures are used to determine which GC column best separates

the two components and what are the appropriate settings on the GC. Table Dl summarizes the GC

columns and settings used for each of the systems in this report.

In the third step of the procedure, density-versus-area data are obtained for the pure

components from the GC. A specific volume is chosen for injection of the sample into the GC. This

volume is filled at different pressures of the pure fluid in a region where the virial equation for the

pure fluid can accurately predict the density. The pressure and temperature of the standard volume

are recorded before each injection. The density of the injected gas is calculated from the virial

equation of state. The fourth step of the procedure characterizes any peak interaction or sample size dependence

by injecting different quanfities of the standard mixtures into the GC. The same sample volume used

for the pure fluids is filled to different pressures with the standard mixtures. The area-percent-versussample-size data are recorded.

The fifth step is determining the correct calibration equation to use with the calibration data.

For short-packed columns the area percent of the GC peaks should be independent of the amount of

the standard mixture injected. The calibration equation for this type of column is based on

combining the area-versus-density data of the two pure components. Usually the area-versus-density

relationships for each of the pure components can be represented by a simple quadratic equation.

The standard mixture data is used to optimize the fit and to check the accuracy of the equation.

A capillary column or long-packed column shows a dependence of area percent on the sample

size injected. The calibration equations for these systems are more complicated than for short-packed

columns because they include a size dependence as well as a composition dependence that requires

all of the mixture and pure component data to be used in the fit. The procedure is more complicated

than the procedure for a short-packed column, but is necessary in order to obtain accurate results. The final step checks the sampling system for losses due to adsorption or inhomogeneous

sampling. The apparatus is filled with one of the standard mixtures, and several samples are analyzed

using the sampling system. The composition calculated from the GC should be the same as the

composition of the standard mixture. If the composition is not the same as the standard mixture, one

of the components may be adsorbing to the walls of the sampling system or the sample is not well

mixed before it reaches the GC.

The GC calibration is only accurate within Uie peak area range of its calibration. If the samples

withdrawn from the system are too large or too small, the calculated compositions wiU be in error.

D-2
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calibration.

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APPENDIX E:

REPRINTS OF PAPERS RELATED TO THE LEMMON-JACOBSEN

MODEL [1] AND REFPROP 6.0 [2]

E-l
Thermodynamic Properties of Mixtures of

Refrigerants R-32, R-125, R-134a, and R-152ai

E.W. Lemmon2,3,4 and R.T Jacobsen2

1 Paper presented at the Thirteenth Symposium on Thermophysical Properties, June 22-27, 1997,

Boulder, Colorado, U.S.A.

2 Center for Applied Thermodynamic Studies, University of Idaho, Moscow, ID 83844-1011

3 Current Address: Physical and Chemical Properties Division, National Institute of Standards and

Technology, 325 Broadway, Boulder, CO 80303

4 To whom correspondence should be addressed.

E-2
ABSTRACT

A mixture model explicit in Helmholtz energy has been developed which is capable of

predicting thermodynamic properties of refrigerant mixtures containing R-32, R-125, R-134a, or

R-152a. The Helmholtz energy of the mixture is the sum of the ideal gas contribution, the

compressibility (or real gas) contribution, and the contribution from mixing. The contribution from

mixing is given by a single equation which is applied to all mixtures used in this work. The

independent variables are the reduced density and reduced temperature. The model may be used to

calculate thermodynamic properties of mixtures at various compositions including dew- and

bubble- point properties and critical points. It incorporates the most accurate pubhshed equation of

state for each pure fluid. The model may be used to calculate the properties of mixtures generally

within the experimental accuracies of the available measured properties. The estimated accuracy of

calculated properties is ±0.1% in density, ±0.5% in the speed of sound, and ±1% in heat

capacities. Calculated bubble-point pressures are generally accurate to within 1%.

KEY WORDS: equation of state; mixtures; R-32; R-125; R-134a; R-152a; thermodynamic

properties.

E-3
1. INTRODUCTION

Applications requiring the use of equations of state capable of accurate prediction of

thermodynamic properties of environmentally-safe refrigerant mixtures are increasing. New

refrigerants and mixtures of refrigerants are used as environmentally acceptable replacements for

chlorofluorocarbons and hydrochlorofluorocarbons in refrigeration, heat pumps, foam-blowing,

and other applications. Mixture equations are required to evaluate the performance of possible

alternatives to the fully halogenated chlorofluorocarbons as working fluids. Many of the refrigerant

mixtures form azeotropes. This is often desirable in vapor compression cycles because the mixture

behaves as a pure fluid during phase changes.

The model presented here for mixtures of refrigerants is part of a more generalized model

reported by Lemmon and Jacobsen [1] for mixtures of cryogens, hydrocarbons, and refrigerants.

This model was initially reported by Lemmon [2] in a slightly different format. The model is based

on corresponding states theory. Reducing parameters which are dependent on the mole fraction are

used to modify absolute values of density and temperature. The Helmholtz energy for an ideal

solution is determined at the reduced density and temperature of the mixture using accurate pure

fluid equations of state for the mixture components.

The portion of the model describing the Helmholtz energy contribution from mixing is

nearly the same for all binary mixtures, and relatively simple scaling factors are used to determine

its magnitude for a particular application. Experimental data from many different mixtures were

used to determine the function. The model is capable of predicting mixture properties for fluids

with limited experimental databases. In addition, all vapor and liquid thermodynamic properties,

including density, energy, entropy, heat capacity, sound speed, vapor-liquid equilibrium, and the

mixture critical temperature, pressure, and density, can be calculated accurately using this

approach.

The mixtures studied in this work are:

• R-32/R-125

•R-32/R-134a

E-4
• R-125/R-134a

•R-32/R-125/R-134a

•R-134a/R-152a

Data for these mixtures were used to evaluate the behavior of the model, to calculate the

coefficients of the generalized refrigerant mixture equation and to calculate the individual

parameters for each binary mixture. Table I lists the pure fluid equations of state used in this work.

All equations were converted for convenience from the published format to a fundamental form

widely used in system analysis and design.

2. THE MIXTURE EQUATION FOR ENVIRONMENTALLY-SAFE

REFRIGERANTS

The Helmholtz energy for mixtures of refrigerants can be calculated using

I E a^(5,x,x) = a^ (5,T,x)+ a (5,x,x). (1)

The Helmholtz energy for the ideal solution is

I

*" r r I a^(5,x,x) = Y.^i a.(5,x) + a.(5,x) +lax, , (2)

where r is the number of fluids in the mixture, a? is the ideal gas Helmholtz energy of component

/, and af is the residual Helmholtz energy of component i. Equations for the ideal gas Helmholtz

energy and residual Helmholtz energy for the pure fluids are given in the references shown in

Table L

The contribution to the Helmholtz energy from mixing used in this work is

£ * * ^°

1 (3)

a = Z E ^p^,^„I^^*5*X* + Xr_32Xr_i25(A^1i5X + A^,2S)+^R-32-^R-125^135'

p = \ q=p + \ k=\

where the A^^, ik, andy^ are coefficients and exponents obtained from linear regression of

experimental mixture data. The coefficients, A^^, of the equation are determined using a stepwise

least-squares technique with a search and selection procedure which selects an optimum group of

terms from a large comprehensive function based on statistical evaluation of the significance of

E-5
individual terms. All thermodynamic properties can be calculated from the Helmholtz energy using

differentiation with respect to density or temperature as described by Lemmon et al. [3],

The independent parameters of the equation for the contribution to the Helmholtz energy

from mixing given in this work are reduced density and temperature and mixture composition. In

addition, one other parameter, Fpq, is required which relates the excess-like properties of one

binary mixture to those of another. This value is determined with a nonlinear fitting procedure that

minimizes the sum of squares of the deviations between the equation and the data for a small select

set of measurements.

The reduced density and temperature of the mixture are

5 = p/p^ and (4)

T = r/r, (5)

where Tj and py are the reducing values,

T,= ixj, + i i x^xl (6)

P; =

p=\ ' p=\q=p+\

' X. ' '

P = l Pc /> = ' ?=/' + !

-1

(7)

and where ^^^ and t,pq are used to define the shapes of the reducing parameter lines. The units for

temperature and density are kelvins and moldm-3. These reducing parameters are not the same as

the critical parameters of the mixture and the use of these parameters allows the calculation of VLE

properties above the reducing temperature. These parameters define the relationship between the

critical lines of different binary mixtures and are determined simultaneously in the nonlinear fit with

the generalized factor. Details of the fitting procedures to determine the generalized factors and

mixture parameters are given by Lemmon [2] and are not repeated here. The values of the

coefficients and exponents in Eq. (3) are given in Table n. The generalized factors and mixture

parameters are given in Table HI.

E-6
3. COMPARISONS OF CALCULATED REFRIGERANT MIXTUREPROPERTIES

TO EXPERIMENTAL DATA

Summary comparisons of values calculated using the mixture equation to P -p-T data,

isochoric heat capacity data, sound speed data and VLE data for refrigerant mixtures are given in

Table IV, as well as the temperature and composition range for the first component listed. Details

of the property calculation methods including those for VLE states are given by Lemmon [2].

The accuracies of calculated values of various properties are determined by comparing them

to measured values. Stadstical analyses are used to determine the overall estimated accuracy of the

model in representing the data sets used in determining the coefficients of equations of state, and to

define the ranges of estimated accuracies for various properties calculated with the formulation.

The statistics used to evaluate the equation are based on the percent deviation for

any property, X,

%AX = 100

/ ^Y —Y daU -^ calc ^ daU (8)

Using this definition, the statistics in Table IV are defined as:

1

" I I AAD= -X |%AX,|, (9)

Bias= -^f;(%AX,). (10)

Two of the terms listed in Eq. (3) are specific to the R-32/R-125 binary mixture and

another is specific to the R-32/R-134a binary mixture. The mixture model behavior is sensitive to

the accuracy of the pure fluid equations of state. The equation of Outcalt and McLinden [5] for R32 was used in place of the equation from Tillner-Roth and Yokozeki [4] to determine the

sensitivity of the deviations of calculated properties to the selection of the R-32 equation. Replacing

the equation increased the AAD for the data of Magee [1 1] by 20%. The deviations for other data in

the supercritical region increased as well. The database for mixtures containing R-32 may be more

reliable than that for the pure fluid. Additional details of this analysis are given by Lemmon [2].

This issue may be resolved by additional measurements and a revised correlation for R-32.

E-7
Likewise, there are similar deficiencies in the R-125 database in the critical region.

4. ACCURACY ASSESSMENT

An assessment has been made to determine the accuracy of the mixture model reported

here. Generally, the equation is accurate to ±0.1% in density, ±1% in heat capacity, and ±0.5% for

the speed of sound. For binary mixtures, calculated bubble point pressures are generally within

±1%. The model is valid between from 200 to 450 K up to 60 MPa as verified by experimental

data. In regions where there are no binary mixture data, the accuracy is estimated to be of the same

magnitude. However, this cannot be verified by the authors until experimental data are available to

support these conclusions. Although the equation was developed using mostly binary data, it is

accurate in calculating the properties of mixtures with three or more constituents. This conclusion

is based upon comparisons of calculated values to the limited data available for multicomponent

systems.

Future measurements will confirm whether the equation is valid for other mixtures and in

regions not covered by the experimental data used in the development of this model. These data

will enable continued evaluation and development of the model.

REFERENCES

1. E.W. Lemmon and R.T Jacobsen, submitted to Int. J. Thermophys. (1997).

2. E.W. Lemmon, A Generalized Model for the Prediction of the Thermodynamic Properties of

Mixtures Including Vapor-Liquid Equilibrium . Ph.D. Dissertation, University of Idaho

(1996).

3. E.W. Lemmon, R.T Jacobsen, S.G. Penoncello, and S.W. Beyerlein, Adv. Cryo. Eng.

2£:1891 (1994).

4. R. Tillner-Rotii and A. Yokozeki, to be submitted to /. Phys. Chem. Ref. Data (1997).

5. S.L. Outcalt and M.O. McLinden, Int. J. Thermophys. W\):19 (1995).

6. R. TiUner-Roth and H.D. Baehr, /. Phys. Chem. Ref. Data 21(5):657 (1994).

7. R. Tillner-Roth, Int. J. Thermophys. I6(l):91 (1995).

£-8
8. CD. Holcomb, VLE data reported in: Haynes, W.M, Thermophvsical Properties of HCFC

Alternatives. National Institute of Standards and Technology, Boulder, Colorado, Final Report

for ARTI MCLR Project Number 660-50800 (1996).

9. H. Kiyoura, J. Takebe, H. Uchida, H. Sato, and K. Watanabe, /. Chem. Eng. Data 41:1409

(1996).

10. M. Kleemiss, Thermodynamic Properties of Binary and Ternary Mixtures of R134a, R32,

R125 and R143a - Measurements and Equations of State, Dissertation, University of

Hannover, Germany (1996).

11. J.W. Magee, private communication, NIST, Boulder, Colorado, USA (1995).

12. K. Oguchi, T. Takaishi, N. Yada, T. Namiki, and T. Sato, Kanagawa Institute of

Technology, Japan (1995).

13. C.-C. Piao, I. Iwata, and M. Noguchi, to be submitted to /. Chem. Eng. Data, Mech. Eng.

Lab., Daikin Industries, Ltd., Japan (1996).

14. T. Sato, H. Kiyoura, H. Sato, and K. Watanabe, Int. J. Thermophys. 17:43 (1996).

15. L.A. Weber and D.R. Defibaugh, Int. J. Thermophys. 11(5):863 (1994).

16. H.-L. Zhang, H. Sato, and K. Watanabe, /. Chem. Eng. Data 41:1401 (1996).

17. D.R. Defibaugh and G. Morrison, Int. J. Refrig. 1S(8):518 (1995).

18. K. Fujiwara, H. Momota, M. Noguchi, 13th Japan Symp. Thermophys. Prop. A 11.

6

:61

(1992).

19. Y. Higashi, 19th International Congress ofRefrigeration, Netherlands IVa:297 (1995).

20. M. Nagel and K. Bier, Int. J. Refrig. 1S:534 (1995).

21. J.V. Widiatmo, H. Sato, and K. Watanabe, High Temp.-High Press. 21:677 (1993).

22. T. Sato, H. Kiyoura, H. Sato, and K. Watanabe, /. Chem. Eng. Data :^(4):855 (1994).

23. J.V. Widiatmo, T. Fujimine, H. Sato, and K. Watanabe, to be submitted to /. Chem. Eng.

Data (1997).

24. Y. Higashi, Int. J. Thermophys. 1^(5): 1175 (1995).

25. J.V. Widiatmo, H. Sato, and K. Watanabe, Fluid Phase Equilib. 92:199 (1994).

E-9
26. T. Hozumi, H. Sato, and K. Watanabe, 4th Asian Thermophysical Properties Conference,

Japan, p. 307 (1995).

27. M. Higuchi and Y. Higashi, Proc. 16th Japan Symp. Thermophys, Prop., Hiroshima, p. 5

(1995).

28. Y, Higashi, private communication. Department of Mechanical Engineering, Iwaki Meisei

University, Japan (1996).

29. M. Dressner and K. Bier, Thermodvnamic Measurements in Binary Svstems of Nevy

Refrigerants. Fortschr.-Ber. VDI 1(332) (1993).

30. R. Tillner-Roth, /. Chem. Thermodyn. 2^.U\9 (1993).

31. A.J. Grebenkov, Yu.G. Kotelevsky, V.V. Saplitza, O.V. Beljaeva, T.A. Zajatz, and B.D.

Timofeev, Proc, CFC's: The Day After, IIR Comm. Bl, B2, El, E2; Padova, Italy (1994).

32. M. Kleiber, M., Fluid Phase Equilib., 92: 149 (1994).

Table I. Pure Fluid Equations of State for Refrigerants Used in the Mixture Model.

Fluid Author Temperature Maximum

Range (K) Pressure (MPa)

R-32 Tillner-Roth and Yokozeki [4] 160-500 60

R-125 Outcalt and McLinden [5] 174-500 68

R-134a Tillner-Roth and Baehr [6] 169.85-500 70

R-152a TUhier-Roth [7] 154.65-435 30

E-10
Table n. Coefficients and Exponents of the Mixture Equation

Nk Ik Jk

1 -0.245476271425 x lO-i

2 -0.241206117483

3 -0.513801950309 x 10-2

4 -0.239824834123 x lO-i

5 0.259772344008

6 -0.172014123104

7 0.429490028551 x lO-i

8 -0.202108593862 x 10-3

9 -0.382984234857 x 10-2

10 0.262992331354 x 10-5

11 -0.198606229861 x lO-i

12 0.143226453485

13 0.369107330061 x lO-i

1 2

1 4

1 -2

2 1

3 4

4 4

5 4

6

6 4

8 -2

1 1

1

1

Table DI. Parameters of the Mixture Equation

Binary Mixture p? 5P? spq

R-32/R-125

R-32/R-134a

R-125/R-134a

R-134a/R-152a

-0.789585

-0.314574

0.104729

0.125997

-O.00530177

-0.00392017

0.0

0.00487924

14.663390

5.737916

-2.736948

-2.041245

E-U
Table IV. Comparisons of Mixture Properties Calculated from the Model to Refrigerant Mixture

Data

Author No.

Points

Temp.

Range

Comp.

Range

AAD Bias

R-32/R.125 -P.p -r

Holcomb [8] 45 279-341 0.24-0.96 0.940 0.361

ICiyoura et al. [9] 94 330-440 0.37-0.61 0.487 -0.486

Kleemiss [10] 415 243-413 0.50-0.51 0.187 -0.035

Magee[ll] 228 200-400 0.50 0.145 0.039

Oguchi et al. [12] 6 355-430 0.87 0.275 -0.138

Piaoetal. [13] 533 263-393 0.37-0.90 0.245 -0.098

Satoetal. [14] 156 320-440 0.70-0.90 0.333 -0.328

Weber and Defibaugh [15] 17 338-373 0.55 0.497 -0.497

Zhang et al. [16] 124 300-380 0.50-0.70 0.279 -0.273

OveraU 1618 200-440 0.24-0.96 0.263 -0.112

R-32/R-125 - VLE

Defibaugh and Morrison [17] 10 249-338 0.76 0.151 -0.060

Fujiwara et al. [18] 8 273 0.06-0.90 2.060 2.060

Higashi [19] 45 283-346 0.23-0.90 0.329 0.322

Holcomb [8] 30 280-340 0.34-0.95 0.311 0.172

Kleemiss [10] 23 224-333 0.48-0.52 0.364 -0.227

Nagel and Bier [20] 34 205-345 0.24-0.95 0.504 0.019

Oguchi et al. [12] 11 250-350 0.87 0.432 0.432

Piaoetal. [13] 10 263-283 0.37-0.90* 0.665 -0.648

Widiatmo et al. [21] 24 280-310 0.20-0.90 0.530 0.383

Overall 195 205-350 0.06-0.95 0.476 0.192

E-12
Table IV. (Continued)

Author No.

Points

Temp.

Range

Comp.

Range

AAD Bias

R-32/R-134a -P-p -T

Holcomb [8] 44 279-340 0.13-0.97 1.123 0.535

Kleemiss [10] 390 243-413 0.50-0.56 0.090 0.001

Magee [11] 219 200-400 0.50 0.121 -0.079

Oguchi et al. [12] 19 363-473 0.39 0.426 0.407

Piaoetal. [13] 633 263-393 0.33-0.89 0.405 -0.225

Sato et al. [22] 220 320-440 0.33-0.89 0.147 -0.011

Weber and Defibaugh [15] 17 338-373 0.51 1.048 -1.048

Widiatmo et al. [23] 22 280-330 0.40 0.105 -0.003

Overall 1564 200-473 0.13-0.97 0.274 -0.095

R-32/R-134a - Cy

Magee [11] 131 205-343 0.50 0.373 0.294

R-32/R-134a - co

Hozumi et al. [26] 144 303-343 0.16-0.90 0.030 0.030

R-32/R-134a - VLE

Defibaugh and Morrison [17] 25 253-358 0.50-0.55 0.398 0.260

Fujiwara et al. [18] 6 273-273 0.20-0.92 3.262 -3.262

Higashi [24] 39 283-365 0.12-0.67 1.171 1.104

Holcomb [8] 48 280-340 0.16-0.78 0.400 0.127

Kleemiss [10] 16 223-343 0.42-0.52 0.229 0.118

Nagel and Bier [20] 50 203-369 0.21-0.77 0.480 0.136

Oguchi etal. [12] 34 238-301 0.27-0.71 0.868 0.776

Piaoetal. [13] 10 261-283 0.33-0.89* 0.565 -0.565

Widiatmo et al. [25] 30 280-340 0.33-0.89 1.646 1.646

Overall 258 203-369 0.12-0.92 0.801 0.445

E-13
Table IV. (Continued)

Author No.

Points

Temp,

Range

Comp.

Range

AAD Bias

R 125/R-134a -F=p ^T

Holcomb [8] 17 280-342 0.35-0.72 0.297 -0.166

Kleemiss [10] 407 243-413 0.50-0.51 0.125 -0.103

Magee[ll] 268 200-400 0.50 0.110 -0.076

Weber and Defibaugh [15] 18 303-373 0.50 0.378 -0.353

Widiatmo et al. [23] 110 280-350 0.09-0.92 0.112 -0.093

Overall 820 200-413 0.09-0.92 0.128 -0.100

R-125/R-134a - VLE

Higuchi and Higashi [27] 55 283-365 0.18-0.78 0.373 0.030

Holcomb [8] 40 280-340 0.26-0.65 0.521 -0.032

Kleemiss [10] 24 224-343 0.46-0.51 0.432 -0.326

Nagel and Bier [20] 31 206-365 0.25-0.75 0.541 0.219

Widiatmo et al. [23] 75 280-350 0.09-0.92 0.712 0.711

OveraU 225 206-365 0.09-0.92 0.542 0.235

R-32/R-125/R-134a - P p -7

Holcomb [8] 42 244-346 0.20-0.68 0.984 0.490

Kiyoura et al. [9] 105 315-440 0.38-0.52 0.481 0.254

Kleemiss [10] 369 243-413 0.33-0.35 0.093 -0.019

Oguchietal. [12] 12 365-430 0.38-0.47 0.090 -0.023

Piaoetal. [13] 994 263-393 0.19-0.47 0.307 -0.115

Widiatmo et al. [23] 53 280-340 0.38-0.46 0.185 -0.181

Overall 1575 243-440 0.19-0.68 0.278 -0.056

E-14
Table IV. (Continued)

Author No.

Points

Temp.

Range

Comp.

Range

AAD Bias

R-32/R-125/R-134a - VLE

Higashi [28] 52 273-359 0.17-0.54 0.624 -0.235

Holcomb [8] 58 221-345 0.05-0.60 0.901 0.030

Kleemiss [10] 44 222-353 0.14-0.66 0.372 -0.306

Nagel and Bier [20] 29 205-362 0.19-0.43 0.601 -0.160

Piaoetal. [13] 31 270-326 0.32-0.38 0.664 0.165

Widiatmo et al. [23] 43 280-340 0.35-0.46 0.504 0.500

Overall 257 205-362 0.05-0.66 0.625 -0.007

R-134a/R-152a - P -p -T

Dressner and Bier [29] 139 333-423 0.49-0.54 0.225 -0.076

Tillner-Roth [30] 1679 243-433 0.25-0.75 0.073 0.034

Weber and Defibaugh [15] 11 353-373 0.50 0.042 0.017

OveraU 1829 243-433 0.25-0.75 0.085 0.025

R-134a/R-152a - co

GrebenJcovetal. [31] 120 230-336 0.69 0.466 -0.446

R-134a/R-152a - VLE

Defibaugh and Morrison [17] 13 248-368 0.78 0.284 0.009

Kleiber [32] 25 255-298 0.31-0.98 0.146 -0.134

Tillner-Roth [30] 65 243-378 0.23-0.75 0.071 -0.035

OveraU 103 243-378 0.12-0.98 0.116 -0.054

"Dew point compositions

E-15
Proceedings, 6th International Refrigeration Conference at Purdue West Lafayette, IN, July 23-26, 1996, pp 409^14

A NEXT GENERATION REFRIGERANT PROPERTIES DATABASE'

Mark O. McLinden and Sanford A. Klein^

Physical and Chemical Properties Division

National Institute of Standards and Technology

Boulder, Colorado 80303 USA

ABSTRACT

A completely revised version of the REFPROP refrigerant properties database is described. This program is based on the most accurate pure fluid and mixture models currently available. It implements three models for the thermodynamic properties of pure fluids: the modified Benedict- Webb-Rubin (MBWR) and Helmholtz equations of state, and an extended corresponding

states (ECS) model. Mixture calculations employ a new model which applies mixing rules to the pure-fluid Helmholtz energies.

Viscosity and thermal conductivity are modeled with either fluid-specific correlations or a new variation on the ECS method. These models are implemented in a suite of subroutines written in standard FORTRAN. A separate graphical user interface provides a convenient means of accessing the models. It will generate tables and/or plots for any user-specified mixture. Numerous options to customize the output are available as well as copy and paste capabilities to and from other applications.

INTRODUCTION

Knowledge of the thermophysical properties is essential for the evaluation of alternative refrigerants and the design of equipment using them. The REFPROP computer database from the National Institute of Standards and Technology (NIST)

(Huber et al. 1995) has been one of the more widely used tools designed to provide these data. In the initial versions of REFPROP (Gallagher et al. 1993), the intent was to provide data on a wide variety of fluids to allow screening studies of

possible replacements for the CFC or HCFC refrigerants. For many of these fluids, only sparse data were available, and,

consequently, the database relied primarily on a simple model with few adjustable parameters—the Camahan-Starling-DeSantis

(CSD) equation of state. As the alternative refrigerants move from the laboratory to use in commercial equipment, highly accurate properties are required for a more limited set of fluids. We describe a completely revised REFPROP program

(designated as Version 6) based on the most accurate pure fluid and mixture models currently available.

THERMODYNAMIC MODELS

REFPROP 6 calculates the thermodynamic properties using comprehensive equations of state. This approach ensures thermodynamic consistency and allows calculations at all conditions. Other approaches, such as the combination of a vapor- phase model with vapor pressure and liquid density equations may not be applicable in the compressed liquid and supercritical

regions and do not always give reliable results for derived properties such as heat capacity and speed of sound.

Pure Fluid Models

Three models are used for the thermodynamic properties of pure components, depending on the availability of data. The

first is the modified Benedict-Webb-Rubin (MBWR) equation of state. This model was first proposed by Jacobsen and Stewart

(1973) and has been applied to a wide variety of fluids, including hydrocarbons, cryogenic fluids, and refrigerants. It is capable of accurately representing the properties of a fluid over wide ranges of temperature, pressure, and density. The MBWR

equation is the basis for the current international standard for the properties of R 123 (Younglove and McLinden, 1994).

The MBWR equation expresses pressure as an explicit function of temperature and molar density and is of the form,

9 15 P = ^a,p''+exp[(p/p-"f]^ay"-'^ , (1)

n=\ n=10 where the a, are simple functions of temperature resulting in a total of 32 adjustable parameters. For a complete description of

the energy quantities (enthalpy, entropy, etc.), the MBWR equation is combined with an expression for the molar heat capacity of

the ideal-gas state, that is, vapor in the limit of zero pressure. A form combining polynomial and theoretical terms is used:

exp(»J-l]

All of the thermodynamic properties can be computed from Eqs. (1-3), as detailed by Younglove and McLinden (1994)

^P ^Zf^^Lr^ , , ,i2

. with ^k^-z: (2,3)

k [exp(«J-l ^

'Contribution of the National Institute of Standards and Technology, not subject to copyright in the United States. ^Pernnaneni address: Solar Energy Laboratory, University of Wisconsin, Madison, Wisconsin 53706 USA

E-16
The second high-accuracy pure-fluid equation of state is in terms of reduced nnolar Heimholtz free energy: A

= A= a"^ + a^ = In 5 + y a,T'' + Y a^t''5'^' expf-yS'' ) , (4)

where the first two terms on the right side of Eq. (4) constitute the ideal-gas contribution a'^\ they are analogous to the combination of the ai term in the MBWR equation of state (ai = RT) and the ideal-gas heat capacity (Eqs. 2 and 3). The second summation is the residual, or real-fluid, contribution a". The temperature and density are expressed in the dimensionless

variables x = T*/T and 5 = p/p*, where T* and p* are reducing parameters which are often, but not always, equal to the critical parameters. The a, and a^ are numerical coefficients fitted to experimental data and the exponents /,, t^, d,, and d^ are typically

determined by a selection algorithm starting with a large bank of terms. The parameter y is equal to for terms with /;; = 0; it is equal to 1 for terms with li^*- 0. This "Heimholtz model" is the basis for the international standard formulation for RI34a

(Tillner-Roth and Baehr, 1994).

This model is sometimes termed the "fundamental equation" because it gives a complete description of the thermodynamic

properties, as discussed by Tillner-Roth and Baehr (1994), but the MBWR equation of state combined with a Clf equation is entirely equivalent. The sources for the high-accuracy formulations implemented in the database are summarized in Table I. The accuracy of these formulations vary, but those for R123 are typical: experimental data are reproduced with average absolute deviations of 0.04 % for densities, 0.05 % for vapor pressures, and 0.75 % for heat capacities.

The third pure-fluid model is the extended corresponding states (ECS) model of Huber and Ely (1994). It is used for fluids with limited experimental data. Simple corresponding states is based on the assumption that different fluids obey, in reduced coordinates, the same intermolecular force law. This assumption leads to the conclusion that, with the appropriate

scaling of temperature and density, the reduced residual Heimholtz energies and compressibilities (Z = p/RTp } of the unknown

fluid "/' and a reference fluid "0" (for which an accurate, wide-ranging equation of state is available) are equal:

a'j(Tj,Pj) = a'o{To,Po) , and Zj[Tj,Pj) = Zo{Tq,Po] (5,6)

When combined with an expression for the ideal gas heat capacity (such as Eqs. 2 and 3), all other thermodynamic properties can be calculated. The reference fluid is evaluated at a "conformal" temperature and density:

fj To = -^ = Tj

T-tTIf 'J 0(7)

and (7.8)

where the multipliers 1// and hj are termed reducing ratios; they are composed of ratios of the critical parameters. Simple

corresponding states was developed for spherically symmetric molecules. The ECS model extends the method to other types of molecules by the introduction of the "shape factors" 9 and (]). These shape factors are taken here to be functions only of

temperature. (If data sufficient to fit the density dependence were available, it would be preferable to develop a full equation of

state.) The shape factors are fitted to experimental data, typically vapor pressures and saturated liquid densities. The reference

fluid is chosen to be chemically similar to the fluid of interest. The fluids represented with the ECS model are listed in Table 2.

Table 1 —High-accuracy pure-fluid equations of state currently implemented in refprop 6

Fluid Model

Rll Heimholtz

R12 Heimholtz

R22 Heimholtz

R32 MBWR

RI13 Heimholtz

R123 MBWR

R124 Heimholtz

R125 MBWR

R134a Heimholtz

R143a MBWR

R152a MBWR

R 170 (ethane) MBWR

R290 (propane) MBWR

R600 (butane) MBWR

R600a(isobutane) MBWR

R717 (ammonia) Heimholtz

R744 (CO2) MBWR

Source Jacobsen ef a/. (1992)

Marx era/. (1992)

Kame'i etal. (1995)

Outcalt and McLinden (1995)

Marx era/. (1992)

Younglove and McLinden (1994)

deVnes etal. (1995)

Outcalt and McLinden (1995)

Tillner-Roth and Baehr ( 1 994)

Outcalt and McLinden ( 1994)

Outcalt and McLinden (1996)

Younglove and Ely (1987)

Younglove and Ely (1987)

Younglove and Ely (1987)

Younelove and Ely (1987)

Tillne'r-Roth era/. (1993)

Ely etal. (1987)

Limits of Application

Temperature (K) Pressure (MPa) Density (kg/m^)

162.68 (T,p)- 625

1 16.10 (T,p)- 525

115.73 (T,p)-550

136.34 (T,p)- 500

236.93 (Ttp) - 525 166(Ttp)-500

100-400

172.52 (T,p)-500

169.85 (T,p)-460

161.34 (T,p)- 500

154.56 (T,p)- 500

90.35 (T,p) - 600

85.87 (T,p) - 600

134.86 (T,p)- 500

113.55 (T,p)-600

195.49 (T,p)- 700

216.58 (T,p)- 500

0-30

0-200

0-60

0-60

0-200

0-40

0-40

0-60

0-70

0-40

0-60

0-70

0- 100 0-70

0-35

0- 1000 0-40

0- 1768 0- 1830 0- 1721 0- 1429 0- 1705 0- 1774 0- 1903 0- 1691 0- 1591 0- 1329 0- 1193 0-664

- 752 0-754

- 745

0-900

0-1178

E-17
Table 2—Fluids currently represented with the extended corresponding states (ECS) model in REFPROP 6

Fluid R13

R14

R23

R114

Reference Fluid R12

R12

R134a

R12

Fluid R115

R116

R134

R141b

Reference Fluid R12

R12

R134a

R22

Fluid Reference Fluid R142b R22

R227ea R134a

R236fa R134a

R245ca R134a

Fluid Reference Fluid RC270 propane RC318 propane R1270 propane RE 170 propane

Mixture \ [ode!

The thermodynamic properties of mixtures are calculated with a new model which was developed, in slightly different forms, independently by Tillner-Roth (1993) and Lemmon (1996). It applies mixing rules to the Helmholtz energy of the mixture components:

n n-\ n

7 = 1 /)=! (/=/)+!

This mixing formula may be applied directly to the Helmholtz equation of state. Application to the MBWR equation of state and

the ideal gas heat capacity expression used with the MBWR and ECS models requires transformations:

a' ^—nP-RTp)dV , and (10)

--\ C;!dT--\ -!^dT , (11) RTh.. '^ RJt., T

^,<y^V_V_i + in RT R

rp ^

^ TrefPref J

where href and Sref are an arbitrary reference enthalpy and entropy at the reference state specified by Tref and Pref- The first summation in Eq. (9) represents the ideal solution; it consists of ideal gas (superscript id) and residual or real fluid (superscript r) terms for each of the pure fiuids in the n component mixture. The Xj In jc terms arise from the entropy of mixing of ideal gases where Xj is the mole fraction of component). The double summation accounts for the "excess" free energy or "departure" from ideal solution. The Fp^ are generalizing parameters which relate the behavior of one binary pair with

another; Fp^ multiplies the a"^"^ term(s), which are empirical functions fitted to experimental binary mixture data. The a'^ and

"^no"^ functions in Eqs (9 and 10) are not evaluated at the temperature and density of the mixture Tm,x and p„,> but, rather, at a reduced temperature and density x and 6. These T and 5 are very much in the spirit of the conformal temperature and density of

the ECS method and are a key innovation in this model. Several mixing rules for the reducing parameters are used, including

T = -—, With T' = Y,^''T.pr''p^,('^rVi'' ' ^"d 02)

p

with P P

p = \ </=!

,1/3-^ (13)

If only limited vapor-liquid equilibrium (VLE) data are available the a""" term is taken to be zero, and only the ^^^^ and/or

kv,pq parameters are fitted. The kT,pa parameter is most closely associated with bubble point pressures, and it is necessary to reproduce azeotropic behavior. The kv,pq parameter is associated with volume changes on mixing. (Ternary and higher order

mixtures are modeled in terms of their constituent binary pairs: kjp^ = 1 and kyp^ = 1 for p = a.) If extensive data, including

single-phase pressure-volume-temperature (PVT) and heat capacity data, are available, the a^^y " function can be determined. The Fnii parameter is used (either alone or in combination with kT,pq and kv,p(,) to generalize the detailed mixture behavior descnoed by the a^""' function to other, similar, binary pairs. Lemmon (1996) has determined an o!'^^"'* function based on

data for 28 binary pairs of hydrocarbons, inorganics, and HFC's (but using different mixing rules than bqs. 12 and 13).

This "mixture Helmholtz model" provides a number of advantages. By applying mixing rules to the Helmholtz energy of the mixture components, it allows the use of high-accuracy equations of state for the components, and the properties of the mixture

will reduce exactly to the pure components as the composition approaches a mole fraction of 1. Different components in a mixture may be modeled with different forms; for example, a MBWR equation may be mixed with a Helmholtz equation of state.

If the components are modeled with the ECS method, this mixture model allows the use of a different reference fluid for each component. The mixture is modeled in a fundamental way, and thus the departure function is a relatively small contribution to the total Helmholtz energy for most refrigerant mixtures. The great flexibility of the adjustable parameters in this model allows

an accurate representation of a wide variety of mixtures, provided sufficient experimental data are available.

E-18
TRANSPORT PROPERTY MODELS

Pure Fluid Models

The transport properties of viscosity and thermal conductivity are modeled with the residua! concept. In this representation, the property % (representing either viscosity rj or thermal conductivity X) is composed of three contributions:

X = x''(7') + Ax'"(r,p) + Ax'(r,p). (14)

X"' is a dilute gas term which is a function only of temperature, and Ax'^ is a residual term accounting for the behavior at higher

densities, including Hquid densities. Ax'^ is primarily a function of density, but it may also be a function of temperature and

density. The thermal conductivity approaches infinity at the critical point, and this critical enhancement is expressed by Ax*^- This term is significant for thermal conductivity even quite far from the critical point. For viscosity the enhancement is small except extremely close to the critical point and may be safely ignored in all practical applications. A variety of fluid-specific correlations for viscosity and thermal conductivity, based on the residual concept, have been implemented in the database.

Where fluid-specific correlations are not available, the transport properties are modeled with the extended corresponding

states method of Klein et al. (1996). This method is a modification of the ECS model of Huber et al. (1992a, 1992bj and

shares many of the concepts of the ECS model for the thermodynamic properties described above. The dilute gas term is modeled using kinetic theory (Hirschfelder et al. 1954). The residual term is scaled to the properties of a reference fluid:

1/2^-2/3 Ax;(7',p) = AxS(ro,Po)y}''%-

1/2 (15)

where M is molar mass. The determination of the reducing ratios involves an iterative procedure. Eqs. (5) and (6) are iterated

to find the/y and hj satisfying those two conformal relations. The conformal temperature Tq is defined in Eq. (7). The

conformal density po in the residual term (but not the critical enhancement) is further modified from that in the thermodynamic

ECS method (Eq. 8) by the introduction of a third shape factor 4*:

crit Po = P,%f(7'mp). (16) Py

The shape factor T is a linear function of density and is determined by a least-squares fit to experimental data. Separate T

functions are determined for viscosity and thermal conductivity.

Mixture Model

The viscosity of a mixture is calculated with the extended corresponding states method of Klein et al. (1996):

ll(r,p,x) = Tl'^(r,;c) + ATlo(r//„p/2jF^+A7l*(p,;c) , with F^= fT^h^^'^gf . (17,18)

The low-density contribution Ti"' is calculated with kinetic theory (Hirschfelder et al. 1954) and Ari* is an Enskog-theory hard- sphere correction for size and mass differences given by Ely (1981). The reducing ratios /^ and hx are determined by applying

Eqs. (5) and (6) to the mixture. The "mass shape factor" for the mixture, gx, is determined by mixing rules applied to the component mass shape factors, defined as

ATi,(r,,p,)

'^ A,o(r//,.p/./^ '^ • ^''^

The residual viscosity of component; in Eq. (19) is evaluated, using either a fluid-specific correlation or the pure-fluid ECS

model described above, at the conformal conditions ^_[nuJj_ ^^^

p PmiA (20,21)

This formulation yields the pure-fluid values exactly as the composition approaches a mole fraction of 1. A analogous method

is used for thermal conductivity. For the transport properties, R134a is used as the reference fluid for all mixture components.

PROPERTY SUBROUTINES

The property models described above are implemented as a suite of FORTRAN subroutines. These routines have been

completely rewritten from earlier versions of REFPROP. Source code is provided with the database so that users may link the property routines with their own application. ("Translation" routines will allow applications written using the subroutine calls

E-19
of earlier versions of REFPROP to use the new subroutines.) The routines are written in ANSI-standard FORTRAN 77 and are compatible with FORTRAN 90. They are written in a structured format, are internally documented with extensive comments, and

have been tested on a variety of compilers.

The fluid or mixture of mterest is specified with a (required) call to the subroutine "SETUP." This routine reads the coefficients to the NlST-recommended models for that fluid. Alternative property models and/or nonstandard reference states may be specified by calls to additional (optional) setup routines. Routines are provided to calculate thermodynamic and

transport properties and surface tension at a given (7, p, x) state. Iterative routines provide saturation properties at a specified

(7, a) or (P, x) state. Flash calculations calculate single- or two-phase states at specified {P, h, x), (P, T, x), etc.

The routines mentioned above are independent of the model. Underlying these routines are sets of "core" routines for each of the models implemented in the database. Each such set is highly modular and is contained in a separate file. Coefficients needed for a particular model are stored in common blocks, but these commons are referenced only by routines in the same file. These sets of subroutines, thus, resemble "units" in the Pascal language with clearly demarcated "interface" and

"local" declarations. This structure is intended to simplify the addition of future models to the database and will make such

additions almost totally transparent to the user.

Numerical coefficients to the property models are stored in text files. There is one file per fluid and one file containing

coefficients for the mixture departure functions. These files are read (once) upon the call to SETUP. Fluids can be added to or

deleted from the database without recompilation.

USER INTERFACE

A graphical user interface has been developed for REFPROP 6 independent of the code implementing the algorithms. The

interface provides a convenient means to calculate and display thermodynamic and transport properties. It is written for the Windows^'^ operating system. (Mention of commercial products is to fully specify the database and does not constitute endorsement by NIST or imply that they are necessarily the best suited for the purpose.) The interface is written in Pascal; it accesses the FORTRAN property subroutines via a dynamic link library. Screen shots of the interface program are shown in Figures 1 and 2. The program is controlled through the use of the following pull-down menus (visible at the top of Figure 1):

File provides commands to save and print generated tables and plots. Individual items or entire sessions with multiple

windows may be saved or recalled. The standard "print setup" and "quit" commands are also present. The Edit menu provides copy and paste commands which allow selected data to be exchanged with other applications.

The Options menu provides commands for selecting the unit system, properties of interest, and the reference state. These

options may be stored for recall at a later time. A user-defined set of preferences is loaded upon program startup. The pure fluid or mixture of interest is specified with commands in the Substance menu. Most of the refrigerant mixtures

of current commercial interest (those having an ASHRAE R400 or R500-series designation) are predefined. In addition, new mixtures can be specified and saved by combining up to five of the pure components listed in Tables 1 and 2.

REFPROP - NIST Refrlgeraiil Properties Database

£ilc £dil Options Substance Calculate £lot N^indow tJclp R152B/123 |30/70|: saL T=-40.0 to 140.0 |C| Temperature

|C1

Pre*»ure(l> IMPilDen*My<l) |ka(ml|Enthalpy^l) IKJ/Wgl -10 00

-30 00

-20 -10

0.02561 04207

1384 9 1362 4

1564

167 1 T-xplot:R152a/I23

00 20 40 60 80 1 00 R123 Composition (mass fraction] R1S2a

Name fl152vH123 Flu<l fll52a / H123 (30 / 701 Helo

R123-Z.Z-diclilafa-I.I.I-tiilluotoethane CASK 30G-83-2I

Holai matt

152.93

Triple pt. lemp. IGEOOIKJ

Critical pretttae Critical dentity 3.BE2 (MPal 3.59E [kmol/m^] Range of applicability Hin. temperature Man. temperature 1EG0|K| E0O.O|K|

Boiling pt. temp. 300.97 |K|

Acentric factor 0.2B

MaK. pretture 40.0 (lilPal

Critical temp. 45E.83 (K|

Dipole at NBP

1.3G(debye|

MaM. dentity

11 GOIkmol/m^l

NIST Rec aWR MBWH equation ol slaloiYounglovo 1 McLinden |1394|. LITERATURE REFERENCE

B A Youriglova and M Q McLnden An Nefnational ilandaid equalion-ol-itale lormulation o( the thetmodynamc piopeiliei ol leltigefant 123 (2.2dichlo(o-1 .1 ,1 trilluoioethanel. J Phys. Chem. Rcf, Data 23 731779(19341

\Equalion ot ttate/VircontyXThetmal conductivilyXSuilace leniion/

Figure 1. Screen shot of interface showing a typical data table and temperature-composition diagram.

Figure 2. Pure-component information screen available from the interface.

E-20
The Calculate menu initiates the calculations that generate a property table. Each property selected for display is shown in a separate column of the table. Two types of tables are provided. The first type, such as that shown in Figure 1 , provides

properties at saturation or with a property (such as temperature or pressure) held constant with another selected property

varying over a specified range. The second type allows the user to .select the independent variables. Values of the independent variables may then be entered with the keyboard, read from a file, or pasted from another application.

The Plot menu provides publication-quality x-y plots of any variables appearing in a table. In addition, temperatureentropy, pressure-enthalpy, temperature-composition and pressure-composition diagrams may be generated automatically. Controls are provided to modify the plot size, axis scaling, plot symbols, line type, legend, and other plot features. Each table or plot appears in a separate window and can be accessed, resized, or retitled with commands in the Window

menu. The number of windows is limited only by available memory.

A complete online-help system can be accessed through the Help menu. A status line at the bottom of the screen displays the currently specified mixture, composition, and reference state. Clicking on the status line will call up a screen for each of the components providing documentation for fluid constants, the source of the models, and their range of applicability, as shown in Figure 2.

CONCLUSIONS

The new refprop 6 database implements a variety of high-accuracy models for the thermodynamic and transport

properties of refrigerants and their mixtures. These models are implemented as a suite of FORTRAN 77 subroutines written in a modular fashion; this structure will facilitate the incorporation of additional fiuids and future models. A graphical user interface provides a convenient means of accessing the models and producing tables and plots of any specified mixture.

ACKNOWLEDGEMENTS

We gratefully acknowledge many helpful discussions with D. Friend, M. Huber, A. Laesecke, and R. Perkins. The stay of S.A. Klein at NIST was funded, in part, by the University of Wisconsin and by a grant from the National Science Foundation

under Agreement No. 9527385. Any opinions, findings, and conclusions or recommendations expressed in this publicadon are those of the authors and do not necessarily reflect the views of the National Science Foundation.

REFERENCES

de Vries, B., Tillner-Roth, R. and Baehr, H.D. (1995). 19th International Congress of Refrigeration, IVa: 582-589.

Ely, J.F. (1981). 7. Res. NBS 86: 597-604.

Ely, J.F., Magee, J.W. and Haynes, W.M. (1987). Research Report RR-1 10, Gas Processors Association, Tulsa, OK.

Gallagher, J., Huber, M., Morrison, G. and McLinden, M. (1993). NIST Standard Reference Database 23, version 4.0. Standard Reference Data Program, National Institute of Standards and Technology, Gaithersburg, MD.

Hirschfelder, J.O., Curtiss, C.F. and Bird, R.B. (1954). Molecular Theory of Gases and Liquids, John Wiley and Sons. Inc. Huber, M.L. and Ely, J.F. (1992a). Fluid Phase Equilibria 80: 239-248.

Huber, M.L., Friend, D.G. and Ely, J.F. (1992b). Fluid Phase Equilibria 80: 249-261.

Huber, M.L. and Ely, J.F. (1994). Int. J. Refrigeration 17: 18-31. Huber, M., Gallagher, J., McLinden, M. and Morrison, G. (1995). NIST Standard Reference Database 23, version 5.0. Standard Reference Data Program, National Institute of Standards and Technology, Gaithersburg, MD.

Jacobsen, R.T. and Stewart, R.B. (1973). J. Phvs. Chem. Ref Data 2: 757-922.

Jacobsen, R.T., Penoncello, S.G. and Lemmon; E.W. (1992). Fluid Phase Equilibria 80: 45-56. Kamei, A., Beyerlein, S.W. and Jacobsen, R.T. (1995). Int. J. Thermophysics 16: 1 155-1 164. Klein, S.A., McLinden, M.O. and Laesecke, A. (1996). An improved extended corresponding states method for estimation of

viscosity of pure refrigerants and mixtures. Int. J. Refrigeration (submitted).

Lemmon, E.W. (1996). PhD Thesis, Mechanical Engineering, University of Idaho, Moscow, ID. Marx, v., PruB, A. and Wagner, W. (1992). Neue Zustandsgleichungen fur R 12, R 22, R 1 1 und R 1 13. Beschreibung des thermodynamishchen Zustandsverhaltens bei Temperaturen bis 525 K und Driicken bis 200 MPa, VDI Verlag.

Outcalt, S.L. and McLinden, M.O. (1995). Int. J. Thermophysics 16: 79-89.

Outcalt, S.L. and McLinden, M.O. (1996). J. Phys. Chem. Ref Data 25: 605-636.

Outcalt, S.L. and McLinden, M.O. (1994). In: Thermophysical Properties of HFC- 143a and HFC- 152a, final report to the Air- Conditioning and Refrigeration Technology Institute, Arlington, VA, report no. DOE/CE/238 10-39. Tillner-Roth, R. (1993). Dr.-Ing. Thesis, Institut fijr Thermodynamik, Universitat Hannover, Germany.

Tillner-Roth, R. and Baehr, H.D. (1994). J. Phys. Chem. Ref Data 23: 657-729.

Tillner-Roth, R., Harms-Watzenberg, F. and Baehr, H.D. (1993). DKV-Tagungsbericht 20, II: 167-181. Younglove, B.A. and Ely, J.F. (1987). J. Phys. Chem. Ref Data 16: 577-798.

Younglove, B.A. and McLinden, M.O. (1994). J. Phys. Chem. Ref Data 23: l'i\-119.

E-21

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