Documentation and Assessment of the 

Transport Property Model for Mixtures 

Implemented in NIST REFPROP

(Version 8.0)

Justin C. Chichester

Marcia L. Huber

NISTIR 6650
 
 

NISTIR 6650

Documentation and Assessment of the 

Transport Property Model for Mixtures 

Implemented in NIST REFPROP 

(Version 8.0)

Justin C. Chichester

Marcia L. Huber

Physical and Chemical Properties Division 

Chemical Science and Technology Laboratory 

National Institute of Standards and Technology 

Boulder, CO 80305-3328 

June 2008

U.S. DEPARTMENT OF COMMERCE 

Carlos M. Gutierrez, Secretary 

NATIONAL INSTITUTE OF STANDARDS AND TECHNOLOGY 

James M. Turner, Deputy Director 
 
iii 

Contents 

1. Introduction......................................................................................................................2 

2. Viscosity Model............................ ...................................................................................3 

 2.1 Pure Fluids..................................................................................................................3 

 2.2 Mixtures......................................................................................................................6 

3. Thermal Conductivity Model...........................................................................................8 

 3.1 Pure Fluids..................................................................................................................8 

 3.2 Mixtures.....................................................................................................................11 

4. Results.............................................................................................................................14 

5. Conclusions.....................................................................................................................29 

6. References........................................................................................................................44 

1 

Documentation and Assessment of the Transport Property Model for Mixtures 

Implemented in NIST REFPROP (Version 8.0) 

Justin Chichester and Marcia L. Huber 

National Institute of Standards and Technology∗

Boulder CO USA 80305-3328 

In this report, we describe an extended corresponding states model for 

viscosity and thermal conductivity of mixtures implemented in Version 8 of the 

NIST computer program REFPROP. The model is a modification of a one-fluid, 

extended corresponding states (ECS) model for thermal conductivity and viscosity 

originally developed by Ely and Hanley (Ind. Eng. Chem. Fundam., 1981, 20:323-

332). We apply the model to selected mixtures representative of the fluids available 

in the NIST 23 database REFPROP (v8), such as refrigerants and natural gas 

constituent fluids, and present comparisons with experimental data. Comparisons are 

given for gas, liquid, and supercritical conditions. 

 

Keywords: extended corresponding states; natural gas; refrigerants; thermal conductivity; 

viscosity. 

 

∗ Physical and Chemical Properties Division, Chemical Science and Technology Laboratory. 
2 

1. Introduction 

The NIST Reference Fluid Thermodynamic and Transport Properties Database, 

NIST23 (REFPROP) [1], provides thermophysical properties for 84 pure fluids and for 

mixtures of up to 20 components. For pure fluids, it contains equation of state formulations 

for the thermodynamic properties, and models or correlations for the viscosity and thermal 

conductivity. The references for the individual pure fluid formulations are provided in the 

program itself through the “fluid information” screen. For mixtures, there is presently only 

one model available for the thermal conductivity and viscosity. It is based upon the extended 

corresponding states concept. In 1981, Ely and Hanley published a one-fluid, extended 

corresponding states model for viscosity [2], and in 1983 applied the model to thermal 

conductivity [3]. Since that time, there have been several modifications to the Ely-Hanley 

method, each improving upon the original model [4-10]. Variations of this model have been 

successfully implemented in two older NIST databases, NIST4 (Supertrapp) [11] and 

NIST14 (DDMIX) [12], and most recently in NIST23 (REFPROP) [13]. In this report we 

will describe the formulation used in the most recent version of the REFPROP computer 

program [1] and give comparisons with experimental data for the viscosity and thermal 

conductivity of selected mixtures of interest, including refrigerant mixtures and natural gas 

mixtures. 
3 

2. Viscosity Model 

2.1 Pure Fluids 

In the extended corresponding states model used here, we represent the viscosity of a 

pure fluid as a sum of a dilute gas and a residual contribution, and apply a corresponding 

states principle to the residual contribution only [2], 

( , ) ( ) ( , ) ( ) ( , ) ( , ), 0 0 0

* * η T ρ =η T + Δη T ρ = η T + Δη T ρ Fη T ρ (1) 

where the superscript * denotes a dilute gas value, and the subscript 0 denotes a reference 

fluid value. The viscosity of the reference fluid is evaluated at a conformal temperature and 

density T0 and ρ0 given by 

T T / f 0 = (2) 

and 

. ρ 0 = ρh (3) 

The quantities f and h are called equivalent substance reducing ratios, and relate the 

reference fluid to the fluid of interest using a ratio of critical parameters (denoted by the 

subscript c) and functions of temperature and density known as shape functions θ and Φ, 

θ

c0

c

T

T f = (4) 

and 

φ

ρ

ρ

c

c0 h = . (5) 
4 

The shape factors can be considered functions of both temperature and density. For 

small, nonpolar, almost-spherical molecules they are nearly unity and can be thought of as 

functions to compensate for deviations from a spherical shape. Different approaches have 

been taken to determine shape factors [4, 5, 8] usually depending on the amount and quality 

of p,V,T information available for the fluid; in this work, we generally have available 

accurate formulations for the thermodynamic properties of the pure fluids, either in terms of 

a Helmholtz free energy equation or a pVT equation of state (EOS), and we use a form of the 

“exact” shape factor method [4]. It is a requirement in this method to first determine the 

thermodynamic shape factors. 

The dilute-gas viscosity in Eq. (1) is given by the Chapman-Enskog theory [14] 

, 16

5

( ) 2 (2,2)

* B

Ω = πσ

η

mk T

T (6) 

where the dilute-gas viscosity is η*, m is the molecular mass, kB is the Boltzmann constant, 

and T is the absolute temperature. We will further assume that a Lennard-Jones 12-6 

potential applies, and use the Lennard-Jones collision diameter for σ. Neufeld et al. [15] 

gave the following empirical correlation for the calculation of the collision integral Ω(2,2)

* * (2,2) 0.14874 0.77320 2.43787 1.16145( *) 0.52487 2.16178 T T T e e − − − Ω = + + , (7) 

with the dimensionless temperature T* = kBT/ε, and ε the minimum of the Lennard-Jones 

pair-potential energy. The range of validity of this empirical correlation is 0.3 < T* < 100. 

The factor Fη in Eq. (1) is found by using the expression 
5 

1/ 2

0

1/ 2 2 / 3

⎥

⎦

⎤ ⎢

⎣

⎡ = −

M

M Fη f h

, (8) 

where M is the molar mass of the fluid and M0 is the molar mass of the reference fluid. The 

model as developed to this point is predictive, and does not use any information on the 

viscosity of the fluid (except for the dilute-gas piece that requires Lennard-Jones ε and σ). 

The functions f and h are found from thermodynamic data and are described in Klein et al. 

[9]. In order to improve the representation of the viscosity, an empirical correction factor 

may be used if there are experimental viscosity data available. We then evaluate Eq. (1) at 

ρ0,v instead of ρ0, where [9] 

( , ) ( , ) ( ), ρ 0,v T ρ = ρ 0 T ρ ψ ρ r (9) 

and ψ is a polynomial in reduced density ρr = ρ/ρc of the form 

( ) , r

0

r

k n

k

k ψ ρ ∑c ρ =

= 

(10) 

where the coefficients ck are constants found from fitting the experimental viscosity data. As 

indicated in Eq. (1), in order to evaluate the viscosity of a particular fluid, the value of the 

residual viscosity of a reference fluid is required. It is not necessary to use the same 

reference fluid for all fluids. However, when the model is used in a predictive mode, it is 

best to select the reference fluid that is most similar in chemical nature to the fluid of 

interest. The reference fluid should also have a very accurate equation of state and viscosity 

surface. When using pure fluid experimental viscosity to essentially “correct” the viscosity, 
6 

the choice of reference fluid is not as important, since an empirical correction factor 

determined from data is applied, as in Eqs. (9) through (10). 

2.2 Mixtures 

The extension of the model to mixtures is similar to that presented for pure fluids, but 

involves an extra step. First, one represents the properties of the mixture in terms of a 

hypothetical pure fluid (denoted here by the subscript x) that is obtained through the use of 

mixing rules, 

j ij ij

n

i

i

n

j

f x hx ∑ ∑x x f h = =

− = 1 1

1

 (11)

, 1 1

j ij

n

i

i

n

j

x h ∑ ∑x x h = =

= (12)

 

and combining rules, 

(1 ) ij i j ij, fη f = f f − k (13) 

 ( ) (1 ) 8 ,

3 1/ 3 1/ 3

ij hη

i j

ij k

h h

h − + = , (14) 

where we introduce binary interaction parameters kij,fη and kij,hη that can be determined by 

fitting mixture experimental data when available. Otherwise, they are set to zero. 
7 

The iterative method used to determine fi and hi by use of equations of state is 

described in detail in the manuscript of Klein et al. [9]. In that work, the authors note that 

although the procedure for finding the shape factors is straightforward, it is complicated in 

practice and the iterative procedure can fail at both high densities and very low densities. In 

addition, we noted some failures with fuel mixtures in the supercritical region that occurred 

with the use of the Helmholtz equations of state, but did not occur with a much simpler 

equation of state such as the Peng-Robinson. We suspect that this might be due to in part to 

the iterations entering the two-phase region where the Helmholtz equations can have very 

complicated behavior. To partially alleviate this problem, we assume that at some high 

temperature, here selected as 1000 K, the individual pure-fluid shape factors will have a 

limiting value of 

x

c j

j x T

T

f f , = (15) 

 

c j

x hj hx ρ ,

ρ = , (16) 

where the subscript x denotes the mixture and j denotes the pure fluid. In the fluid region 

where the compressibility factor (z=p/ρRT) of the fluid mixture is above 0.3, and the 

temperature is above the critical temperature of the mixture, we linearly interpolate the 

value of the fj and hj between its limiting value in Eqs. (15) through (16) and the value at the 

critical temperature. 

We assume that the viscosity of the mixture obeys a corresponding states principle, 

( , , ) ( , ) ( , , ) ( , ) ( , ) ( , ), 0 0 0

* * η T ρ x =η T x + Δη T ρ x =η T x + Δη T ρ Fη T ρ (17) 
8 

where the factor Fη for mixtures is found with 

1/ 2 2 / 3 1/ 2

x x x F f h g −η = , (18) 

1/ 2 4 / 3

1

1/ 2 4 / 3 1/ 2

1/ 2 1

x x

n

j

i j ij ij ij

n

i

x f h

x x f h M

g

∑ ∑ = = =

 (19) 

. 1/ 1/

2

i j

ij M M

M

+ = (20) 

For the dilute gas viscosity, binary interaction parameters kij,σ and kij,ε may be used, where 

ij ij i j σ = (1− k ,σ ) σ σ (21) 

/ (1 ) ( / )( / ) B , B B k k k k ij ij i j ε ε ε = − ε . (22) 

When there are sufficient dilute-gas viscosity data available, these parameters are 

obtained by fitting experimental data. Otherwise, they are set to zero. 

3. Thermal Conductivity Model 

3.1 Pure Fluids 

We start with the procedure of Ely and Hanley [3] and represent the thermal 

conductivity of a fluid as the sum of translational (from collisions between molecules) and 

internal (due to internal motions of the molecule) modes of energy transfer, 

 ). ( , ) ( ) ( , int trans λ T ρ = λ T + λ T ρ (23) 
9 

The translational contribution may be further divided into a translational dilute-gas 

contribution (denoted here by a superscript *) that is a function only of temperature, a 

residual contribution, and a critical enhancement, 

 ), ( , ) ( ) ( , ) ( , trans * r crit λ T ρ = λ T + λ T ρ + λ T ρ (24) 

leading to the following expression for the thermal conductivity 

 ). ( , ) ( ) ( ) ( , ) ( , int * r crit λ T ρ = λ T + λ T + λ T ρ + λ T ρ (25) 

We use an Eucken correlation for the internal contribution 

 , 2

5 ( ) *

*

int int ⎥

⎦

⎤ ⎢

⎣

⎡ = C − R

M

f T p

η λ (26) 

where Cp* is the ideal-gas heat capacity in J/(mol·K), R is the molar gas constant [16] (8.314 

472 J/(mol·K)), η* is the dilute-gas viscosity (μPa·s) as given in Eq. (6), fint is set to 1.32·10-

3

, and λ is in W/(m·K). If sufficient dilute-gas thermal conductivity data are available, fint is 

fit to a polynomial in temperature, 

 . fint = a0 + a1T (27) 

For the dilute-gas translational contribution (in W/(m·K)) we use 

 , 4

15 10 ( )

3 *

*

M

R T η λ

− × = (28) 

where the dilute gas viscosity η* is from Eq. (6). The residual contribution is found using 

extended corresponding states: 

 ( , ) ( , ) , 0 0

r

0

r λ T ρ = λ T ρ Fλ (29) 

with 
10 

 

1/ 2

1/ 2 2 / 3 0

⎥

⎥

⎦

⎤

⎢

⎢

⎣

⎡ = −

M

M

F f h λ . (30) 

In order to improve the representation of the thermal conductivity, an empirical 

correction factor may be used if there are experimental thermal conductivity data available. 

We then evaluate Eq. (29) at ρ0,k instead of ρ0 , where [8] 

 ), ( , ) ( , ) ( ρ 0,k T ρ = ρ 0 T ρ χ ρ r (31) 

and χ is a polynomial in reduced density ρr = ρ/ρc of the form 

 , ( ) r

0

r

k n

k

χ ρ ∑bk ρ == (32) 

where the coefficients bk are found from fitting the experimental thermal conductivity data. 

 The critical contribution is computed using a simplified crossover model developed 

by Olchowy and Sengers [17], 

( ) 0

crit 0 B

6 ( , ) = Ω − Ωπηξ

ρ λ ρC R k T

T p , (33) 

where the heat capacity at constant pressure, Cp (T,ρ), is obtained from the equation of state, 

R0 = 1.03 is a universal constant [18], and the viscosity, η(T,ρ), is obtained from the method 

described earlier. The crossover functions Ω and Ω0 are determined by 

arctan( ) ( ) , 2

d d

⎥

⎥

⎦

⎤

⎢

⎢

⎣

⎡

+ ⎟

⎟

⎠

⎞

⎜

⎜

⎝

⎛ − Ω = ξ ξ

π

q C

C

q C

C C

p

v

p

p v (34) 
11 

.

( )

3

1 ( )

1 1 exp2

2

1 d

d

0

⎥

⎥

⎥

⎥

⎥

⎦

⎤

⎢

⎢

⎢

⎢

⎢

⎣

⎡

⎟

⎟

⎟

⎟

⎟

⎠

⎞

⎜

⎜

⎜

⎜

⎜

⎝

⎛

⎟

⎟

⎠

⎞ ⎜

⎜

⎝

⎛

+

− Ω = −

−

ρ

ξ ρ ξ

π c q

q

 (35) 

The heat capacity at constant volume, Cv(T,ρ), is obtained from the equation of state, and the 

correlation length ξ is given by 

γ γ

ρ ρ ρ ρ

ρ

ρ ξ ξ

/

R R

/

2

c

c

0

( , ) ( , ) v

T T

v

p

T

T

T

p

p T

⎥

⎥

⎦

⎤

⎢

⎢

⎣

⎡

∂

∂ − ∂

∂ ⎥

⎦

⎤ ⎢

⎣

⎡

Γ = . (36) 

The partial derivative of density with respect to pressure ∂ρ/∂p|T, is evaluated with the 

equation of state at the system temperature T and a reference temperature, TR. The reference 

temperature is a value where the critical enhancement is assumed to be negligible: TR = 

1.5Tc. The exponents γ = 1.239 and ν = 0.63 are universal constants [18]. The critical 

amplitudes Γ and ξ0 are system-dependent and can be determined by the asymptotic behavior 

of the equation of state in the critical region. For most fluids, we set these to the fixed values 

Γ = 0.0496 and ξ0 = 1.94 × 10-10 m. The thermal conductivity at the critical point itself is 

infinite. The cutoff wavenumber qd (or alternatively, its inverse qd

-1) is system-dependent 

and determined from regression of experimental data when available or estimated [19]. 

3.2 Mixtures 

The extension of the model to mixtures is similar to that presented for viscosity, 

 ) ( , , ) ( , ) ( , ) ( , ) ( , , * int r crit λ T ρ x = λ T x + λ T x + Δλ T ρ + λ T ρ x (37) 

and only the residual contribution is treated with corresponding states. 

 ) ( , , ) ( , ) ( , ) ( , ) ( , ) ( , , crit

0 0 0

* int λ T ρ x λ T x λ T x λ T ρ F T ρ λ T ρ x = + + Δ λ + . (38) 
12 

The internal and translational dilute-gas contributions for the mixture are found with the 

empirical mixing rule, 

∑

∑ =

=

+

+ =

n

j

ji

n

i

i

j j j

x

x T T

T x T x

1

1

int *

*

mix

int

mix

( ( ) ( )) ( , ) ( , )

φ

λ λ

λ λ , (39) 

with 

( )

( )1/ 2

2 * * 1/ 4

,

8(1 / )

(1 ) 1 / ( / )

j i

ij j i j i

ji M M

k M M

+

− + = η η φ λ

. (40) 

All quantities are evaluated at the mixture temperature T, and the dilute-gas viscosity is 

found with Eq. (6). The parameter kij,λ is an empirical binary interaction parameter for the 

dilute gas region determined by fitting experimental data. 

 The residual contribution for the mixture requires calculation of a mixture Fλ with 

 , 1/ 2 2 / 3 1/ 2

x hx gx F f −λ = (41) 

( )4 / 3

1/ 2

1

1/ 2

1

1/ 2 4 / 3

1/ 2

1/ 2 0

1/ 1/

2 ( ) ij

i j

n

j

i j ij

n

x x i

x h

g g

x x f f h

M

g

−

= = ⎟

⎟

⎠

⎞

⎜

⎜

⎝

⎛

+ = ∑ ∑ , (42) 

with gi found with 

4 / 3

2

r

0 0

r

0

0 ( , )

( , ) −

⎟

⎟

⎠

⎞

⎜

⎜

⎝

⎛ = j j

j j j

i f h

T

T

g M

λ ρ

λ ρ . (43) 

Similar to the procedure for the calculation of the viscosity, we use the combining rules 
13 

(1 ) ij i j ij, fλ f = f f − k (44) 

( ) (1 ) 8 ,

3 1/ 3 1/ 3

ij hλ

i j

ij k

h h

h − + =

, (45) 

where we introduce binary interaction parameters kij,fλ and kij,hλ that can be determined by 

fitting experimental data when available. Otherwise, they are set to zero. 

 As observed for pure fluids, there also is an enhancement of the thermal conductivity 

observed near the critical point; however, theory for mixtures is not as well developed as for 

pure fluids. In REFPROP 8 we make a simple approximation where we obtain the 

parameters necessary for the simplified Olchowy-Sengers model (Eqs. 33 through 36, above) 

by taking a mole fraction average, 

i c

1

TR 1.5 x T

n

i

∑ i

=

= (46) 

∑=

=

n

i

i i q x q

1

d d (47) 

∑=

Γ = Γ

n

i

i i x

1 (48) 

 . 1

0 ∑=

=

n

i

i i ξ x ξ (49) 

The other parameters γ , ν are universal and are unchanged from their pure-fluid values. The 

thermodynamic properties are evaluated at the mixture T and ρ, not at scaled values, and the 

mixture Tc, pc and ρc are used. In addition, at the critical point the thermal conductivity of a 

pure fluid is infinite [20]; for a mixture, theory and experiment suggest that a finite value is 
14 

achieved [21]. Therefore we have imposed a somewhat arbitrary upper limit (100 % of the 

background value) on the magnitude of the enhancement possible in a mixture. 

4. Results 

 We tested the model on a limited set of mixture data containing both refrigerant 

mixtures and hydrocarbon mixtures. When possible, we obtained binary interaction 

parameters to improve the representation of the data. The parameters were obtained by least 

squares regression of experimental data. The resulting values for the binary interaction 

parameters for both the dilute and residual viscosity are given in Table 1 below. These 

binary interaction parameters are contained in the file HMX.BNC that is distributed with the 

REFPROP computer program.* In addition, we provide a list of the Lennard-Jones 12-6 

parameters for pure fluids in REFPROP8 in Table 2. References for these values are 

provided, however, often it was necessary to estimate the values. The estimation methods of 

Chung et al. [22] and a corresponding-states estimation method [7] were used. 

Table 3 gives the deviations of the model compared with experimental data for 

selected binary systems. We use the following definitions for average absolute deviation 

(AAD), bias, and root-mean-square (RMS) deviation: 

∑=

= −

n

i i

i

n

AAD

1

exp

cal

1 100

η

η , (50) 

 

∑= ⎟

⎟

⎠

⎞

⎜

⎜

⎝

⎛ = −

n

i i

i

n

BIAS

1

exp

cal

1 100

η

η , (51) 

and 
15 

1 . 100 1 100

2

1

exp

cal 2

1

exp

cal

2

⎟

⎟

⎠

⎞

⎜

⎜

⎝

⎛

⎟

⎟

⎠

⎞

⎜

⎜

⎝

⎛ − − ⎟

⎟

⎠

⎞

⎜

⎜

⎝

⎛

⎟

⎟

⎠

⎞

⎜

⎜

⎝

⎛ = ∑ − ∑ = =

n

i i

i

n

i i

i

n n

RMS

η

η

η

η (52) 

In general, the model represents the viscosity of the mixtures to within 5 % to 10 %. One 

notable exception is the data for the CO2/decane system measured by Cullick and Mathis 

[23]. The model does not represent this system well even with the use of interaction 

parameters. This is probably due to the large size differences between the molecules. It has 

been noted previously [24] that one-fluid corresponding states types of models have 

difficulties representing systems where the molecules differ greatly in size. The trend of 

increasing deviations as the size difference between the molecules increases is also 

demonstrated by the data of Aucejo et al. [25] in the systems of pentane with increasingly 

larger alkanes, with the largest deviations observed for the pentane/dodecane system. Further 

work in this area is needed. 

 There also are large discrepancies between the model predictions and the viscosity 

data of Heide [26] for the binary mixture of R143a and R125. This system has also been 

studied by Laesecke [27] and is represented well by the model, so we suspect that there may 

have been an experimental problem with the Heide [26] data for this system. 
16 

Table 1a. Values of interaction parameters for viscosity for selected refrigerant, 

cryogen and hydrogen mixtures. 

System kij,σ kij,ε kij,fη kij,hη Source of 

data 

R134a/R32 0.0437 -0.4678 -0.0645 -0.0305 [28] [29] [30] 

R125/propane -0.21 0.3776 0.0 0.0 [30] 

R22/propane -0.1284 0.3331 0.0 0.0 [31] 

R115/propane -0.2744 0.414 0.0 0.0 [31] 

R32/propane 0.2857 -0.4346 0.0 0.0 [28] 

R12/R22 -0.0213 0.0698 0.0 0.0 [31] 

R13/R22 -0.0332 0.1389 0.0 0.0 [31] 

R13B1/R22 -0.2505 0.7028 0.0 0.0 [31] 

R14/R22 0.0523 -0.1194 0.0 0.0 [31] 

R14/CO -0.0421 -0.0638 0.0 0.0 [32] 

R22/R152a -0.1382 0.4355 0.0 0.0 [31] 

R32/R124 0.0 0.0 -0.0813 -0.1158 [29] 

R125/R134a -0.0045 0.0038 -0.0829 -0.0106 [28] [29] [33] 

hydrogen/methane -0.2744 -0.1738 0.0 0.0 [22, 34] 

hydrogen/nitrogen -0.6225 0.1388 0.0 0.0 [34] 

hydrogen/CO2 -0.9218 0.8531 0.0 0.0 [34] 

hydrogen/oxygen -0.7661 0.3683 0.0 0.0 [34] 

hydrogen/CO -0.7599 0.8536 0.0 0.0 [34] 

hydrogen/helium -0.0480 -1.9733 0.0 0.0 [34] 

hydrogen/argon -1.0538 0.8177 0.0 0.0 [34] 

hydrogen/neon -0.6105 -0.4449 0.0 0.0 [34] 

hydrogen/ammonia -0.5275 0.8419 0.0 0.0 [34] 

hydrogen/ethylene -0.7131 0.8778 0.0 0.0 [34] 

hydrogen/propylene -0.5744 0.7662 0.0 0.0 [34] 

hydrogen/cis-butene -0.7844 0.8518 0.0 0.0 [34] 

helium/argon -0.187 -2 0.0 0.0 [35] 

krypton/CO -0.08889 -0.05653 0.0 0.0 [32] 
17 

Table 1b. Values of interaction parameters for viscosity for selected natural gas 

component and hydrocarbon mixtures. 

System kij,σ kij,ε kij,fη kij,hη Source of 

data 

methane/CO -0.0165 -0.2493 0.0 0.0 [32] 

methane/CO2 -0.00677 -0.3342 0.0 0.0 [36] 

methane/ethane 0.0331 -0.2549 0.0 0.0 [37] 

methane/propane -0.0182 -0.0263 0.0 0.0 [37] 

methane/butane 0.0030 -0.1292 0.0 0.0 [36] 

methane/nitrogen -0.3594 0.9076 -0.0612 0.0398 [38] 

ethane/propane 0.0557 -0.2792 0.0 0.0 [37] 

ethane/butane 0.0350 -0.1858 0.0 0.0 [37] 

ethane/hydrogen -0.4544 0.3877 0.0 0.0 [34] 

propane/butane 0.0825 -0.3471 0.0 0.0 [37] 

propane/hydrogen -0.5926 0.7304 0.0 0.0 [34] 

pentane/heptane 0.0 0.0 -0.0892 -0.1474 [25] 

pentane/dodecane 0.0 0.0 -0.1188 -0.4911 [25] 

pentane/decane 0.0 0.0 -0.0965 -0.2775 [25] 

pentane/nonane 0.0 0.0 -0.1195 -0.1610 [25] 

hexane/octane 0.0 0.0 -0.0519 -0.0010 [25] 

hexane/dodecane 0.0 0.0 -0.0168 -0.2999 [25] 

hexane/decane 0.0 0.0 -0.0550 -0.1310 [25] 

hexane/nonane 0.0 0.0 -0.0263 -0.0738 [25] 

hexane/cyclohexane 0.0 0.0 -0.12 0.2243 [39] 

heptane/dodecane 0.0 0.0 -0.0716 -0.1600 [25] 

heptane/decane 0.0 0.0 -0.0374 -0.0677 [25] 

heptane/nonane 0.0 0.0 -0.0704 -0.0117 [25] 

octane/dodecane 0.0 0.0 -0.0802 -0.1160 [25] 

octane/decane 0.0 0.0 -0.0299 -0.0488 [25] 

nonane/dodecane 0.0 0.0 -0.0634 -0.0803 [25] 

nonane/decane 0.0 0.0 -0.0671 -0.0241 [25] 

decane/dodecane 0.0 0.0 -0.0499 -0.0614 [25] 
18 

Table 2. Lennard-Jones 12-6 parameters for selected pure fluids in REFPROP v8. 

Fluid ε/kB(K) σ(nm) Ref. Fluid ε/kB(K) σ(nm) Ref. 

R11 363.6 0.5447 Estimated ammonia 386 0.2957 [40] 

R12 297.24 0.5186 Estimated argon 143.2 0.335 [41] 

R13 204.00 0.4971 [42] butane 280.51 0.573 [43] 

R14 164.44 0.4543 [44] CO 91.7 0.369 [12] 

R32 289.65 0.4098 Fit to [45] CO2 251.196 0.3751 [46] 

R115 201.90 0.5876 [47] cyclohexane 297.1 0.6182 [43] 

R116 226.16 0.5249 Estimated H2O,D2O 809.1 0.2641 [43] 

R124 275.80 0.5501 [48] decane 490.51 0.686 [49] 

R125 249.00 0.5190 [50] dimethyl ether 395 0.4307 [43] 

R141b 370.44 0.5493 Estimated ethane 245 0.4362 [51] 

R142b 278.20 0.5362 [52] ethanol 362.6 0.453 [43] 

R218 266.35 0.5800 Estimated ethylene 224.7 0.4163 [43] 

R227ea 289.34 0.5746 Estimated H2S 301.1 0.36237 [43] 

R236ea 318.33 0.5604 Estimated helium 10.22 0.2551 [43] 

R236fa 307.24 0.5644 Estimated heptane 400 0.64947 [12] 

R245ca 345.44 0.5505 Estimated hexane 399.3 0.5949 [43] 

R245fa 329.72 0.5529 Estimated hydrogen 59.7 0.2827 [43] 

RC318 299.76 0.5947 Estimated isobutene 332 0.5026 Estimated 

R113 376.035 0.6019 Estimated isohexane 395.2 0.5799 Estimated 

Krypton 178.9 0.3655 [43] isopentane 341.06 0.56232 [12] 

Methane 174 0.36652 [53] isobutane 307.55 0.46445 [54] 

Methanol 481.8 0.3626 [43] nitrogen 98.94 0.3656 [41] 

N2O 232.4 0.3828 [43] nonane 472.127 0.66383 [49] 

Neon 32.8 0.282 [43] oxygen 118.5 0.3428 [41] 

neopentane 191 0.644 [55] propane 263.88 0.49748 [56] 

Octane 452.09 0.63617 [49] R23 243.91 0.4278 [57] 

Pentane 341.1 0.5784 [12] R41 244.88 0.4123 Estimated 

Propylene 298.9 0.4678 [43] R114 323.26 0.5770 Estimated 

R22 284.724 0.4666 fit to [58] 

 
19 
20 
21 
22 
23 

 We also made comparisons of the viscosity prediction of the REFPROP model for 

multicomponent mixtures. Comparisons with a ternary refrigerant mixture R32/R125/R134A 

measured by Laesecke et al. [28] are presented in Table 4 and Figures 1a-c. In the Figures, 

the percent deviation is defined as 100*(calculated value-experimental)/experimental. Two 

different compositions 0.33/0.33/0.34 and 0.3/0.1/0.6 mole fraction of R32/R125/R134a 

were measured from 246 to 340 K at pressures up to 2853 kPa with average absolute 

deviations of less than two percent. The estimated uncertainty of this data is given by the 

author as 2.4 %. We also made comparisons with several natural gas mixtures. Figures 2a-c 

show the deviations between the REFPROP model and experimental data for several 

different natural gas mixtures as a function of density, pressure and temperature. Schley et 

al. [65] measured two different natural gases at pressures to 20 MPa. The "L" gas is a low 

calorific gas with a relatively high content (9.75 mole %) of nitrogen; the high calorific "H" 

natural gas contains only 1.5 % nitrogen. Both samples are represented well by the 

REFPROP model. We note that a binary interaction parameter for methane/nitrogen 

viscosity was added to the REFPROP program after the initial release of Version 8 in April 

2007; this parameter significantly improved the results for the high-nitrogen natural gas. 

Carr [66] measured three natural gas samples. Sample 1 is a high-ethane natural gas (73.5 

mole % methane, 25.7 % ethane, 0.6 % N2 and 0.2 % propane), Sample 2 is a high-nitrogen 

natural gas (73.1 % methane, 15.8 % N2, 6.1 % ethane, 3.4% propane, 0.8 % helium, 0.6 % 

n-butane and 0.2 % isobutane) obtained from a transmission line, and Sample 3 is a 

simulated low-ethane natural gas (95.6 % methane, 3.6 % ethane, 0.5 % propane and 0.3 % 

N2). With the exception of the Sample 1 gas at pressures above 20 MPa, the model 

represents the data well. The Assael et al. [67] natural gas is a 5 component mixture of 
24 

methane, ethane, propane, CO2 and nitrogen ( 84.8, 8.4, 0.5, 5.6, and 0.66 mol % 

respectively) at pressures up to 14 MPa and is represented with an average absolute 

deviation of less than two percent, but shows systematic deviations with the largest 

deviations at the lowest temperatures. Systematic deviations are even more pronounced in 

the data from Langelandsvik et al. [68] who measured the viscosity of natural gas samples 

from three different locations in the North Sea. The primary constituents in gas 1 are: 

methane 90.2 %, ethane 6.3 %, propane 0.8 % and CO2 1.8 %; in gas 2: methane 80.0 %, 

ethane 9.3 %, propane, 5.0 %, CO2 2.2 %; and in gas 3: methane 92.2 %, ethane 4.4 %, 

propane 0.5 %, CO2 1 %. REFPROP shows systematic negative deviations for all the 

Langelandsvik et al. [68] samples with the deviations increasing as the pressure increases at 

the lowest temperatures. It was noted in the Langelandsvik et al. [68] manuscript that there 

are unexplained differences between the measurements of Schley et al. [65] and the work of 

Langelandsvik et al. [68]. The REFPROP program is in better agreement with the 

measurements of Schley et al. [65]. 

 The thermal conductivity model was tested on a limited set of thermal conductivity 

mixture data containing both refrigerant mixtures and hydrocarbon mixtures. When possible, 

we obtained binary interaction parameters to improve the representation of the data. The 

resulting values for the binary interaction parameters for both the dilute and residual thermal 

conductivity are given in Table 5 below. 
25 

Table 5. Values of interaction parameters for thermal conductivity for selected 

systems. 

System kij,λ kij,fλ kij,hλ Data source 

methane/CO2 -0.15 0.0 0.0 [69] [70] 

methane/CO -0.07 0.0 0.0 [71] 

methane/propane -0.2 0.0 0.0 [72] 

methane/ethane -0.12 0.0 0.0 [73] 

R22/R142b -0.76 0.0 0.0 [74] [75] 

R32/propane 0.18 -0.08 0.09 [76] 

R32/R134a 0.08 -0.09 0.1 [76] 

R32/R125 0.0 -0.15 0.16 [77, 78] [79] 

R134a/propane 0.07 -0.04 0.20 [76] 

 

 Comparisons for both binary and multicomponent mixtures were made. The results 

are summarized in Tables 6-8 below. Most systems that do not contain data near the critical 

region have deviations of less than 5 % with the notable exception of the binary system 

R32/R125 measured by Kim et al. [80], which shows deviations over 10 %. Since this same 

binary system was measured by other researchers [77-79] and the model agrees well with the 

experimental data, there must be an unexplained problem with the data of Kim et al. [80] 

for this binary system. Perkins et al. [76] made extensive measurements on binary mixtures 

containing R32, R125, R134a, and propane. The measurements covered the gas, liquid and 

some near-critical points. Figures 3a-d show the deviations for Perkins et al. [76] 

measurements as a function of density. The REFPROP model represents the data well except 

for points where the critical enhancement is significant; in these cases the model 

significantly underestimates the thermal conductivity. 

 Roder and Friend [73] made extensive measurements on the thermal conductivity of a 

binary mixture of methane and ethane for three different concentrations (0.7, 0.5, and 0.35 
26 

mol fraction methane) over a range of conditions that included regions where the 

enhancement is significant. Sakonidou et al. [21] also measured the methane/ethane system 

in the critical region. Figures 4a-d show the percent deviations of the data from the model 

with and without the critical enhancement. The figure demonstrates the improvement seen 

when the enhancement term is included; however, it also demonstrates that the present 

model is not adequate, and further work in this area is needed. 

5. Conclusions 

We describe a modified corresponding states model based on an extended 

corresponding states model originally developed by Ely and Hanley [2, 3] for the viscosity 

and thermal conductivity of mixtures. The model has been implemented in the computer 

program REFPROP [1], version 8, available from NIST. Comparisons with selected mixture 

data for viscosity and thermal conductivity are presented. The model represents the viscosity 

and thermal conductivity of mixtures to within 5 to 10 %, except for mixtures containing 

large size differences, where the deviations are larger. A model for the thermal conductivity 

of mixtures in the critical region is also implemented, and although it provides some 

enhancement to the thermal conductivity, comparisons with experimental data indicate that 

the model significantly underpredicts the thermal conductivity in the critical region, and 

further research in this area is recommended. 

We thank our NIST colleagues Dr. Richard Perkins and Dr. Daniel Friend for 

their helpful comments. JC acknowledges a Professional Research Experiences Program 

(PREP) undergraduate fellowship at NIST.
27 
28 
29 

Refrigerant Mixture R32/R125/R134A

-2

-1

0

1

2

3

4

5

200 220 240 260 280 300 320 340 360

Temperature, K

% Deviation

Mix1 0.33/0.33/0.33

Mix2 0.3/0.1/0.6

Figure 1a. Percent deviation of viscosity as a function of temperature for a mixture of 

refrigerants R32, R125 and R134a. 
30 

Refrigerant Mixture R32/R125/R134A

-2

-1

0

1

2

3

4

5

10 11 12 13 14 15 16

Density, mol/L

% Deviations

Mix1 0.33/0.33/0.33

Mix2 0.3/0.1/0.6

Figure 1b. Percent deviation of viscosity as a function of density for a mixture of 

refrigerants R32, R125 and R134a. 
31 

Refrigerant Mixture R32/R125/R134A

-2

-1

0

1

2

3

4

5

10 510 1010 1510 2010 2510 3010

Pressure, kPa

% Deviations

Mix1 0.33/0.33/0.33

Mix2 0.3/0.1/0.6

Figure 1c. Percent deviation of viscosity as a function of pressure for a mixture of 

refrigerants R32, R125 and R134a. 
32 

 

-10

-5

0

5

10

15

0 5 10 15 20

Density, mol/L

% Deviation

Assael et al. 2001

Carr 1953 Natural Gas 1

Carr 1953 Natural Gas 2

Carr 1953 Natural Gas 3

Schley et al. 2004 Natural Gas H

Schley et al. 2004 Natural Gas L

Langelandsvik et al. 2007 Natural Gas 1

Langelandsvik et al. 2007 Natural Gas 2

Langelandsvik et al. 2007 Natural Gas 3

Figure 2a. Percent deviation of viscosity as a function of density for selected natural gas 

mixtures. 
33 

-10

-5

0

5

10

15

0 10000 20000 30000 40000 50000 60000 70000 80000

Pressure, kPa

% Deviation

Assael et al. 2001

Carr 1953 Natural Gas 1

Carr 1953 Natural Gas 2

Carr 1953 Natural Gas 3

Schley et al. 2004 Natural Gas H

Schley et al. 2004 Natural Gas L

Langelandsvik et al. 2007 Natural Gas 1

Langelandsvik et al. 2007 Natural Gas 2

Langelandsvik et al. 2007 Natural Gas 3

Figure 2b. Percent deviation of viscosity as a function of pressure for selected natural gas 

mixtures. 
34 

-10

-5

0

5

10

15

200 250 300 350 400 450 500

Temperature, K

% Deviation

Assael et al. 2001

Carr 1953 Natural Gas 1

Carr 1953 Natural Gas 2

Carr 1953 Natural Gas 3

Schley et al. 2004 Natural Gas H

Schley et al. 2004 Natural Gas L

Langelandsvik et al. 2007 Natural Gas 1

Langelandsvik et al. 2007 Natural Gas 2

Langelandsvik et al. 2007 Natural Gas 3

Figure 2c. Percent deviation of viscosity as a function of temperature for selected natural 

gas mixtures. 
35 

Propane/R32

-60

-50

-40

-30

-20

-10

0

10

0 2 4 6 8 10 12 14 16 18

Density, mol/L

% Deviation

propane 0.3/R32 0.7

propane 0.7/R32 0.3

Figure 3a. Percent deviation of thermal conductivity as a function of density for binary 

mixtures of propane and R32. 
36 

R32/R134a

-8

-6

-4

-2

0

2

4

0 5 10 15 20

Density, mol/L

% Deviation

0.3 R32/0.7 R134a

0.7 R43/ 0.3 R134a

Figure 3b. Percent deviation of thermal conductivity as a function of density for binary 

mixtures of R134a and R32.
37 

Propane/R134a

-40

-35

-30

-25

-20

-15

-10

-5

0

5

10

0 2 4 6 8 10 12 14

Density, mol/L

% Deviation

0.3 propane/0.7 R134a

0.7 R134a/0.3 propane

Figure 3c. Percent deviation of thermal conductivity as a function of density for binary 

mixtures of propane and R134a.
38 

R125/R134a

-20

-15

-10

-5

0

5

10

0 2 4 6 8 10 12 14

Density, mol/L

% Deviation

0.3 R125/ 0.7 R134a

0.7 R125/ 0.3 R134a

Figure 3d. Percent deviation of thermal conductivity as a function of density for binary 

mixtures of R125 and R134a.
39 

-60

-50

-40

-30

-20

-10

0

10

20

0 5 10 15 20 25 30

Density, mol/L

% Deviation

Roder and Friend, 1985

Sakonidou et al. , 1998

Figure 4a. Percent deviation of thermal conductivity as a function of density for binary 

mixtures of methane and ethane near the critical region with a critical enhancement term 

in the model. 
40 

-60

-50

-40

-30

-20

-10

0

10

20

0 5 10 15 20 25 30

Density, mol/L

% Deviation

Roder and Friend, 1985

Sakonidou et al., 1998

Figure 4b. Percent deviation of thermal conductivity as a function of density for binary 

mixtures of methane and ethane near the critical region without a critical enhancement 

term in the model. 
41 

-60

-50

-40

-30

-20

-10

0

10

20

100 150 200 250 300 350

Temperature, K

% Deviation

Roder and Friend, 1985

Sakonidou et al., 1998

Figure 4c. Percent deviation of thermal conductivity as a function of temperature for 

binary mixtures of methane and ethane near the critical region with a critical 

enhancement term in the model. 
42 

-60

-50

-40

-30

-20

-10

0

10

20

100 150 200 250 300 350

Temperature, K

% Deviation

Roder and Friend, 1985

Sakonidou et al., 1998

Figure 4d. Percent deviation of thermal conductivity as a function of temperature for 

binary mixtures of methane and ethane near the critical region without a critical 

enhancement term in the model. 
43 

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