Accepted Manuscript

Densities and derived thermophysical properties of the 0.9505 CO2 + 0.0495 H2Smixture from 273 K to 353 K and pressures up to 41 MPa

Mahmoud Nazeri, Antonin Chapoy, Alain Valtz, Christophe Coquelet, Bahman Tohidi

PII: S0378-3812(16)30217-5

DOI: 10.1016/j.fluid.2016.04.024

Reference: FLUID 11083

To appear in: Fluid Phase Equilibria

Received Date: 2 March 2016

Revised Date: 8 April 2016

Accepted Date: 24 April 2016

Please cite this article as: M. Nazeri, A. Chapoy, A. Valtz, C. Coquelet, B. Tohidi, Densities and derivedthermophysical properties of the 0.9505 CO2 + 0.0495 H2S mixture from 273 K to 353 K and pressuresup to 41 MPa, Fluid Phase Equilibria (2016), doi: 10.1016/j.fluid.2016.04.024.

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Densities and derived thermophysical properties of the 0.9505 CO2 + 0.0495

H2S mixture from 273 K to 353 K and pressures up to 41 MPa

Mahmoud Nazeria*, Antonin Chapoya,b*, Alain Valtzb, Christophe Coqueletb, Bahman Tohidia

a Hydrates, Flow Assurance & Phase Equilibria Research Group, Institute of Petroleum

Engineering, Heriot-Watt University, Edinburgh, EH14 4AS, UK b MINES ParisTech PSL Research University, CTP-Centre of Thermodynamics of Processes,

35, Rue Saint Honoré, 77305 Fontainebleau, France

* Corresponding Author: Antonin Chapoy (Antonin.chapoy@pet.hw.ac.uk)

ABSTRACT

From the development of sour reserves and acid gas injection into oil reservoirs to the impact

of hydrogen sulfide in CCS scheme, knowledge of the phase behaviour and thermophysical

properties of the hydrogen sulfide + carbon dioxide is essential. Due to the toxicity of CO2 -

H2S system, there are few experimental data available in the literature. The aim of this work

is to investigate the densities and thermodynamic properties of acid gases and liquids.

Densities of the 95.05 mol% CO2 + 4.95 mol% H2S binary system were measured

continuously using a high temperature and pressure Vibrating Tube Densitometer (VTD),

Anton Paar DMA 512 at pressures up to 41 MPa at five different temperatures, 273, 283, 298,

323 and 353 K in gas, liquid and supercritical regions. The specific heat capacity,

compressibility factor, dew point and bubble point of the system have also been derived from

the measured density data. The experimental data then were employed to evaluate the GERG-

2008 equations of state and classical cubic equations of state (PR, SRK and VPT) with a CO2

volume correction model as well as volume translation.

Keywords: Density measurement at high pressure, Carbon capture transport and storage,

CO2 and H2S, Thermophysical properties, Acid gas injection

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1. Introduction

One of the most common impurities in natural gas is carbon dioxide. Most reservoirs around

the world contain carbon dioxide. For instance, one of the biggest sour gas resources in

Prudhoe Bay Alaska with more than 20 trillion cubic feet associated gas (with the oil)

contains roughly 12 mol% carbon dioxide and over 10 ppmv hydrogen sulfide. Similar acid

gas content can be found in the Tangguh field in the South-East Asia, Indonesia. The In

Salah gas field in Algeria also contains around10 mol% CO2. In the Middle East, there are

some sour gas fields typically containing 3 mol% H2S and 5 mol% CO2. A comprehensive

survey by Weeks et al. [1] shows that acid gas content composition can span from a few ppm

to 90 mol% for carbon dioxide and from 0 to 60 mol% for hydrogen sulfide.

CO2 content in natural gas pipelines should not be more than a specific amount which is

typically 2 - 2.5 mol%. Any extra carbon dioxide from natural gas reservoirs, therefore,

should be first removed and then transported to either re-injection to oil fields for enhanced

oil recovery or should be stored in a depleted formations (e.g., Sleipner field [2] in the

Norwegian North Sea or aquifer in In Salah gas field [3] in Algeria).

Sour gas and oil fields contain CO2 and H2S which should be removed from the produced oil

or gas before sending to markets [4]. The desulphurisation process is costly and acid gas

injection is seen as an alternative environmentally friendly method to deal with the toxic acid

gases [5]. In the sweetening process of natural gas or oil produced, H2S and CO2 can be

removed by different processes, e.g., using amine. The gas off the amine regenerator can be

transported by pipelines to an injection well to be stored in geological basins[6]. The

knowledge of phase equilibria [7][8] and thermophysical properties [9] of acid gases and

liquids, particularly density [10], are essential to the design considerations of acid gas

injection schemes [11][12].

CCS technology can provide a potential to cut/reduce the large scale release of CO2

emissions. The process comprises three main steps: capture, transport and storage. Three

capturing technologies are under development: pre-combustion, post-combustion and oxy-

fuel combustion. The aim of each capturing technology is to capture CO2, preventing it from

release to the atmosphere, following this to transport it to a suitable place of storage.

However, the CO2 coming from capture processes will contain a range of impurities as none

of the technologies are efficient enough to produce pure CO2 [13]. The concentration and

type of the impurities will depends on many factors like fuel type, capture technology and the

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design of the plant [14]. The presence of impurities changes thermophysical properties of the

stream such as density, viscosity, specific heat capacity, compressibility and critical pressure,

which can have a significant effect on the hydraulic behaviour of the CO2 stream. [15]

To understand thermodynamic properties of different CO2-mixtures, study of Equations of

State (EoS) is extremely important. In literature several references are available on EoS

suitable/used for CO2 and CO2 mixtures [16] [17] [18] [19], nevertheless a suitable equation

of state for mixtures in appropriate conditions for pipeline transport, in particular with a high

CO2 concentration, has not been clearly defined [20].

The pressure-volume-temperature behaviour, thermal properties and densities of sour

gas/liquid mixtures are essential to proper design of sour natural gas processing units, CO2

transport and injections as well as CCS scheme.

In this communication, new density data are presented for the 0.0495 H2S + 0.9508 CO2

binary system. In total, five sets of isothermal density data have been measured at 273, 283,

298, 323 and 353 K and pressures up to 41 MPa. The new experimental data are compared to

literature data and predictions using the Peng-Robinson (PR-ES) [21], Soave-Redlich-Kwong

(SRK-EoS) [22], Valderrama modification of the [23] Patel-Teja (VPT-EoS) [24] equations

of state. The CO2 volume correction [25] and Peneloux volume translation [26] have been

introduced to these equations of state to improve the density prediction. Finally, apart from

the cubic equations of state, the GERG-2008 EoS[27] which a wide-range multi parameter

equation of state proposed, has also been evaluated using the new experimental data.

2. Literature review

Despite the importance of the volumetric behaviour of acid gases and liquids, due to the

toxicity of the system, limited experimental data are available in the literature on the density

of CO2-H2S binary systems. The first study on the pressure-volume-temperature behaviour of

the CO2-H2S systems has been conducted by Bierlein and Kay [28] for the saturated states of

eight mixtures of CO2-H2S from 273.15 K to the critical temperature of hydrogen sulfide.

Mole fractions of CO2 in this study are 0, 0.0630, 0.1614, 0.2608, 0.3759, 0.4728, 0.6659,

0.8292, 0.9009 and 1. Sobocinski and Kurata [29] experimentally studied the CO2-H2S

systems at lower temperatures, i.e., below 273.15 K, to extend the data obtained by Bierlein

and Kay [28]. They studied seven mixtures from the solid-liquid-vapour region to the critical

region. A research report from the Gas Processors Association about thermodynamic

properties of CO2-H2S mixtures [30] includes density measurements for four mixtures. The

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systems studied by Kellerman et al. [30] were 93.93 mol% CO2 + 6.07 mol% H2S, 90.45

mol% CO2 + 9.55 mol% H2S, 70.67 mol% CO2 + 29.33 mol% H2S and 50.01 mol% CO2 +

49.99 mol% H2S. They performed Burnett and Burnett-isochoric measurements for each

mixture at temperatures ranging from 200 K to 450 K and at pressures from 0.1 to 23 MPa.

All the results of these research reports were also published by Stouffer et al. [31]. The data

published by Kellerman et al. [30] have been evaluated using different cubic equations of

state by Boyle et al. [32] in order to simulate natural gas treatment processes. They have

compared the SRK, SRK-Peneloux, PR, PR-Peneloux, PR-Mathias and Patel-Teja (PT)

equations of state. The overall reported Absolute Average Deviations (AAD) were 3.55, 2.64,

2.27, 2.68, 3.28 and 1.82%, respectively. The original SRK EoS in the vapour region and PT

in the liquid and supercritical regions have the best predictions with the AAE of 0.51, 2.16

and 2.26%, respectively. The phase behaviour of CO2-H2S has also been studied by Chapoy

et al. [33] in order to improve the design of acid gas injection in CCS scheme. The

measurements included VLE tests at 258.41 K, 273.15 K, 293.47 K and 313.02 K at

pressures from 1.0 to 5.5 MPa. They also have mentioned that both PR and SRK equation of

states with classical mixing rules and adjusted binary interaction parameters can predict the

phase behaviour of CO2-H2S systems.

3. Experimental part

3.1 Material

Table 1 lists the chemicals used to prepare the binary mixture. To prepare the binary mixture,

first, the component with the lower vapour pressure at laboratory condition was injected to

the pressure vessel. At about 20 °C (293 K), the vapour pressure of H2S and CO2 are 1.78 and

5.73 MPa, respectively. Therefore, for preparing the CO2+H2S mixture, first H2S was injected

to the high pressure vessel. The exact weight of the injected pure component was measured

using a four digit balance (Model: Sartorius) three times. The average of the readings was

calculated and accordingly the calculated amount for the second pure component was injected

to the high pressure vessel. After injecting both pure components, weight percent and mole

percent of the mixture were calculated. The mole percent for H2S was 4.95 mol% with U(z) =

0.008 mol% and for CO2 was 95.05 mol% with U(z) = 0.286 mol%. The uncertainties are

reported with 95% level of confidence, i.e., k=2.

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3.2 Equipment description

Densities of CO2-H2S system were measured using a high temperature and pressure Vibrating

Tube Densitometer (VTD), Anton Paar DMA 512 and using forced path mechanical

calibration (FPMC) model which is well described by Bouchot et al. [34], Khalil et al. [35]

and Coquelet et al. [36]. A schematic view of the apparatus is shown in Figure 1. A detailed

description of a typical vibrating-tube densitometer is given by Bouchot and Richon [37]. The

main part of the set-up is the Anton Paar DMA 512 densitometer. The U-shape vibrating tube

made from Hastelloy which can work up to 70 MPa pressure and temperature range of 263.15

K to 423.15 K (-10 °C to 150 °C). Temperature of the densitometer can be set by a liquid

bath (model: Lauda RE206) keeping the temperature stable to ±0.01 K. The setup is fully

immersed in a liquid bath (model: West P6100) to keep the temperature constant during the

test. Temperature in the vibrating tube part and liquid bath is measured using four-wire 100-

Ω platinum resistance probes (Pt100) which were calibrated against a 25-Ω reference

thermometer of Tinsley Precision Instrument. The period of the vibration, τ, is recorded with

a HP53131A data acquisition unit. Three pressure transducers (model: Druck PTX611) with

different complementary ranges of 0-10 MPa, 10-30 MPa and 30-70 MPa were used to record

the pressure data by connection to the HP34970A data acquisition unit. All the parts of the

apparatus are connected together by 1.6 mm (1/16 inch) tubes.

3.3 Calibration procedure

First, three pressure transducers and three temperature sensors were calibrated. The model of

pressure transducers is Druck PTX611 with different ranges of 0-10 MPa, 10-30 MPa and 30-

70 MPa. The pressure transducers were calibrated using an electronic balance GE Sensing

PACE 5000 for pressures up to 20 MPa and a dead weight tester Desgranges & Huot 5202S

for pressures from 20 MPa to 40 MPa. The accuracies on pressure measurements after

calibration are ±0.002, ±0.005 and ±0.005 MPa for the pressure transducers ranging from 0-

10 MPa, 10-30 MPa and 30-40 MPa, respectively. Also, two temperature probes in the

vibrating tube part and liquid bath were calibrated against a 25-Ω reference thermometer

(model: Tinsley Precision Instrument). The accuracies on temperature after calibration are

estimated to be ±0.02 K.

The Calibration method for the vibrating tube densitometers using a forced path mechanical

calibration (FPMC) model for the Anton Parr DMA 512 densitometer were well described by

Bouchot and Richon [34]. The U-tube of the DMA 512 densimeter cell can be modelled by

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means of a linear hollow vibrating system whose internal volume is Vi = Vi(T,P) and total

vibrating mass is (M0+ρV i). M0 is the proper mass of the tube under vacuum and ρ the

density of the inner fluid to be determined. This system has a natural transversal stiffness K =

K (T, P) and is vibrating with a period () in an undamped harmonic way under the effect of a

mechanical excitation.

The relation between these quantities can be written in the following general form:

20

20

ρ 1

= − i o

M K

V K

ττ

(1)

where K0 and 0 are the transversal stiffness and the vibrating period of the evacuated tube,

respectively.

The term K/K0 is calculated as:

3

0

0 0

=

LK I

K I L (2)

The first term in the equation above is expressed as:

4 44

4 4 40 0 0 0

−∆= =∆ −

e i

e i

r rI r

I r r r (3)

In Equation (3), the subscription i and e denotes the internal and external radius, respectively.

The second term of the equation is the relative inverse cubic length change expressed by:

( )

3

0 exp 3 = −

T

LP

Lγ

(4)

γT is one of the two unknown parameters that remain in the FPMC model.

The term (M0/V i) can be expressed as:

( )0 0 0

2 200

1

,

= =

i i i

M M M

V T P r L L r Lπ π δ (5)

where (M0/L00) is the second unknown parameter of these equations.

The complete FPMC model can be formulated as:

( )4 2

02 4 2

00 0

1 ( , ) ( , )( , ) exp 3 1

( , ) ( , ) ( ) ( )

∆ = − − ∆ T

i o

M r T P T PT P P

L r T P L T P r T T

τρ γπ δ τ

(6)

The two unknown parameters of γT and (M0/L00) were optimised using the density data of

pure CO2 from Span and Wagner equation of state [16] at across the full desired pressure

ranges and for each measured isotherms.

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3.4 Measurement procedure

All the experiments were conducted using the Anton Paar DMA 512 densitometer. The

maximum desired pressure to conduct experiments was 40 MPa. The vessel for the mixture

preparation has 150 ml volume and the design pressure is 70 MPa. It also includes a piston to

isolate the mixture from the gas injected to pressurise the system.

To conduct each test, after stabilising the desired temperature the entire system was

vacuumed. The vacuumed vibrating period, vacuum pressure and temperature then were

recorded at the desired temperatures (273, 283, 298, 323 and 353 K). The sample then was

injected through the pressurised vessel into the densitometer as a gas phase by opening a

regulated valve at a slow rate till reaching the dew point pressure. For measuring the density

in the liquid phase, the pressure was increased to the maximum desired pressure, i.e., about

40 MPa, then the pressure reduced slowly by opening the outlet valve till reaching the bubble

point. The toxic outlet gas during the depressurisation was neutralised in a column containing

a basic solution (Sodium Hydroxide, NaOH). During the injection and depressurisation steps,

pressure, temperature and period were recorded to calculate the density. To promote the dew

point of the system during the injection of sample, as the first bubble should be produced

inside the densitometer, the temperature of bath controlling the densitometer temperature was

set to be slightly lower (0.2 - 0.3 K) than the temperature of liquid bath for the entire system.

However, to promote the bubble point, the densitometer temperature was set to be slightly

higher than the entire system temperature.

The combined standard uncertainties [38][39] of density measurements for each measured

quantity have been detailed in Appendix A.

4. Modelling

Density prediction using classical cubic equations of state are normally lower than the real

density of fluid particularly in the liquid phase [40]. This is due to the unrealistic critical

compressibility assumption for all compounds in the cubic equations of state [41][42]. To

improve this underestimation in the density, Peneloux et al. [26] introduced a volume

translation parameter to the overestimated molar volume predicted by equations of state. In

this work, to improve the overestimation of molar volume by equations of state, a volume

correction parameter, called CO2 volume correction parameter [25], was subtracted from the

molar volume of carbon dioxide and mixtures containing high CO2 concentrations calculated

by the cubic equations of state.

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new EoS cV V V= − (7)

where EoSV is the molar volume calculated by the desired equation of state. The volume

correction parameter in the Equation (7),cV , is given below:

NCompc c

i ii

V x V= ∑ (8)

zi in the above equation is the composition of component i and Vc is the molar volume

predicted by Equation of state for each composition in the same phase. As this volume

correction parameter is applicable for the CO2-rich mixtures, the molar volume of CO2, c

iV is

defined by:

2 2 2

c Eos mBWRCO Pure CO Pure COV V V= −

(9)

The molar volume for pure CO2 by mBWR equation of state can be calculated from the

equation proposed by McCarty [43][44] below.

29 15

n 2 n 17n n

n 1 n 10

P a (T) a (T) e− − γρ

= =

= ρ + ρ∑ ∑ (10)

In this work, the Peng-Robinson [21], Soave-Redlich-Kwong [22] ,Valderrama [23] modified

version of Patel-Teja [24] and the GERG-2008 equations of state were employed to predict

the densities as implemented in-house thermodynamic package, HWPVT [25][33][45][46].

The modified binary interaction parameters, BIP or kij, of CO2-H2S have been employed to

improve the phase equilibrium predictions. The values for PR, SRK and VPT equations of

state are 0.095, 0.098 and 0.090483, respectively [47]. Figure 2 shows the Pressure –

composition diagram of the CO2 (1) + H2S (2) system [33].

5. Specific heat capacity calculations

The isobaric heat capacity is the temperature derivative of enthalpy at constant pressure. The

residual specific heat capacity ( − ) can be obtained from the measured densities. The

thermodynamic relationship between isobaric heat capacity and pressure-temperature-volume

data is as following [48]:

2p

2

PT

CT

P T

∂ ∂ ν= − ∂ ∂ (11)

We have considered the molar volume as a polynomial function of temperature at each

constant pressure:

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2aT bT cν = + + (12)

2

22aT b 2a

T T

∂ν ∂ ν= + → =∂ ∂ (13)

By replacing Equation (13) in Equation (11) and integrating, the specific heat capacity can be

calculated from the following equation:

p

pi

C P2p

p2C P 0PT

0p p

CT 2aT dC 2aT dP

P T

C C 2aTP C

=

∂ ∂ ν ∫= − = − → = − ∂ ∂

− = − +

∫ ∫ (14)

The constant C can be calculated from the reference fluid experiments, i.e. pure CO2, by the

following equations:

p

CO2

2 CO 22

pi

CO2 CO2 2 CO CO 22 2

C P2p

CO p CO2C P 0PT

0 0p p CO p p CO

CT 2a T dC 2a T dP

P T

C C 2a TP C C C (C ( 2a TP))

=

∂ ∂ ν ∫= − = − → = − ∂ ∂

− = − + ⇒ = − + −

∫ ∫ (15)

Where CpCO2 is the heat capacity of pure CO2 obtained from Span and Wagner equation of

state [16]. Also, 0piC is the ideal gas specific heat capacity and can be calculated from the

equation presented by Aly and Lee (1981) [49]:

( ) ( ) ( )

2 2

0

cos

+ +

=

pi

D FT TC T B C ED Fsinh hT T

(16)

The constants B through F in this equation can be found from Table 2 for the pure

components CO2 and H2S.The calculated heat capacity from the equation above would be in

J.K-1.kmol-1. The ideal gas heat capacity for the mixtures can be calculated using the following

mixing rule:

( )mix i

N

ii 1

0 0p pyC T C

=

=∑ (17)

6. Results and discussions

The densities of CO2-H2S system were measured continuously using a high temperature and

pressure Vibrating Tube Densitometer (VTD), Anton Paar DMA 512. First, pure CO2 were

used to calibrate the densitometer with a forced path mechanical calibration (FPMC)

technique [34]. Then, the densities of a binary system (95.05 mol% CO2 + 4.95 mol% H2S)

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were measured at pressures up to 41.4 MPa at five different temperatures, 273.15, 283.15,

298.15, 323.15 and 353.15 K (0, 10, 25, 50 and 80 °C) in gas, liquid and supercritical

regions. The uncertainty of measurements [38][39] also has been calculated and reported for

all measured data. The thermodynamic properties, i.e., compressibility factor, specific heat

capacity, dew and bubble points have been derived from the measured density data. In the

modelling part, the experimental density data were used to evaluate the CO2 volume

correction model by comparing to the cubic equations of state (PR [21], SRK [22] and VPT

[23]) and Peneloux shift parameter [26]. Also, the measured density data and derived

thermodynamic properties have been employed to evaluate the GERG-2008 [27] equation of

state.

The measured density data as well as their uncertainty and compressibility factor at

corresponding pressure, temperature and phase are declared in Table 3 to Table 7. Each table

shows the measured data at one isotherm. Also, Figure 3 illustrates the experimental densities

measured in this work and predicted densities using GERG-2008 [27] at all measured

temperatures at total pressure ranges. Those at low pressures up to 10 MPa and densities

below 200 kg/m3 are shown on Figure 4. In Figure 3, there are few gaps in the densities

measured at 322.4 K. As explained in the measurement procedure above, the densities were

measured by reducing the pressure through opening the outlet valve. The gap of the data are

due to the sudden release of pressure from the outlet valve.

To validate the measured densities of the investigated system, the densities measured at 322.4

K and 353 K have been compared to the data published by Stouffer et al. [30][31] at two

different isotherms at 325 K and 350 K. The mole fraction of H2S in this work is 0.0495 and

those at work by Stouffer et al. are 0.0607 and 0.0955. Figure 5 and Figure 6 demonstrate the

comparisons at total pressure range and low pressures, respectively. The lines in these figures

are the predicted densities using the GERG-2008 [27] and PR-CO2 in this work. As can be

seen in these figures, the measured densities in this work are in the same range reported by

Stouffer et al. [30][31]. Both measured data also are in good agreement with the predicted

densities using the GERG-2008 and PR-CO2 in this work. Figure 7 presents the deviations of

the GERG-2008 EoS [27] from experimental densities of CO2 - H2S at 322.4 K measured in

this work and at 325 K measured by Stouffer at al. [30][31]. As can be seen in this figure, the

densities in the range from 200 – 700 kg/m3 is in the area close to the critical condition and

the deviation is higher in this area.

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In Table 3 to Table 7, type B uncertainties [38][39] for each measured density are reported

with 95% level of confidence, coverage factor k = 2. The highest uncertainties are in the gas

phase at lower temperature and low pressures as expected. Also the uncertainties of

measurements in the liquid phase near two-phase region and critical point are high. The

average uncertainties in the gas phase is 0.4% with the maximum value of 1.8% at the lowest

temperature of 273 K. While, those in the liquid phase are 0.03% and 0.3%, respectively. The

maximum uncertainty in the liquid phase occurred at 298 K which is close to the critical point

of the mixture. The average and maximum uncertainties in the supercritical region are 0.1%

and 0.4%, respectively.

The Average Absolute Deviations (AAD) shows the deviations of the predicted values from

the experimental results which has been calculated using Equation (18) and summarised in

Table 9 for the measured system.

% = . − .

× 100 (18)

The prediction accuracy of classical cubic equations of state (PR, SRK and VPT) has been

investigated in this work by comparing to the experimentally measured densities of CO2-H2S

system. The modified binary interaction parameters for CO2-H2S in each equation of state

have been employed (kij = 0.095, 0.098 and 0.090483 for PR, SRK and VPT, respectively)

[47]. As Table 9 illustrated, the SRK EoS with the AAD of 1.3% is the most accurate EoS to

predict the densities of CO2-H2S in the gas region while both PR and VPT predicts quite well

the densities in the gas region with the AADs of 1.5% and 1.4%, respectively. In the liquid

and supercritical regions, the VPT EoS predicts the densities of the investigated system with

the AAD of 3.0% in the liquid and 1.8% in the supercritical regions. Those for the PR and

SRK are higher than VPT EoS (4.2% and 8.0% in the liquid region, 2.1% and 8.0% in the

supercritical region for PR and SRK, respectively). Overall, the VPT EoS with AAD of 2.3%

predicts the densities better than PR (AAD = 3.0%) and SRK (AAD = 6.3%) for this binary

mixture.

The effect of applying CO2 volume correction to cubic equations of state in this work could

be analysed by comparing each equation of state and those with CO2 volume correction. This

modification in the volume prediction improved the density prediction significantly. As can

be seen from Table 9, the overall AAD for PR and PR-CO2 are 3.0% and 1.5%, respectively.

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The Maximum Absolute Deviation (MAD) for the PR is 10.6% while that of PR-CO2 is 4.9%

which indicates the improvements in density predictions. The improvement in density

prediction in the liquid phase is more sensible as the AAD reduced from 4.2% to 1.3% by

applying the CO2 volume correction parameter. Also, the AAD for PR with Peneloux shift

model [26] is 2.1% which indicates that PR-CO2 predicts more accurate than PR-Peneloux.

The improvements in density prediction using SRK EoS by applying CO2 volume correction

term is even higher than PR EoS, particularly in the dense liquid / supercritical regions.

Applying the CO2 volume correction to SRK EoS reduces the AAD from 8.0% to 1.6% and

1.3% in the liquid and supercritical regions, respectively. Moreover, the SRK-CO2 (overall

AAD = 1.4%) predicts more accurate than SRK-Peneloux with the overall AAD of 2.6%. The

MAD for SRK and SRK-CO2 are 19.4% and 4.7%, respectively. Also, the overall AAD of

density predictions by VPT EoS is reduces from 2.3% to 1.4% by applying the CO2 volume

correction term. The most improvements can be seen in the liquid region where the AAD

reduces from 23% to 1.2% by applying CO2 correction term to VPT EoS. The MAD for VPT

and VPT-CO2 are 11.6% and 5.4%, respectively. All equations of state with CO2 volume

correction term predicts with the same accuracy in the gas region (AAD = 1.0%) while VPT-

CO2 in the liquid region (AAD = 1.2%) and SRK-CO2 in the supercritical region (AAD =

1.3%) predict better than other equations of state. The overall AAD for PR-CO2, SRK-CO2

and VPT-CO2 are 1.5%, 1.4% and 1.4% where indicates that almost all of the equations of

state investigated in the work predicts with the same accuracy for the investigated system.

Figure 8 shows the deviations of the predictions using PR-CO2 equation of state from the

measured densities in this work.

Apart from the classical cubic equations of state which are very common in the industries due

to the simplicity and availability of them, the GERG-2008 [27] also has been evaluated in this

work by comparing to the measured densities of CO2-H2S binary mixture. As seen in Table 9

the GERG-2008 is as accurate as the cubic equations of state with CO2 volume correction in

the gaseous region with the AAD of 1.0% while it has the best predictions in comparison to

the cubic equations of state in the liquid and supercritical regions with the AAD of 1.0% and

1.3%, respectively. Overall, the GERG-2008 with an AAD of 1.1% predicts the best among

the equations of state studied during the course of this study.

The compressibility factor, or Z-Factor, defined as a ratio of the gas at high pressure and

temperature conditions to the ideal gas volume [42] which shows the deviation of gas

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behaviour from the ideal gas. Z-factor is an imperative property in the design of processing

units, compression stations, pipeline design [50][51] and flow metering [52]. In engineering

purposes and design considerations, compressibility factor as a thermodynamic property can

be determined from either equations of states, predictive corresponding state equations [52]

or experimentally derived correlations [50][53]. However, the accuracy of these methods

should be evaluated by comparing to the experimentally measured data. In this work, the

compressibility factors of the system have been obtained from the measured density data for

each isotherms and are reported on Table 3 to Table 7. Figure 9 and 10 show the measured

compressibility factors as well as predicted values using GERG-2008 EoS [27] over the full

pressure range and at low pressures, respectively. As can be seen in these figures, the

predicted values using the GERG-2008 EoS are in good agreement with the experimentally

obtained data with an overall AAD of 1.1%.

The specific heat capacities were calculated using Equations (11) through Equation (17) at

different pressures of 40, 35, 30, 25, 20 and 15 MPa at five measured temperatures. Table 10

and Figure 11 show the calculated specific heat capacity for this system. The AAD between

the predicted values using GERG-2008 and the experimentally calculated isobaric heat

capacity is 4.8% with a MAD of 9.0%. The model predicts the specific heat capacity higher

than experimental data. It is also worth to note that the isobaric specific heat capacities were

calculated from the measured density data and is not a direct measurement. Therefore, it is

suggested to compare these result with the directly measured data using calorimeters.

Numerous data points were measured on both sides of the dew and bubble points at each

temperature. Analysis of the measured data points can be used in determination of the break

points, which corresponds to dew point or bubble point of the mixture at the measured

temperatures. The importance of the large number of recorded data points is that it allows

fitting the representative correlation to the experimental data in order to find the precise

location of dew and bubble points. Table 11 as well as Figure 12 and Figure 13 show the

measured dew and bubble points for CO2 - H2S binary system. The measured dew / bubble

points also have been compared to the prediction values using the PR-CO2 equation of state.

Table 11 also shows the prediction values in addition of the AAD from the experimental data.

As can be seen, the AAD for the dew/bubble point pressures and those densities are 3.8% and

2.7%, respectively.

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7. Conclusions

The densities of the 0.9505 CO2 – 0.0495 H2S binary system have been measured at different

isotherms 273, 283, 298, 323 and 353 K at pressures up to 41 MPa in the gas, liquid and

supercritical regions. The measured densities have been validated by comparing to the

measured densities by Stouffer et al. [30][31] at two different isotherms. The uncertainty of

measurements also reported and concluded that it can be high at low pressures in the gas

region and in vicinity of the phase boundary of two phase region. The overall average

uncertainty of measurements with 95% level of confidence is 0.2% with the highest quantity

of 1.8% in the gas region. Moreover, by applying thermodynamic equations, the

compressibility factor and specific heat capacities of the investigated system were obtained

from the measured densities at different pressures and temperatures. Numerous

measurements of the density on both sides of dew and bubble point resulted in obtaining the

dew and bubble points for investigated system. In the modelling part, the predictions using

classical cubic equations of states has been compared and concluded that VPT equation of

state with the AAD of 2.3% predicts more accurate than PR and SRK equations of state. To

improve the density predictions by these equations, the CO2 volume correction term then has

been applied to them. It has been concluded that introducing CO2 volume correction to

equations of state can improve the prediction accuracy for the acid gas and liquid system

particularly in the liquid phase. The accuracy of equations of state with CO2 correction term

are approximately as same as each other with the AAD of 1.4% for SRK-CO2 and VPT-CO2

and 1.5% for PR-CO2. Apart from the cubic equations of states, the prediction by GERG-

2008 has been evaluated by comparing to the experimental data. The conclusion from this

evaluation is that GERG-2008 equation of state has the best prediction with the AAD of 1.1%

among the equations of states investigated during the course of this work. The derived

thermodynamic properties, i.e., compressibility factor and isobaric heat capacity, as well as

dew / bubble points have been compared to the predictions by models. From the

comparisons, it has been concluded that there are a good agreement between the derived

properties and predicted values. The AAD of predictions by GERG-2008 are 1.1% and 4.8%

for compressibility factor and isobaric heat capacities, respectively. For the dew/bubble point

pressure and those densities, the AAD of predictions by PR-CO2 are 3.8% and 2.7%,

respectively.

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8. Acknowledgements

This work is a part of the joint industrial project titled “Impact of Common Impurities on

Carbon Dioxide Capture, Transport and Storage” [54] where conducted jointly at Hydrates,

Flow Assurance & Phase Equilibria Research Group, Institute of Petroleum

Engineering, Heriot-Watt University in Edinburgh, UK and CTP - Centre Thermodynamique

des Procédés, MINES ParisTech in Fontainebleau, France. The physical properties to be

investigated during the project are phase equilibira, hydrates, solid formation, density,

viscosity, interfacial tension, solubility in brine and pH. The first phase of project was

conducted from October 2011 to October 2014. The JIPs was supported by Chevron, GALP

Energia, Linde AG Engineering Division, OMV, Petroleum Expert, Statoil, TOTAL and

National Grid Carbon Ltd, which is gratefully acknowledged. The participation of National

Grid Carbon in the JIP was funded by the European Commission’s European Energy

Programme for Recovery. The authors would also like to thank the members of the steering

committee for their fruitful comments and discussions.

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Table 1 - Details of the chemicals, suppliers and purities of the components used in this study.

Chemical Name Source Initial Purity a Certification Analysis Method b H2S Air Liquide 0.995 vol Air Liquide Certified SM

CO2 Air Liquide 0.99995 vol Air Liquide Certified SM a No additional purification is carried out for all samples. b SM: Supplier method

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Table 2 - Constants B through F in the Equation (16) by Aly and Lee [49]

Cpi constants B C D E F

CO2 2.94E+04 3.45E+04 -1.43E+03 2.64E+04 5.88E+02

H2S 3.33E+04 2.61E+04 9.13E+02 -1.80E+04 9.49E+02

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Table 3 - Experimental results of the 0.9505 mole CO2 + 0.0495 mole H2S system at 273 K a

No Phase T/K p/MPa ρ/kg.m-3 Z No Phase T/K p/MPa ρ/kg.m-3 Z

Exp. uc(ρ) Exp. uc(ρ) 1 Gas 272.55 0.301 5.9 0.1 0.98 48 Liquid 272.58 27.519 1017.1 0.2 0.52 2 Gas 272.55 0.402 8.0 0.1 0.96 49 Liquid 272.59 26.955 1015.2 0.2 0.51 3 Gas 272.55 0.501 9.9 0.1 0.97 50 Liquid 272.59 26.517 1013.8 0.3 0.50 4 Gas 272.54 0.603 12.1 0.1 0.96 51 Liquid 272.58 26.025 1012.1 0.3 0.49 5 Gas 272.54 0.892 18.5 0.1 0.93 52 Liquid 272.58 25.470 1010.3 0.3 0.48 6 Gas 272.54 1.006 20.9 0.1 0.92 53 Liquid 272.58 24.975 1008.5 0.3 0.48 7 Gas 272.56 1.104 23.0 0.1 0.92 54 Liquid 272.58 24.527 1006.9 0.2 0.47 8 Gas 272.56 1.296 27.7 0.1 0.90 55 Liquid 272.58 23.985 1004.9 0.2 0.46 9 Gas 272.56 1.499 32.4 0.1 0.89 56 Liquid 272.58 23.514 1003.2 0.2 0.45 10 Gas 272.55 1.696 37.5 0.1 0.87 57 Liquid 272.58 23.027 1001.3 0.3 0.44 11 Gas 272.56 1.905 42.8 0.1 0.85 58 Liquid 272.58 22.532 999.5 0.3 0.43 12 Gas 272.56 2.003 45.2 0.2 0.85 59 Liquid 272.58 22.003 997.4 0.2 0.42 13 Gas 272.56 2.205 51.1 0.2 0.83 60 Liquid 272.58 21.490 995.3 0.3 0.41 14 Gas 272.56 2.407 56.8 0.2 0.81 61 Liquid 272.58 20.990 993.2 0.3 0.41 15 Gas 272.56 2.502 59.8 0.2 0.80 62 Liquid 272.58 20.515 991.4 0.3 0.40 16 Gas 272.55 2.709 66.4 0.2 0.78 63 Liquid 272.58 20.011 989.1 0.3 0.39 17 Gas 272.56 2.906 73.0 0.2 0.76 64 Liquid 272.58 19.576 987.6 0.3 0.38 18 Gas 272.55 2.996 76.3 0.2 0.75 65 Liquid 272.58 19.012 985.2 0.3 0.37 19 Gas 272.55 3.196 83.9 0.2 0.73 66 Liquid 272.58 18.540 983.3 0.3 0.36 20 Liquid 272.57 41.417 1058.1 0.2 0.75 67 Liquid 272.58 18.046 981.3 0.3 0.35 21 Liquid 272.58 41.400 1058.2 0.1 0.75 68 Liquid 272.57 17.556 979.1 0.3 0.34 22 Liquid 272.57 40.955 1057.0 0.2 0.74 69 Liquid 272.57 17.016 976.8 0.3 0.33 23 Liquid 272.57 40.046 1054.9 0.2 0.73 70 Liquid 272.58 16.524 974.5 0.3 0.33 24 Liquid 272.58 39.463 1053.2 0.2 0.72 71 Liquid 272.58 16.022 972.1 0.3 0.32 25 Liquid 272.57 39.008 1052.0 0.2 0.71 72 Liquid 272.58 15.502 969.5 0.3 0.31 26 Liquid 272.56 38.445 1050.7 0.2 0.70 73 Liquid 272.58 15.013 967.1 0.3 0.30 27 Liquid 272.56 38.049 1049.5 0.2 0.70 74 Liquid 272.59 14.525 964.4 0.3 0.29 28 Liquid 272.56 37.557 1048.1 0.2 0.69 75 Liquid 272.58 14.024 962.0 0.3 0.28 29 Liquid 272.56 36.988 1046.6 0.2 0.68 76 Liquid 272.59 13.502 959.4 0.3 0.27 30 Liquid 272.56 36.561 1045.5 0.2 0.67 77 Liquid 272.58 13.010 956.8 0.3 0.26 31 Liquid 272.57 36.058 1044.1 0.2 0.66 78 Liquid 272.58 12.531 954.3 0.3 0.25 32 Liquid 272.57 35.493 1042.3 0.3 0.65 79 Liquid 272.59 11.999 951.2 0.3 0.24 33 Liquid 272.56 35.060 1041.1 0.3 0.65 80 Liquid 272.59 11.511 948.4 0.4 0.23 34 Liquid 272.57 34.499 1039.5 0.2 0.64 81 Liquid 272.59 11.005 945.1 0.4 0.22 35 Liquid 272.57 33.999 1038.0 0.2 0.63 82 Liquid 272.59 10.504 942.3 0.4 0.21 36 Liquid 272.57 33.498 1036.6 0.3 0.62 83 Liquid 272.59 10.064 939.4 0.4 0.21 37 Liquid 272.58 33.011 1034.9 0.2 0.61 84 Liquid 272.60 9.499 936.2 0.3 0.19 38 Liquid 272.57 32.499 1033.1 0.2 0.60 85 Liquid 273.23 8.730 929.2 0.3 0.18 39 Liquid 272.58 31.936 1031.5 0.2 0.59 86 Liquid 273.29 8.504 926.7 0.3 0.18 40 Liquid 272.58 31.446 1030.1 0.3 0.59 87 Liquid 273.30 7.996 922.9 0.3 0.17 41 Liquid 272.57 31.002 1028.8 0.2 0.58 88 Liquid 273.31 7.524 919.4 0.3 0.16 42 Liquid 272.57 30.442 1027.2 0.2 0.57 89 Liquid 273.31 7.043 915.9 0.3 0.15 43 Liquid 272.57 30.022 1025.6 0.2 0.56 90 Liquid 273.31 6.504 911.6 0.3 0.14 44 Liquid 272.58 29.540 1023.9 0.2 0.55 91 Liquid 273.32 5.999 907.1 0.3 0.13 45 Liquid 272.58 29.015 1022.3 0.2 0.54 92 Liquid 273.31 5.510 902.8 0.3 0.12 46 Liquid 272.58 28.510 1020.5 0.2 0.54 93 Liquid 273.34 5.004 897.4 0.4 0.11 47 Liquid 272.58 28.014 1018.9 0.2 0.53 94 Liquid 273.33 4.504 892.6 0.4 0.10

a Standard uncertainties, u, are u(T) = 0.02 K, u(p) = 0.002 MPa for pressures up to 10 MPa and u(p) = 0.005 MPa for pressures from 10-40 MPa

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Table 4 - Experimental results of the 0.9505 mole CO2 + 0.0495 mole H2S system at 283 K a

No Phase T/K p/MPa ρ/kg.m-3 Z No Phase T/K p/MPa ρ/kg.m-3 Z

Exp. uc(ρ) Exp. uc(ρ) 1 Gas 282.69 0.600 11.8 0.1 0.94 49 Liquid 283.47 27.510 998.9 0.3 0.51 2 Gas 282.70 0.820 16.6 0.1 0.91 50 Liquid 283.47 27.006 996.9 0.3 0.50 3 Gas 282.70 1.004 20.1 0.1 0.92 51 Liquid 283.46 26.500 994.9 0.3 0.49 4 Gas 282.70 1.194 25.0 0.1 0.88 52 Liquid 283.47 26.012 993.0 0.3 0.48 5 Gas 282.69 1.405 29.0 0.1 0.90 53 Liquid 283.47 25.518 991.0 0.3 0.48 6 Gas 282.70 1.631 35.0 0.1 0.86 54 Liquid 283.47 25.000 988.9 0.3 0.47 7 Gas 282.69 1.794 39.0 0.1 0.85 55 Liquid 283.47 24.503 986.9 0.3 0.46 8 Gas 282.70 2.040 45.1 0.1 0.84 56 Liquid 283.47 24.005 984.8 0.3 0.45 9 Gas 282.71 2.196 49.1 0.1 0.83 57 Liquid 283.47 23.501 982.6 0.3 0.44 10 Gas 282.70 2.423 55.3 0.2 0.81 58 Liquid 283.47 23.016 980.5 0.3 0.43 11 Gas 282.71 2.634 61.3 0.2 0.80 59 Liquid 283.47 22.501 978.3 0.3 0.42 12 Gas 282.71 2.835 67.2 0.2 0.78 60 Liquid 283.47 22.004 976.1 0.3 0.42 13 Gas 282.71 3.018 72.9 0.2 0.77 61 Liquid 283.47 21.510 973.8 0.3 0.41 14 Gas 282.71 3.243 80.2 0.2 0.75 62 Liquid 283.47 21.001 971.5 0.3 0.40 15 Gas 282.70 3.401 85.8 0.2 0.73 63 Liquid 283.47 20.508 969.2 0.3 0.39 16 Gas 282.71 3.592 92.8 0.2 0.72 64 Liquid 283.47 20.010 966.8 0.3 0.38 17 Gas 282.71 3.803 101.2 0.3 0.70 65 Liquid 283.47 19.503 964.5 0.3 0.37 18 Gas 282.72 4.012 110.3 0.3 0.67 66 Liquid 283.47 19.008 962.0 0.3 0.36 19 Gas 282.72 4.099 114.3 0.3 0.66 67 Liquid 283.47 18.516 959.4 0.3 0.36 20 Gas 282.71 4.201 119.4 0.3 0.65 68 Liquid 283.47 18.016 957.0 0.3 0.35 21 Gas 282.70 4.303 125.3 0.4 0.64 69 Liquid 283.47 17.517 954.5 0.3 0.34 22 Gas 282.73 4.361 133.9 0.5 0.60 70 Liquid 283.47 17.011 951.8 0.4 0.33 23 Liquid 283.47 40.443 1041.5 0.2 0.72 71 Liquid 283.47 16.506 949.0 0.4 0.32 24 Liquid 283.47 40.004 1040.3 0.2 0.71 72 Liquid 283.48 16.008 946.2 0.4 0.31 25 Liquid 283.47 39.503 1038.9 0.2 0.70 73 Liquid 283.47 15.509 943.3 0.4 0.30 26 Liquid 283.46 39.000 1037.5 0.2 0.69 74 Liquid 283.47 15.004 940.4 0.4 0.29 27 Liquid 283.47 38.528 1036.2 0.2 0.69 75 Liquid 283.48 14.511 937.3 0.4 0.29 28 Liquid 283.46 38.004 1034.6 0.2 0.68 76 Liquid 283.47 14.002 934.4 0.4 0.28 29 Liquid 283.47 37.525 1033.1 0.2 0.67 77 Liquid 283.47 13.504 931.2 0.4 0.27 30 Liquid 283.47 37.000 1031.5 0.3 0.66 78 Liquid 283.47 13.000 928.0 0.4 0.26 31 Liquid 283.46 36.513 1030.0 0.2 0.65 79 Liquid 283.47 12.500 924.7 0.4 0.25 32 Liquid 283.47 36.014 1028.3 0.2 0.65 80 Liquid 283.48 12.002 921.3 0.4 0.24 33 Liquid 283.46 35.507 1026.8 0.2 0.64 81 Liquid 283.48 11.502 917.6 0.4 0.23 34 Liquid 283.46 35.004 1025.2 0.2 0.63 82 Liquid 283.49 11.002 914.1 0.4 0.22 35 Liquid 283.46 34.516 1023.7 0.2 0.62 83 Liquid 283.48 10.503 910.4 0.5 0.21 36 Liquid 283.47 34.012 1022.0 0.2 0.61 84 Liquid 283.48 10.062 907.0 0.5 0.20 37 Liquid 283.47 33.525 1020.3 0.2 0.61 85 Liquid 283.49 9.506 903.0 0.5 0.19 38 Liquid 283.47 33.004 1018.6 0.2 0.60 86 Liquid 283.49 9.003 898.8 0.5 0.18 39 Liquid 283.47 32.470 1017.0 0.2 0.59 87 Liquid 283.49 8.504 894.3 0.5 0.18 40 Liquid 283.46 32.001 1015.4 0.2 0.58 88 Liquid 283.49 8.001 889.8 0.5 0.17 41 Liquid 283.46 31.532 1013.7 0.2 0.57 89 Liquid 283.49 7.466 884.7 0.5 0.16 42 Liquid 283.46 31.005 1011.9 0.3 0.57 90 Liquid 283.49 7.137 882.0 0.6 0.15 43 Liquid 283.47 30.502 1010.1 0.2 0.56 91 Liquid 283.48 6.383 875.9 0.6 0.13 44 Liquid 283.47 30.014 1008.2 0.2 0.55 92 Liquid 283.47 5.962 871.5 0.6 0.13 45 Liquid 283.46 29.517 1006.4 0.2 0.54 93 Liquid 283.47 5.586 867.8 0.7 0.12 46 Liquid 283.47 29.025 1004.6 0.2 0.53 94 Liquid 283.52 5.041 861.2 0.7 0.11 47 Liquid 283.47 28.508 1002.7 0.2 0.52 95 Liquid 283.57 4.945 859.9 0.7 0.11 48 Liquid 283.48 28.010 1000.7 0.3 0.52 96 Liquid 283.92 4.600 826.5 0.8 0.10

a Standard uncertainties, u, are u(T) = 0.02 K, u(p) = 0.002 MPa for pressures up to 10 MPa and u(p) = 0.005 MPa for pressures from 10-40 MPa

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Table 5 - Experimental results of the 0.9505 mole CO2 + 0.0495 mole H2S system at 298 K a

No Phase T/K p/MPa ρ/kg.m-3 Z No Phase T/K p/MPa ρ/kg.m-3 Z Exp. uc(ρ) Exp. uc(ρ) 1 Gas 297.51 0.320 5.7 0.1 0.99 52 Liquid 298.06 27.528 944.3 0.3 0.51 2 Gas 297.51 0.652 11.8 0.1 0.97 53 Liquid 298.06 27.004 941.7 0.3 0.50 3 Gas 297.50 0.710 13.0 0.1 0.96 54 Liquid 298.08 26.531 939.6 0.3 0.50 4 Gas 297.50 0.878 16.3 0.1 0.95 55 Liquid 298.06 26.004 937.0 0.3 0.49 5 Gas 297.50 1.019 19.5 0.1 0.92 56 Liquid 298.06 25.499 934.6 0.3 0.48 6 Gas 297.51 1.232 23.7 0.1 0.91 57 Liquid 298.07 25.023 932.1 0.3 0.47 7 Gas 297.51 1.419 27.5 0.1 0.91 58 Liquid 298.07 24.506 929.6 0.3 0.46 8 Gas 297.51 1.601 31.3 0.1 0.90 59 Liquid 298.07 24.008 926.9 0.3 0.45 9 Gas 297.51 1.810 35.8 0.1 0.89 60 Liquid 298.07 23.511 924.2 0.3 0.45 10 Gas 297.51 2.003 40.0 0.1 0.88 61 Liquid 298.07 23.013 921.5 0.3 0.44 11 Gas 297.51 2.203 44.6 0.1 0.87 62 Liquid 298.07 22.513 918.8 0.3 0.43 12 Gas 297.51 2.413 49.5 0.1 0.86 63 Liquid 298.07 22.015 915.8 0.3 0.42 13 Gas 297.50 2.616 54.4 0.1 0.85 64 Liquid 298.07 21.500 912.9 0.4 0.41 14 Gas 297.50 2.810 59.4 0.1 0.83 65 Liquid 298.06 21.003 910.0 0.4 0.41 15 Gas 297.50 2.998 64.3 0.2 0.82 66 Liquid 298.08 20.503 907.0 0.4 0.40 16 Gas 297.50 3.212 70.1 0.2 0.81 67 Liquid 298.06 19.999 903.9 0.4 0.39 17 Gas 297.51 3.399 75.3 0.2 0.79 68 Liquid 298.07 19.497 900.5 0.4 0.38 18 Gas 297.50 3.611 81.5 0.2 0.78 69 Liquid 298.06 19.001 897.3 0.4 0.37 19 Gas 297.50 3.810 87.7 0.2 0.76 70 Liquid 298.08 18.497 893.7 0.4 0.36 20 Gas 297.51 4.001 93.7 0.2 0.75 71 Liquid 298.08 18.006 890.4 0.4 0.36 21 Gas 297.50 4.206 100.8 0.2 0.73 72 Liquid 298.07 17.508 886.9 0.4 0.35 22 Gas 297.51 4.402 107.8 0.2 0.72 73 Liquid 298.08 17.005 883.1 0.4 0.34 23 Gas 297.50 4.609 115.8 0.3 0.70 74 Liquid 298.08 16.505 879.3 0.4 0.33 24 Gas 297.51 4.802 123.7 0.3 0.68 75 Liquid 298.07 16.004 875.3 0.5 0.32 25 Gas 297.50 4.999 132.4 0.3 0.66 76 Liquid 298.07 15.503 871.1 0.5 0.31 26 Gas 297.50 5.204 142.4 0.3 0.64 77 Liquid 298.07 15.005 867.0 0.5 0.30 27 Gas 297.50 5.404 153.1 0.4 0.62 78 Liquid 298.07 14.504 862.4 0.5 0.30 28 Gas 297.50 5.604 165.3 0.5 0.60 79 Liquid 298.08 14.002 857.8 0.5 0.29 29 Liquid 298.06 40.253 992.4 0.2 0.71 80 Liquid 298.07 13.633 854.0 0.5 0.28 30 Liquid 298.07 39.735 991.0 0.2 0.70 81 Liquid 298.07 13.024 848.3 0.5 0.27 31 Liquid 298.06 39.498 990.0 0.3 0.70 82 Liquid 298.07 12.512 843.3 0.6 0.26 32 Liquid 298.06 39.023 988.6 0.2 0.69 83 Liquid 298.08 11.996 837.8 0.6 0.25 33 Liquid 298.07 38.510 986.9 0.3 0.69 84 Liquid 298.07 11.500 832.0 0.6 0.24 34 Liquid 298.05 37.999 985.2 0.3 0.68 85 Liquid 298.08 11.013 825.9 0.6 0.23 35 Liquid 298.06 37.004 982.1 0.3 0.66 86 Liquid 298.08 10.511 819.3 0.7 0.23 36 Liquid 298.06 36.016 978.6 0.3 0.65 87 Liquid 298.07 9.998 812.8 0.7 0.22 37 Liquid 298.06 35.006 974.9 0.3 0.63 88 Liquid 298.07 9.500 804.5 0.8 0.21 38 Liquid 298.07 34.449 973.0 0.3 0.62 89 Liquid 298.08 9.001 795.9 0.8 0.20 39 Liquid 298.06 33.955 971.2 0.3 0.61 90 Liquid 298.08 8.504 786.3 0.9 0.19 40 Liquid 298.07 33.502 969.5 0.3 0.61 91 Liquid 298.07 8.003 775.2 1.0 0.18 41 Liquid 298.06 33.000 967.6 0.3 0.60 92 Liquid 298.07 7.503 762.2 1.2 0.17 42 Liquid 298.07 32.526 965.7 0.3 0.59 93 Liquid 298.08 7.001 746.0 1.5 0.16 43 Liquid 298.06 32.025 963.8 0.3 0.58 94 Liquid 298.07 6.900 742.3 1.6 0.16 44 Liquid 298.07 31.546 961.8 0.3 0.58 95 Liquid 298.06 6.801 738.3 1.6 0.16 45 Liquid 298.06 31.090 960.2 0.3 0.57 96 Liquid 298.06 6.700 733.8 1.7 0.16 46 Liquid 298.06 30.499 957.6 0.3 0.56 97 Liquid 298.08 6.600 729.1 1.9 0.16 47 Liquid 298.07 29.996 955.4 0.3 0.55 98 Liquid 298.06 6.500 723.9 2.0 0.16 48 Liquid 298.08 29.493 953.2 0.3 0.54 99 Liquid 298.08 6.450 720.9 2.1 0.16 49 Liquid 298.07 29.029 951.1 0.3 0.54 100 Liquid 298.08 6.401 718.3 2.2 0.16 50 Liquid 298.06 28.538 948.9 0.3 0.53 101 Liquid 298.10 6.354 716.1 2.3 0.16 51 Liquid 298.07 28.032 946.6 0.3 0.52

a Standard uncertainties, u, are u(T) = 0.02 K, u(p) = 0.002 MPa for pressures up to 10 MPa and u(p) = 0.005 MPa for pressures from 10-40 MPa

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Table 6 - Experimental results of the 0.9505 mole CO2 + 0.0495 mole H2S system at 323 K a

No Phase T/K p/MPa ρ/kg.m-3 Z No Phase T/K p/MPa ρ/kg.m-3 Z

Exp. uc(ρ) Exp. uc(ρ) 1 SC 322.40 40.495 910.3 0.3 0.72 39 SC 322.43 17.599 735.2 0.6 0.39 2 SC 322.40 39.992 908.1 0.3 0.71 40 SC 322.43 17.106 726.5 0.7 0.38 3 SC 322.42 39.502 906.0 0.3 0.71 41 SC 322.43 16.502 714.5 0.7 0.37 4 SC 322.41 39.014 903.8 0.3 0.70 42 SC 322.44 16.031 705.9 0.8 0.37 5 SC 322.41 38.492 901.2 0.3 0.69 43 SC 322.43 15.744 699.2 0.8 0.37 6 SC 322.42 38.008 899.1 0.3 0.69 44 SC 322.43 15.001 682.0 0.9 0.36 7 SC 322.42 37.495 896.5 0.3 0.68 45 SC 322.42 14.515 671.3 1.0 0.35 8 SC 322.41 37.001 894.2 0.3 0.67 46 SC 322.44 14.019 655.5 1.1 0.35 9 SC 322.42 36.519 891.6 0.3 0.66 47 SC 322.42 13.503 640.1 1.2 0.34 10 SC 322.42 36.000 889.0 0.3 0.66 48 SC 322.43 12.996 620.9 1.3 0.34 11 SC 322.41 35.495 886.3 0.3 0.65 49 SC 322.44 12.023 577.0 1.9 0.34 12 SC 322.42 35.002 883.6 0.3 0.64 50 SC 322.42 11.519 546.6 2.3 0.34 13 SC 322.42 34.503 880.8 0.3 0.64 51 SC 322.39 9.505 341.5 1.2 0.45 14 SC 322.42 34.004 878.0 0.3 0.63 52 SC 322.43 9.001 280.6 0.8 0.52 15 SC 322.42 33.498 875.2 0.3 0.62 53 SC 322.45 8.500 247.2 0.6 0.56 16 SC 322.42 33.004 872.4 0.3 0.61 54 SC 322.46 8.000 217.9 0.5 0.60 17 SC 322.43 32.503 869.6 0.3 0.61 55 Gas 322.46 7.503 192.7 0.4 0.63 18 SC 322.43 32.002 866.5 0.3 0.60 56 Gas 322.47 7.001 170.9 0.3 0.66 19 SC 322.42 31.506 863.7 0.3 0.59 57 Gas 322.47 6.501 151.5 0.3 0.70 20 SC 322.43 31.004 860.5 0.3 0.58 58 Gas 322.47 6.004 133.9 0.2 0.73 21 SC 322.43 30.505 857.5 0.3 0.58 59 Gas 322.47 5.503 118.8 0.2 0.75 22 SC 322.43 30.006 854.1 0.3 0.57 60 Gas 322.48 5.004 103.9 0.2 0.78 23 SC 322.43 29.502 850.6 0.4 0.56 61 Gas 322.48 4.501 90.7 0.2 0.81 24 SC 322.42 29.009 847.7 0.4 0.56 62 Gas 322.48 3.994 77.8 0.1 0.83 25 SC 322.42 27.425 837.1 0.4 0.53 63 Gas 322.48 3.544 66.6 0.1 0.86 26 SC 322.43 27.003 833.6 0.4 0.53 64 Gas 322.47 3.003 55.8 0.1 0.87 27 SC 322.44 26.506 829.2 0.4 0.52 65 Gas 322.48 2.803 51.5 0.1 0.88 28 SC 322.43 26.031 824.9 0.4 0.51 66 Gas 322.48 2.504 45.3 0.1 0.90 29 SC 322.42 25.008 817.6 0.4 0.50 67 Gas 322.48 2.249 40.2 0.1 0.91 30 SC 322.43 24.503 813.0 0.4 0.49 68 Gas 322.49 2.003 35.3 0.1 0.92 31 SC 322.44 24.134 809.6 0.4 0.48 69 Gas 322.49 1.749 30.7 0.1 0.92 32 SC 322.43 20.544 780.3 0.5 0.43 70 Gas 322.48 1.500 25.9 0.1 0.94 33 SC 322.40 20.400 775.4 0.5 0.43 71 Gas 322.49 1.306 22.1 0.1 0.96 34 SC 322.43 20.014 767.7 0.5 0.42 72 Gas 322.48 1.002 16.8 0.1 0.97 35 SC 322.43 19.480 761.0 0.6 0.42 73 Gas 322.49 0.801 13.5 0.1 0.96 36 SC 322.43 19.007 754.0 0.6 0.41 74 Gas 322.49 0.602 10.1 0.1 0.97 37 SC 322.44 18.728 748.8 0.6 0.41 75 Gas 322.49 0.498 8.0 0.1 1.01 38 SC 322.42 18.005 740.8 0.6 0.39 76 Gas 322.48 0.395 6.4 0.1 1.00

a Standard uncertainties, u, are u(T) = 0.02 K, u(p) = 0.002 MPa for pressures up to 10 MPa and u(p) = 0.005 MPa for pressures from 10-40 MPa

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Table 7 - Experimental results of the 0.9505 mole CO2 + 0.0495 mole H2S system at 353 K a

No Phase T/K p/MPa ρ/kg.m-3 Z No Phase T/K p/MPa ρ/kg.m-3 Z

Exp. uc(ρ) Exp. uc(ρ) 1 SC 352.98 40.721 814.9 0.3 0.74 46 SC 352.96 19.005 551.2 0.9 0.51 2 SC 352.99 40.702 814.4 0.3 0.74 47 SC 352.98 18.502 536.7 0.9 0.51 3 SC 352.97 40.515 813.3 0.3 0.74 48 SC 352.98 18.005 522.1 1.0 0.51 4 SC 352.98 40.059 810.8 0.3 0.73 49 SC 352.96 17.499 505.7 1.0 0.51 5 SC 352.98 39.509 807.3 0.3 0.73 50 SC 352.97 17.001 488.9 1.1 0.52 6 SC 352.98 39.054 803.9 0.3 0.72 51 SC 352.97 16.500 471.0 1.1 0.52 7 SC 352.95 38.517 800.4 0.3 0.71 52 SC 352.98 16.003 452.2 1.2 0.52 8 SC 352.96 38.004 797.0 0.3 0.71 53 SC 353.00 15.503 432.9 1.2 0.53 9 SC 352.96 37.470 794.1 0.3 0.70 54 SC 352.99 15.001 412.5 1.2 0.54 10 SC 352.96 37.004 790.2 0.3 0.69 55 SC 352.98 14.500 391.7 1.2 0.55 11 SC 352.96 36.555 786.3 0.3 0.69 56 SC 352.98 14.003 371.6 1.2 0.56 12 SC 352.96 35.917 782.8 0.3 0.68 57 SC 352.98 13.497 350.2 1.2 0.57 13 SC 352.93 35.508 779.2 0.4 0.68 58 SC 352.97 13.007 329.4 1.2 0.59 14 SC 352.96 35.000 774.5 0.4 0.67 59 SC 353.00 12.511 308.8 1.1 0.60 15 SC 352.96 34.514 770.9 0.4 0.66 60 SC 352.99 12.007 288.6 1.1 0.62 16 SC 352.96 33.999 767.0 0.4 0.66 61 SC 352.98 11.500 268.9 1.0 0.63 17 SC 352.94 33.504 763.1 0.4 0.65 62 SC 352.98 10.998 250.1 0.9 0.65 18 SC 352.96 33.000 758.5 0.4 0.65 63 SC 352.99 10.504 232.8 0.9 0.67 19 SC 352.96 32.501 753.9 0.4 0.64 64 SC 352.99 9.987 217.7 0.3 0.68 20 SC 352.96 32.001 749.7 0.4 0.63 65 SC 352.99 9.504 202.4 0.3 0.70 21 SC 352.96 31.506 745.0 0.4 0.63 66 Gas 352.99 9.009 187.0 0.2 0.71 22 SC 352.96 31.043 741.1 0.4 0.62 67 Gas 352.98 8.504 172.4 0.2 0.73 23 SC 352.97 30.503 735.1 0.4 0.62 68 Gas 353.00 8.009 158.7 0.2 0.75 24 SC 352.97 30.006 730.2 0.4 0.61 69 Gas 352.99 7.526 145.7 0.2 0.77 25 SC 352.97 29.502 724.7 0.4 0.60 70 Gas 353.00 7.006 132.8 0.2 0.78 26 SC 352.98 29.007 719.7 0.4 0.60 71 Gas 352.98 6.499 120.5 0.2 0.80 27 SC 352.98 28.505 714.4 0.5 0.59 72 Gas 352.98 6.017 109.3 0.1 0.82 28 SC 352.99 28.006 708.6 0.5 0.59 73 Gas 352.98 5.511 98.1 0.1 0.83 29 SC 352.97 27.507 702.5 0.5 0.58 74 Gas 352.99 5.014 87.3 0.1 0.85 30 SC 352.98 27.002 696.2 0.5 0.58 75 Gas 352.98 4.496 77.2 0.1 0.86 31 SC 352.98 26.504 690.1 0.5 0.57 76 Gas 352.98 4.015 67.8 0.1 0.88 32 SC 352.98 25.999 683.7 0.5 0.56 77 Gas 352.98 3.498 57.6 0.1 0.90 33 SC 352.97 25.500 676.5 0.5 0.56 78 Gas 352.97 3.004 48.9 0.1 0.91 34 SC 352.97 25.005 669.8 0.5 0.55 79 Gas 352.98 2.801 44.8 0.1 0.93 35 SC 352.96 24.499 662.4 0.6 0.55 80 Gas 352.97 2.517 40.0 0.1 0.93 36 SC 352.98 24.005 654.6 0.6 0.54 81 Gas 352.97 2.253 35.5 0.1 0.94 37 SC 352.98 23.503 646.1 0.6 0.54 82 Gas 352.97 2.019 31.5 0.1 0.95 38 SC 352.98 23.002 637.9 0.6 0.53 83 Gas 352.98 1.745 27.0 0.1 0.96 39 SC 352.99 22.505 629.4 0.7 0.53 84 Gas 352.99 1.523 23.3 0.1 0.97 40 SC 352.98 22.006 619.7 0.7 0.53 85 Gas 352.98 1.267 19.0 0.1 0.99 41 SC 352.98 21.508 610.3 0.7 0.52 86 Gas 352.99 0.998 15.1 0.1 0.98 42 SC 352.96 20.999 599.7 0.7 0.52 87 Gas 352.99 0.754 11.4 0.1 0.98 43 SC 352.96 20.506 589.4 0.8 0.52 88 Gas 353.00 0.502 7.6 0.1 0.98 44 SC 352.96 19.998 577.5 0.8 0.51 89 Gas 352.99 0.399 5.9 0.1 1.00 45 SC 352.97 19.503 564.8 0.9 0.51

a Standard uncertainties, u, are u(T) = 0.02 K, u(p) = 0.002 MPa for pressures up to 10 MPa and u(p) = 0.005 MPa for pressures from 10-40 MPa

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Table 8 - Uncertainties with 95% level of confidence (k=2) for the measured densities of the 0.9505 mole CO2 + 0.0495 mole H2S system

Uncertainties U(ρ)

T / K Phase Data Average Max.

No. kg.m-3 % kg.m-3 %

273 Gas 19 0.1 0.6 0.2 1.8

Liquid 75 0.3 0.0 0.4 0.0

SC 0 0 0 0.0 0.0

Total 94 0.3 0.1 0.4 1.8

283 Gas 22 0.2 0.3 0.5 0.9

Liquid 74 0.3 0.0 0.8 0.1

SC 0 0 0 0.0 0.0

Total 96 0.3 0.1 0.8 0.9

298 Gas 28 0.2 0.4 0.5 1.7

Liquid 73 0.6 0.1 2.3 0.3

SC 0 0 0 0.0 0.0

Total 101 0.5 0.2 2.3 1.7

323 Gas 22 0.1 0.4 0.4 1.4

Liquid 0 0 0 0.0 0.0

SC 54 0.6 0.1 2.3 0.4

Total 76 0.4 0.2 2.3 1.4

353 Gas 24 0.1 0.3 0.2 1.4

Liquid 0 0 0 0.0 0.0

SC 65 0.6 0.1 1.2 0.4

Total 89 0.5 0.2 1.2 1.4

Total Gas 115 0.2 0.4 0.5 1.8

Liquid 222 0.4 0.0 2.3 0.3

SC 119 0.6 0.1 2.3 0.4

Total 456 0.4 0.2 2.3 1.8

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Table 9 - Summarised AADs for the measured density of the 0.9505 mole CO2 + 0.0495 mole H2S system

Gas Liquid SC Overall

No of Exp. Data 115 222 119 456

Average PR 1.5 4.2 2.1 3.0

Absolute SRK 1.3 8.0 8.0 6.3

Deviation VPT 1.4 3.0 1.8 2.3

(%) PR-CO2 1.0 1.3 2.3 1.5

SRK-CO2 1.0 1.6 1.3 1.4

VPT-CO2 1.0 1.2 2.4 1.4

PR-Peneloux 1.3 2.3 2.6 2.1

SRK-Peneloux 1.1 3.0 3.2 2.6

GERG-2008 1.0 1.0 1.3 1.1

Maximum PR 4.5 10.6 6.5 10.6

Absolute SRK 7.1 19.4 12.2 19.4

Deviation VPT 4.6 11.6 6.5 11.6

(%) PR-CO2 4.4 4.0 4.9 4.9

SRK-CO2 4.7 3.7 3.8 4.7

VPT-CO2 4.4 4.0 5.4 5.4

PR-Peneloux 4.6 13.0 6.4 13.0

SRK-Peneloux 6.2 15.7 8.6 15.7

GERG-2008 4.3 2.5 5.7 5.7

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Table 10 - Specific heat capacity calculations for the 0.9505 mole CO2 + 0.0495 mole H2S system a

T/K p/MPa ρ/

kg.m-3 ν/

(m3.mol-1) Cp

0/ (J.K-1.mol-1)

CpExp/ (J.K-1.mol-1)

U(Cp)/ (J.K-1.mol-1)

CpGERG/ (J.K-1.mol-1)

Dev./(%)

40 MPa 272.57 40.046 1054.95 4.13E-05 35.78 73.27 0.10 78.10 6.6 283.47 40.004 1040.33 4.18E-05 36.35 73.34 0.11 78.33 6.8

298.06 40.253 992.41 4.39E-05 37.09 73.42 0.11 78.71 7.2

322.40 39.992 908.08 4.79E-05 38.25 74.01 0.12 79.79 7.8

352.98 40.059 810.79 5.37E-05 39.58 74.00 0.13 80.64 9.0

35 MPa 272.56 35.060 1041.12 4.18E-05 35.78 74.86 0.13 79.31 5.9 283.46 35.004 1025.18 4.24E-05 36.35 75.19 0.13 79.80 6.1

298.06 35.006 974.89 4.46E-05 37.09 75.78 0.14 80.65 6.4

322.42 35.002 883.60 4.93E-05 38.25 77.01 0.15 82.43 7.0

352.96 35.000 774.52 5.62E-05 39.58 78.03 0.17 84.17 7.9

30 MPa 272.57 30.022 1025.61 4.24E-05 35.78 77.03 0.16 80.82 4.9 283.47 30.014 1008.20 4.32E-05 36.35 77.64 0.17 81.65 5.2

298.07 29.996 955.45 4.55E-05 37.09 78.87 0.18 83.11 5.4

322.43 30.006 854.09 5.10E-05 38.25 81.42 0.19 86.16 5.8

352.97 30.006 730.17 5.96E-05 39.58 84.11 0.21 89.33 6.2

25 MPa 272.58 24.975 1008.49 4.32E-05 35.78 79.89 0.21 82.76 3.6 283.47 25.000 988.85 4.40E-05 36.35 80.97 0.22 84.11 3.9

298.07 25.023 932.12 4.67E-05 37.09 83.34 0.24 86.50 3.8

322.42 25.008 817.64 5.32E-05 38.25 88.23 0.27 91.89 4.2

352.97 25.005 669.79 6.50E-05 39.58 93.93 0.28 97.82 4.1

20 MPa 272.58 20.011 989.08 4.40E-05 35.78 82.07 0.34 85.28 3.9 283.47 20.010 966.84 4.50E-05 36.35 83.86 0.36 87.51 4.3

298.06 19.999 903.89 4.81E-05 37.09 87.85 0.39 91.67 4.3

322.43 20.014 767.74 5.67E-05 38.25 98.41 0.45 102.20 3.8

352.96 19.998 577.46 7.54E-05 39.58 108.87 0.46 113.28 4.1

15 MPa 272.58 15.013 967.11 4.50E-05 35.78 88.80 0.82 88.83 0.0 283.47 15.004 940.37 4.63E-05 36.35 92.27 0.87 92.67 0.4

298.07 15.005 867.00 5.02E-05 37.09 100.54 0.95 100.66 0.1

322.43 15.001 682.02 6.38E-05 38.25 131.03 1.18 129.00 -1.6

352.99 15.001 412.49 0.0001055 39.58 126.51 1.24 129.50 2.4

AAD (%) 4.8 a Standard uncertainties, u, are u(T) = 0.02 K and u(p) = 0.005 MPa for pressures from 10-40 MPa

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Table 11 - Estimated dew and bubble points from the measured density data for the 0.9505 mole CO2 + 0.04905 mole H2S system a

Type Experimental Prediction (PR-CO2) Deviations (%)

T/K p/MPa ρ/ kg.m-3 uc(ρ)/ kg.m-3 p/MPa ρ/ kg.m-3 p ρ

Dew Point 272.56 3.435 97.4 0.3 3.379 93.3 -1.6 -4.2

Dew Point 282.70 4.334 127.4 0.4 4.387 125.6 1.2 -1.4

Dew Point 297.50 6.166 225.0 0.6 6.232 236.0 1.1 4.9

Bubble Point 273.32 4.193 889.9 0.5 3.445 905.5 -17.8 1.8

Bubble Point 283.90 4.621 861.4 0.7 4.537 850.5 -1.8 -1.3

Bubble Point 298.07 6.320 714.3 2.3 6.313 731.9 -0.1 2.5

Absolute Average Deviation (AAD) (%) 3.8 2.7 a Standard uncertainties, u, are u(T) = 0.02 K, u(p) = 0.002 MPa for pressures up to 10 MPa

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Figure 1 - Schematic diagram of the densitometer apparatus

1) Anton Paar DMA 512 densitometer, 2) High pressure fluid vessel, 3) Capillary Valve

(ROLSITM) [55], 4) Lauda liquid bath, 5) Liquid bath of circuit, 6) Pressure transducers, 7)

Temperature probes, 8) Vacuum pump, 9) Neutralisation column, 10) Data acquisition unit,

V1) Circuit valve, V2) Flow controlling ball valve, V3)Vacuum valve, V4) Venting valve,

V5) Neutralisation valve

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Figure 2 - Pressure – composition diagram of the CO2 (1) + H2S (2) system. a) T =258.41 K. (), this work. b) T = 273.15 K. (), this work; (), T=273.15 K from [28]. c) T = 293.47 K. (), this work; () T = 293.15 K from [28]. d) T = 313.02 K. (), this work (), T = 313.15 K, from [28]. Black lines: calculated bubble and dew lines using the SRK model with kij = 0.098. Dotted lines: calculated bubble and dew lines using the adjusted PR-EoS model. Grey lines: calculated bubble and dew lines using the SRK-EoS model with kij =0. Broken lines (— – —): GERG-2008 EoS.

0

0.5

1

1.5

2

2.5

0 0.2 0.4 0.6 0.8 1

P /

MP

a

x1,y1

a)

0

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4

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P /

MP

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x1,y1

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9

0 0.2 0.4 0.6 0.8 1

P /

MP

a

x1,y1

d)

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Figure 3 - Experimental and predicted densities of the 0.9505 mole CO2 + 0.0495 mole H2S system. Experimental results: (∆) 273. 15 K, (*) 283.15 K, () 298.15 K, (◊) 323.15 K and () 353.15 K.

Lines: Predictions using the GERG-2008 EoS

0

5

10

15

20

25

30

35

40

45

0 200 400 600 800 1000 1200

P/

MP

a

Density / kg.m-3

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Figure 4 – Low pressure experimental and predicted densities of the 0.9505 mole CO2 + 0.0495 mole H2S system. Experimental results: (∆) 273. 15 K, (*) 283.15 K, () 298.15 K, (◊) 323.15 K and ()

353.15 K. Lines: Predictions using the GERG-2008 EoS

0

1

2

3

4

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9

10

0 40 80 120 160 200

P/

MP

a

Density / kg.m-3

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Figure 5 – Comparison of experimental and predicted densities results of the CO2 + H2S system at two

isotherms with the literature data, Experimental results in this work: (◊) 322.4 K and () 353.15 K. Stouffer et al. [7]: (♦) 325 K and ZH2S = 0.0955, ()325 K and ZH2S = 0.0607, ()350 K and ZH2S = 0.0607, ()350 K and ZH2S = 0.0955, Lines: Predictions using the GERG-2008 EoS, Dash lines:

Predictions using the PR-CO2 EoS

0

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45

0 200 400 600 800 1000

P/

MP

a

Density / kg.m-3

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Figure 6 – Comparison of experimental and predicted densities of the CO2 + H2S system at two

isotherms with the literature data at lower pressures, Experimental results in this work: (◊) 322.4 K and () 353.15 K. Stouffer et al. [7]: (♦) 325 K and ZH2S = 0.0955, ()325 K and ZH2S = 0.0607,

()350 K and ZH2S = 0.0607, ()350 K and ZH2S = 0.0955, Lines: Predictions using the GERG-2008 EoS, Dash lines: Predictions using the PR-CO2 EoS

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1

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10

0 50 100 150 200

P/

MP

a

Density / kg.m-3

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Figure 7 – Deviations of GERG-2008 EoS from experimental densities of CO2 - H2S

This work: (◊) 322.4 K and ZH2S = 0.0495, Stouffer et al. [7]: (♦) 325 K and ZH2S = 0.0955, () 325 K and ZH2S = 0.0607

-5.0

-4.0

-3.0

-2.0

-1.0

0.0

1.0

2.0

3.0

4.0

5.0

0 200 400 600 800 1000

(ρG

ER

G-ρ

Exp

)/ρ

Exp

×1

00

Density / kg.m-3

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Figure 8 – Deviations of the predictions using PR-CO2 equation of state from the measured data

-5.0

-4.0

-3.0

-2.0

-1.0

0.0

1.0

2.0

3.0

4.0

5.0

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(ρP

R-C

O2-ρ

Exp

)/ρ

Exp

×1

00

P / MPa

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Figure 9 – Compressibility factor of the 0.9505 mole CO2 + 0.0495 mole H2S system. Experimental results: (∆) 273. 15 K, (*) 283.15 K, () 298.15 K, (◊) 323.15 K and () 353.15 K.

Lines: Predictions using the GERG-2008 EoS

0.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0

0 5 10 15 20 25 30 35 40 45

Z

P / MPa

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Figure 10 – Compressibility factor of the 0.9505 mole CO2 + 0.0495 mole H2S system.

Experimental results: (∆) 273. 15 K, (*) 283.15 K, () 298.15 K, (◊) 323.15 K and () 353.15 K. Lines: Predictions using the GERG-2008 EoS

0.50

0.55

0.60

0.65

0.70

0.75

0.80

0.85

0.90

0.95

1.00

0 2 4 6 8 10

Z

P / MPa

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Figure 11 - Specific heat capacity of the 0.9505 mole CO2 + 0.0495 mole H2S system.

() calculated Cp with 2% error, Lines: predictions using the GERG-2008 EoS

50

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65

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75

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95

100

270 290 310 330 350 370

Cp

/ (J

.K-1

.mo

l-1)

T / K

40 MPa

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270 290 310 330 350 370

Cp

/ (J

.K-1

.mo

l-1)

T / K

35 MPa

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270 290 310 330 350 370

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.K-1

.mo

l-1)

T / K

30 MPa

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270 290 310 330 350 370

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T / K

25 MPa

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20 MPa

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Cp

/ (J

.K-1

.mo

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T / K

15 MPa

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Figure 12 – Dew point determination of the 0.9505 mole CO2 + 0.0495 mole H2S system for different

isotherms, (∆) 273. 15 K, (*) 283.15 K, () 298.15 K

0.0

1.0

2.0

3.0

4.0

5.0

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7.0

0 50 100 150 200 250 300

P /

MP

a

Density / kg.m-3

Dew points

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Figure 13 – Bubble point determination of the 0.9505 mole CO2 + 0.0495 mole H2S system for different isotherms, (∆) 273. 15 K, (*) 283.15 K, () 298.15 K

0

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P/M

Pa

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Bubble points

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Appendix A:

Density Measurement uncertainties

The combined standard uncertainties [38][39] of density measurements for each measured

quantity have been calculated using the root sum of the squares of uncertainties as shown in

the following equation.

$% = &$'( + $(*( + $+( (A1)

where $' is the estimated uncertainties due to the temperature, $(* the estimated

uncertainties due to pressure and $+ the estimated uncertainties of oscillation period.

The estimated uncertainties due to temperature variations, $', has been calculated from

the equation below:

$' = ,-./.01( . $'( (A2)

In the above equation, u(T) is the standard estimated uncertainty of temperature probe and is

considered to be U(T) = 0.04 K. The density gradient due to the temperature variations, -./.01, has been calculated from the equation below.

-./.01 = (20 0320 − 020 (A3)

The upper and lower limits of densities due to temperature effect, 0320456020, were

estimated from REFPROP v8.0 [56].

The similar procedures have been followed to estimate the uncertainties due to the pressure

and period of oscillations. The standard uncertainty of pressure transducer, u(p)=0.005 MPa

and standard uncertainty of oscillation period, u(τ)=0.005 µs [57].

$(* = ,-./.1( . $*( (A4)

-./.1 = (2 032 − 02 (A5)

$+ = ,-./.71( . $( (A6)

-./.71 = (27 0327 − 027 (A7)

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Isobaric heat capacity uncertainties

The combined standard uncertainty [38][39] of isobaric heat capacities were calculated from

the equation below.

$% = &$'( + $*( + $4( (A8)

The uncertainties due to the temperature and pressure, u(T) and u(p), can be calculated

through Equation A9 to Equation A12.

$' = ,-.89.0 1( . $'( (A9)

-.89.0 1 = (20 -0320 − 0201 (A10)

$* = ,-.89. 1( . $*( (A11)

-.89. 1 = (2 -32 − 21 (A12)

The uncertainties due to the parameter a in Equation (14) is determined by changing in the

densities. The accuracy in the density measurement is estimated to be ±1.0 kg/m3 and then

the upper and lower limits of parameter a were determined from the polynomial Equation

(12). Then, the upper and lower Cp were calculated through Equation (14). The uncertainties

due to parameter a then can be calculated through equations below.

$4 = ,-.89.: 1( . $4( (A13)

-.89.: 1 = (2: -:32: − :2:1 (A14)

Expanded uncertainties are defined as ;<, > = 2 = 2 × $@ in this work.

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Highlights:

• Pressure-volume-temperature properties of a 0.95CO2+0.05H2S were investigated.

• Densities of the above mixture have been measured at high pressures.

• To validate, the experimental results have been compared to literature data.

• Thermodynamic properties have been derived from the experimental data.

• Different EoSs have been evaluated using the new experimental data.