J Kittel

F Grosjean

C Forot

C Duret-Thual

P Fourny

Impact of time of exposure on HIC testing of very high strength steel in low H 2 S environments

The methodology for testing hydrogen induced cracking (HIC) resistance of carbon steel in mild sour service is discussed. When H 2 S content is low, the use of steel grades with intermediate HIC resistance often presents many advantages, such as higher mechanical properties and economical benefits. As a counterpart, fit-forpurpose testing procedures are required in order to assess the HIC performance at the intended service conditions. With regards low H 2 S environments, time of exposure is recognised as the critical testing parameter. The environment encountered in the annular space of flexible pipelines is very often mildly severe. Due to the high confinement, pH level is higher than in bulk conditions with the same gas composition but a larger V/S ratio between the volume of electrolyte and the exposed steel surface. H 2 S content in the annulus is also less than in the fluid transported in the bore of the pipeline, due to permeation through inner thermoplastic sheaths and consumption. Qualification tests of very high strength steel wires are therefore usually conducted with one month exposure at the predicted pH and P H2S in the annulus, while standard testing procedures referenced in ISO 15156 specify 96 h immersion but a higher H 2 S level.

In the present study, different grades of high strength steel wires were tested in environments containing as low as 100 ppmv H 2 S, with exposure times varying from 96 hours and up to 1 year. For the steels grades that have been studied, it is confirmed that the actual qualification procedure consisting in 1 month exposure is appropriate to detect HIC susceptibility.

Introduction

Development of deep and ultra-deep water fields often requires the use of high strength and very high strength steel wires (UTS≥1000MPa) in flexible pipes in order to sustain mechanical loads [START_REF] Taravel-Condat | Qualification of very high strength carbon steel wires for use in flexible pipes containing small amount of H 2 S, Eurocorr[END_REF][START_REF] Désamais | Use of High Strength Steel Wires for Flexible Pipe in Low Sour Service Conditions: Impact on Deep Water Applications[END_REF]. On the other hand, these fields are most of the time considered as containing H 2 S at design stage due to the possible souring of the reservoir. Hopefully, in flexible pipe, steel wires are not in direct contact with the conveyed fluid. They are located in the annular space (see Figure 1) which is a specific medium for corrosion, less severe than the bore. These high strength steel wires need to be qualified with regards to H 2 S (Sulfide Stress Cracking and Hydrogen Induced Cracking) [START_REF] Désamais | Use of High Strength Steel Wires for Flexible Pipe in Low Sour Service Conditions: Impact on Deep Water Applications[END_REF]. Nevertheless, testing procedures referenced in ISO15156 are not appropriate for these high strength wires that have a low H 2 S resistance. Therefore for SSC and HIC qualification, environment closer than the one encountered in flexible pipe annulus are tested:

 H 2 S and CO 2 partial pressures in the annulus are calculated thanks to a dedicated permeability model [START_REF] Benjelloun-Dabagui | MOLDI TM : A fluid permeability model to calculate the annulus composition in flexible pipes[END_REF] that considers permeation through the polymer pressure and external sheathes. Furthermore, as presented in [START_REF] Désamais | On The Beneficial Influence of a Very Low Supply of H 2 S on the Hydrogen Embrittlement Resistance of Carbon Steel Wires in Flexible Pipe Annulus[END_REF], the flow rate of H 2 S arriving in the annulus of flexible pipe by permeation through the polymer pressure sheath is very low. The consequence is that when the annulus is flooded, the thermodynamic equilibrium calculated thanks to the permeability model is not reached. Indeed, due to the high confinement, H 2 S arriving in the annulus is partly consumed by iron sulfide precipitation. As a consequence H 2 S levels in flexible pipe annulus is even lower than presented previously.  pH is determined from long term experiments in confined environment, where the ratio between the volume of electrolyte and the surface of steel is extremely low, typically between 0.02 and 0.06 mL/cm² [START_REF] Ropital | Eurocorr[END_REF][START_REF] Rubin | Corrosion rates of carbon steel in confined environments[END_REF][START_REF] Remita | A kinetic model for CO2 corrosion in confined aqueous environments[END_REF][START_REF] Remita | Experimental and theoretical investigation of the uniform corrosion in the annulus of offshore flexible pipelines[END_REF]. Representative environments simulating service conditions are thus characterized by relatively high pH, between 5 and 6, and P H2S often below a few mbar. Fit-forpurpose tests are therefore systematically carried-out to check H 2 S suitability of steel grade to target annulus environment. And since H 2 S severity is extremely mild, experimental procedures have to be adapted. The goal of this paper is to discuss methodology for HIC testing of carbon steel wire in mild sour environments. In particular, great attention will be paid to time of exposure. It is well known that high pH and low H 2 S induce slow ingress of hydrogen in the steel [START_REF] Kittel | Effect of Membrane thickness on hydrogen permeation in steels during wet hydrogen sulphide exposure[END_REF][START_REF] Kittel | Hydrogen induced cracking (HIC) testing of low alloy steel in sour environment: Impact of time of exposure on the extent of damage[END_REF], thus long incubation times before HIC starts. Recent work on linepipe steels have shown that exposure periods of a few weeks were necessary at P H2S below 10 mbar [START_REF] Kittel | HIC laboratory testing assessment of low alloy steel linepipe[END_REF]. In this paper, exposure times from a few days and up to one year were performed, to evaluate the evolution of cracking with time, and to determine a reasonable and safe testing exposure period for qualification tests.

HIC experiments in low H S environments 2.1 Experimental methods

Tested materials

Wires of two carbon steel grades were tested, the first (A) used for sweet service and the second one (B) for mildly sour environments. Most specimens were 120 mm long, 12 mm large and 5 mm thick. The tested surface condition was either as received (raw surface) or polished. In the as received condition, a thin layer of mixed oxide and phosphate is present due to processing conditions. The applied polishing procedure was in agreement with the HIC NACE standard TM0284 (P320 SiC grit paper). All the cutting edges of the specimens were also polished (P600 SiC grit paper). The specimens were degreased with acetone just before starting the test, otherwise they were oiled and stored in a dessicator.

Experimental procedure

The test matrix included two pH (5.5 and 6), four H 2 S levels in a CO 2 -H 2 S mixture (100, 300, 1000 and 3000 ppm H 2 S i.e. partial pressures of 0.1, 0.3, 1 and 3 mbar).

The tests were run from 4 days to 1 year. In each condition 10 specimens of each grade were introduced, half as-received and half polished. Several glass cells were used simultaneously to conduct the tests. The solution composition was 35 g/l NaCl, with NaHCO 3 addition to attain the target value of pH. Solutions were prepared with demineralized water and analytical grade chemicals. Gas pre-mix of the required compositions were used to saturate the solution with CO 2 -H 2 S mixture. The V/S ratio was 20 ml/cm 2 .

After placing the specimens in the cell, the cell and the solution were de aerated separately, before introducing the solution. The gas mixture was then bubbled at a high rate for saturation and then at a lower rate to maintain the H 2 S level throughout the exposure. A measurement of the sulfide concentration in the solution was performed after saturation, with a specific electrode, in order to ensure that the H 2 S saturation was met. The pH was measured either continuously with a pH electrode or checked regularly. If needed, it was adjusted by addition of acidic or basic solutions.

A view of a test cell with the specimens is given in Figure 2a. An example of pH control and measurement during a one year test is provided in Figure 2b.

HIC Evaluation by ultrasonic testing

Ultrasonic inspection was applied using an immersion transducer and a 2 axis scanner. The system was calibrated according to an in-house technique. The total area of each specimen was scanned to detect the presence of internal defects. To eliminate some side-effects, a distance of 10 mm at each end of the sample was not considered for this study.

Results

Effect of surface condition

The impact of surface preparation is illustrated on Figure 3 for the sweet steel grade (A) tested at pH 6 and 3 mbar H 2 S. This impact vanishes at long immersion times.

For short tests below 1 month, cracking appears faster on polished specimens than on as received specimens. However, after 1 month exposure even if there is still more samples with HIC on polished samples, by testing enough samples, it is sufficient to detect the onset of cracking in both cases. Similar results were obtained for the mild sour grade (B).

Effect of the duration of test

Figure 3 illustrates the evolution of HIC extent of sweet steel (A) exposed at pH 6 and 3 mbar H 2 S, for exposure periods from 96 hours to 3 months. The results show that the usual 96 hours exposure test is clearly too short to determine the susceptibility of the steel to HIC at such low H 2 S levels. In this environment, the indication obtained after one month are clear, whatever the surface condition.

The effect of even lower H 2 S content is presented on Figure 4, with exposures up to 1 year. It appears that the H 2 S limit for this sweet grade lies between 0.1 mbar and 0.3 mbar at pH 6.

It is particularly interesting to note that no significant evolution of HIC extent is observed between 3 months and 1 year exposure, suggesting that an equilibrium state was reached. These results also highlight some HIC resistance of this sweet steel grade at 0.1 mbar H 2 S and pH 6. This is in accordance with previous results obtained after one month exposure. Similar tests in more severe conditions were performed with a mild sour steel grade (B), known to offer some resistance to HIC. Tests were performed at pH 5.5, where the limit of use of this steel grade is above 1 mbar of H 2 S, as determined by extensive qualification tests based on 1 month exposure time. Longer tests were therefore conducted in order to check that the limits of use of this steel were not falsely influenced by too short exposure times. The results of tests at pH 5.5 and 1 mbar H 2 S are presented on Figure 5, and clearly show that no HIC was detected even after 1 year immersion. Similarly as Figure 4 showed that 0.1 mbar H 2 S at pH 6 is included into the safe domain for the sweet grade, we can conclude that 1 mbar H 2 S at pH 5.5 is into the safe domain of the mild sour grade. It is also confirmed that excessively long term exposures are not necessary, since all experiments show the same type of HIC response (positive or negative) after 1 month and after 1 year exposure. It can thus be concluded from these tests that 1 month exposure time of steel armor wires offers good confidence for qualification.

Conclusions and perspectives

A large campaign of HIC tests has been launched on high strength and very high strength steels wires in very low severe conditions representative of flexible pipe annulus environment, i.e. H 2 S partial pressure below a few mbar and pH above 5.

These tests have shown that there is no major impact of surface preparation of tested specimens between polished and as received surface for 1 month duration or more.

HIC tests performed at different durations between 96hours and 1 month with low H 2 S level show that HIC damage increase with time and that a short time test is misleading as HIC has not yet appeared at 96 hours.

In order to demonstrate that 1 month testing is enough in such mild sour environment, longer duration have been tested up to 1 year. For the two steels in the different conditions tested, one month duration has been demonstrated to be sufficient. Indeed no major evolution of HIC damage has been seen between 1 month and 1 year.

One should notice that such results are probably material and environment dependent and therefore not applicable as such for other microstructures or environments. Further studies are ongoing to better understand other environment conditions and materials.

Figure 1 :

 1 Figure 1: Example of a flexible pipeline structure

Figure 2 a

 2 Figure 2 a-view of a test cell with the specimens; b-pH evolution during 1 year for a pH 5.5 test

Figure 3 :

 3 Figure 3: Evolution of HIC extent on sweet steel (A) specimens exposed at ambient temperature and pressure in a 35 g/L NaCl solution saturated with 997 mbar CO 2 + 3 mbar H 2 S and at pH 6.

Figure 4 :

 4 Figure 4: Evolution of HIC extent on sweet steel (A) specimens exposed at ambient temperature and pressure in a 35 g/L NaCl solution saturated with 999.7 mbar CO 2 + 0.3 mbar H 2 S and with 999.9 mbar CO 2 + 0.1 mbar H 2 S .

Figure 5 :

 5 Figure 5: HIC results on a mild sour high strength carbon steel grade tested at pH 5.5 under 1 mbar H 2 S for 3 months to 1 year


C Plennevaux

N Ferrando

J Kittel

M Fregonese

B Normand

T Cassagne

F Ropital

M Bonis

M Frégonèse

pH prediction in concentrated aqueous solutions under high pressure of acid gases and high temperature

published or not. The documents may come

pH prediction in concentrated aqueous solutions under high pressure of acid gases and high temperature

Introduction 1

Produced waters contain dissolved acid gases such as carbon dioxide (CO 2 ) and sometimes hydrogen sulfide (H 2 S). These dissolved gases induce uniform corrosion of low alloy steels. Risks of hydrogen embrittlement and cracking are also specifically associated with wet H 2 S. The risk level is strongly related to in situ pH [START_REF] Nesic | Key issues related to modelling of internal corrosion of oil and gas pipelines -A review[END_REF][START_REF] Malik | The influence of pH and chloride concentration on the corrosion behaviour of AISI 316L steel in aqueous solutions[END_REF][START_REF] Mancia | The effect of environmental modification on the sulphide stress corrosion cracking resistance of 13Cr martensitic stainless steel in H 2 S-CO 2 -Clsystems[END_REF][START_REF] Truman | The influence of chloride content, pH and temperature of test solution on the occurrence of stress corrosion cracking with austenitic stainless steel[END_REF][START_REF] Sury | Activation pH and susceptibility to hydrogen assisted stress corrosion cracking of martensitic stainless steels in sulphate solutions[END_REF]. To make the most accurate material selection in regard to the application, industrial companies and research institutes have developed prediction models of CO 2 and H 2 S corrosion. They provide pH evaluation from water composition and in-situ temperature and partial pressures of acid gases. However usual models are limited to low pressures, relatively low temperatures and to slightly concentrated brines solutions. This is the case of the company model of some of the authors, which is limited to 20 bar of acid gas and 150°C. These limitations were not a problem for typical oil and gas fields operated from the eighties to 2000. However, the increasing number of high pressure (HP) and high temperature (HT) oil and gas fields being developed requires these pH calculation tools to be extended to higher limits of use. The goal of this work is therefore to extend the existing pH evaluation model to HPHT applications and to highly concentrated brines.

Background 2

Before presenting the new model in detail, a brief description of actual methods of pH calculation in oil and gas environments is proposed. Two groups of reactions are usually considered: i/ dissolution of acid gases in the solution and ii/ chemical dissociation of weak acid components. The corresponding reactions are described below.

At the liquid/gas interface, a quantity of CO 2 and H 2 S dissolves in the aqueous solution according to reactions (a) and (b) :

2( ) 2( ) gd CO CO  S CO2 (a) 2 ( ) 2

( ) gd H S H S 

S H2S

(b)

In the liquid phase, dissolved carbon dioxide hydrates to carbonic acid following reaction (c).

HCO H CO

    K 2 (e) 2 H S H HS   K' 1 (f) 2 HS H S     K' 2 (g)
The dissolution of acid gases is described by solubility constants S CO2 and S H2S and equilibrium constants K hyd , K, K' 1 , K 2 and K' 2 define chemical equilibria in the liqui phase as the ratios between the products and the reactants. Theoretically, solubility constants and equilibrium constants depend only on temperature as long as fugacities are used. However for simplicity these constants are frequently expressed vs concentrations and partial pressures. Consequently these constants must include the effects of pressure and ionic strength. Empirical expressions versus temperature and / or ionic strength are often used to describe these equilibrium constants, the effect of pressure being frequently not included. Some of them can be found in references [START_REF]CO 2 corrosion rate calculation model[END_REF][START_REF] Oddo | Simplified calculation of CaCO 3 saturation at high-temperatures and pressures in brine solutions[END_REF][START_REF] Palmer | The chemistry of metal carbonato and carbon dioxide complexes[END_REF][START_REF] Sun | Kinetics of iron carbonate and iron sulfide scale formation in CO 2 /H 2 S corrosion[END_REF].

The main advantage of such methodology is that it allows using the simple formalism of partial pressures and concentrations in a range of temperature, pressure and ionic strength. However, the validity domain remains limited to low pressures, low temperatures and to slightly concentrated brines solutions. As reported in the literature [START_REF] Nyborg | CO 2 corrosion models for oil and gas production systems[END_REF] most of industrial models, including the model of some of the authors' company, are limited to 150°C, 20 bar of CO 2 and an ionic strength of 1 mol.L -1 . In order to extend its range of application, the calculation method of equilibrium constants must be reconsidered taking into account the non-ideal behaviors of gas and liquid phases. This requires more complete thermodynamics, as described in the next section.

Evolution of the model for HPHT conditions 3

At high pressure, the gas phase can not be considered as ideal and fugacity correction needs to be applied to account for interactions between gas molecules. As a consequence, three corrections must be applied in the solubility calculations leading to the ensemble Henry's law [START_REF] Carroll | What is Henry's law?[END_REF][START_REF] Nelson | Selecting representative laboratory test conditions for fit-forpurpose OCTG material evaluations[END_REF] (eq.1):

  ( ) ( )
. . ( ).exp / .

xx i i i i i i x H T RT dP P       (eq. 1)
In this expression, x i is the molar fraction of the component in the liquid phase, H (x) i (T) is the Henry's constant characterizing the solubility of the gas in water on the molar fraction scale and P i is the partial pressure of component i. The activity coefficient defined on the molar fraction scale γ i (x) accounts for non-ideality of dissolved gas, the exponential term is known as the Poynting correction considering the effect of high pressures on the molar volume under infinite dilution and the fugacity coefficient Φ i corrects the gas phase for non ideal behavior.

Similarly, reactions in concentrated solutions are influenced by interactions between dissolved molecules and ions. To consider the non-ideal behavior of the liquid phase, the calculation of chemical equilibrium constants must also be corrected adding activity coefficients (eq.2).

    () ( ) . i i m i i i ii K T a m      (eq. 2)
In this expression, K(T) is the equilibrium constant of one of the chemical reactions (c) to (g), a i , m i and γ i (m) are respectively the activity, the molality and the activity coefficient of component i (on the molality scale) in the liquid phase and υ i is the stoechiometric coefficient of component i for the chemical equilibrium considered.

For the specific application described in this paper, the Henry's constants H i (T) are calculated using Dhima's correlations [START_REF] De Hemptinne | The Henry constant for 20 hydrocarbons, CO 2 and H 2 S in water as a function of pressure and temperature[END_REF], the correlations allowing to describe equilibrium constants K(T) on a wide range of temperature are selected in the literature [START_REF] Austgen | Model of vapor-liquid equilibria for aqueous acid gas-alkanolamine systems .2. Representation of H 2 S and CO 2 solubility in aqueous MDEA and CO 2 solubility in aqueous mixtures of MDEA with MEA or DEA[END_REF][START_REF] Millero | The thermodynamics and kinetics of the hydrogen-sulfide system in naturalwaters[END_REF][START_REF] Olofsson | Thermodynamics of ionization of water over wide ranges of temperature and pressure[END_REF][START_REF] Patterson | Carbonate equilibria in hydrothermal systems: first ionization of carbonic acid in NaCl media to 300°C[END_REF][START_REF] Peiper | Thermodynamics of aqueous carbonate solutions including mixtures of sodium-carbonate, bicarbonate, and chloride[END_REF], fugacity coefficients in the gas phase are calculated using the Soreide and Whitson's model [START_REF] Soreide | Peng-Robinson predictions for hydrocarbons, CO 2 , N 2 , and H 2 S with pure water and NaCI brine[END_REF] and activity coefficients in the liquid phase are calculated applying Pitzer's model [START_REF] Pitzer | Thermodynamics of electrolytes. I. Theoretical basis and general equations[END_REF]. Each of these steps is presented in the next sections.

Gaseous components fugacity coefficients

The Soreide and Whitson [START_REF] Soreide | Peng-Robinson predictions for hydrocarbons, CO 2 , N 2 , and H 2 S with pure water and NaCI brine[END_REF] modification of the Peng-Robinson equation of state (EOS) [START_REF] Peng | A new two-constant equation of state[END_REF] is used for describing the fugacity of CO 2 , H 2 S, H 2 O and CH 4 in the gas phase.

The Peng-Robinson EOS (eq.3) is derived from Van der Waals theory:

    () RT a T P v b v v b b v b      (eq. 3)
In this expression, the pressure is calculated as a function of a repulsive term taking into account the co-volume b and a parameter a regarded as a measure of intermolecular attraction forces. This parameter depends on temperature. R is the gas constant, T is the temperature (K), P the total pressure (Pa) and v the molar volume (m 3 /mol).

Peng and Robinson used this equation to express the fugacity coefficient Φ i of a component i in a mixture [START_REF] Peng | A new two-constant equation of state[END_REF].

      2 2.414 ln 1 ln ln 0.414 22 j ji j i i i xa b b A Z B Z Z B b a b Z B B                           (eq. 4)
where

2 2 T R aP A  , RT bP B  and RT Pv Z  In this expression, x i is the molar fraction of component i, b i is the co-volume of component i, a
ji characterizes the binary system formed by components j and i.

The presence of salts in the formation water also has a great influence on gas solubilities, which is not directly taken into account by the Peng and Robinson expression. Thus, Soreide and Whitson have proposed an extension of the Peng-Robinson EOS for treating the case of water-hydrocarbon mixtures in the presence of sodium chloride in the water phase. One of the main change consists in a modification of the attractive term of pure water.

Equations 4 was used in this study for the calculation of gas fugacities, using constant parameters values from reference [START_REF] De Hemptinne | The Henry constant for 20 hydrocarbons, CO 2 and H 2 S in water as a function of pressure and temperature[END_REF].

Activity coefficients in the liquid phase

The calculation of activity coefficients of components in the liquid phase was performed applying Pitzer's model [START_REF] Pitzer | Thermodynamic properties of aqueous sodium-chloride solutions[END_REF] which derives from Debye-Hückel's method. This activity model is adapted to describe thermodynamic properties of concentrated brines solutions. Debye-Hückel's equations give an expression of activity coefficients of ionic species from the ionic strength of the solution. This method takes into account long distance interactions between ionic species corresponding to electrostatic interactions. This approach is valid for slightly concentrated solutions (maximum molality of salts of about 1mol / kg of water).

To adapt the activity model to highly concentrated brines, Pitzer has added short distance interaction terms to the Debye-Hückel's model corresponding to interactions between ionic species and the solvent. In the Pitzer's formalism, β (0) , β (1) and β (2) represent binary interaction parameters between ionic species and molecules and between cations and anions. Interactions between two cations or two anions θ were neglected. Ternary interactions were also taken into account. C Φ parameter refers to the interaction between two identical molecular or ionic species and a third different one, while ψ refers to the interaction between three different components. Table 1 summarizes all binary and ternary parameters that we considered in our system, considering dissolved CO 2 and H 2 S and the following ionic species : Ca 2+ , Na + , Cl -, H + , HCO 3 -CO 3 2-, HS -, S 2-and HO -.

Table 1: References of Pitzer parameters used in this work

Pitzer parameters β (0) , β (1) , β (2) , C Φ and ψ Reference Validity domain H + / Cl - [START_REF] Holmes | The enthalpy of dilution of HCl(aq) to 648 K and 40 MPa thermodynamic properties[END_REF] 273 K < T < 523 K, P < 400 bar Ca 2+ / Cl - [START_REF] Greenberg | The prediction of mineral solubilities in natural waters: A chemical equilibrium model for the Na-K-Ca-Cl-SO 4 -H 2 O system to high concentration from 0 to 250°C[END_REF] 298 K < T < 523 K Ca 2+ / HCO 3 - [START_REF] Duan | Coupled phase and aqueous species equilibrium of the H 2 O-CO 2 -NaCl-CaCO 3 system from 0 to 250 degrees C, 1 to 1000 bar with NaCl concentrations up to saturation of halite[END_REF] 273 < T < 523 K, P <1000 bar Ca 2+ / CO 3 2- [START_REF] Duan | Coupled phase and aqueous species equilibrium of the H 2 O-CO 2 -NaCl-CaCO 3 system from 0 to 250 degrees C, 1 to 1000 bar with NaCl concentrations up to saturation of halite[END_REF] 273 < T < 523 K, P <1000 bar Ca 2+ / HO - [START_REF] Harvie | The prediction of mineral solubilities in natural waters: The Na-K-Mg-Ca-H-Cl-SO 4 -OH-HCO 3 -CO 3 -CO 2 -H 2 O system to high ionic strengths at 25°C[END_REF] not reported Na + / Cl - [START_REF] Pitzer | Thermodynamic properties of aqueous sodium-chloride solutions[END_REF] 273 K < T < 573K, P < 1000 bar Na + / HCO 3 - [START_REF] Millero | The dissociation of carbonic acid in NaCl solutions as a function of concentration and temperature[END_REF] 273 < T < 523 K Na + / OH - [START_REF] Pabalan | Thermodynamics of NaOH(aq) in Hydrothermal Solutions[END_REF] 273 K < T< 623 K, P < 400 bar Na + / CO 3 2- [START_REF] Millero | The dissociation of carbonic acid in NaCl solutions as a function of concentration and temperature[END_REF] 273 < T < 523 K Cl -/ CO 2 [START_REF] Duan | An improved model calculating CO 2 solubility in pure water and aqueous NaCl solutions from 273 to 533 K and from 0 to 2000 bar[END_REF] 273 K < T < 533 K, P < 2000 bar HCO 3

-/ CO 2 [START_REF] Kurz | Vapor-liquid-solid equilibria in the system NH 3 -CO 2 -H 2 O from around 310 to 470 K: New experimental data and modeling[END_REF] 310 < T < 470 K Ca 2+ / CO 2 [START_REF] Duan | Coupled phase and aqueous species equilibrium of the H 2 O-CO 2 -NaCl-CaCO 3 system from 0 to 250 degrees C, 1 to 1000 bar with NaCl concentrations up to saturation of halite[END_REF] 273 < T < 523 K, P <1000 bar Na + / CO 2 [START_REF] Duan | An improved model calculating CO 2 solubility in pure water and aqueous NaCl solutions from 273 to 533 K and from 0 to 2000 bar[END_REF] 273 K < T < 533 K, P < 2000 bar Na + / H 2 S [START_REF] Carrier | Contribution à l'étude thermodynamique de solutions d'électrolytes sous haute pression, haute température[END_REF] 298 K < T < 623 K HS -/ H 2 S

[32] 313 < T < 393 K, 100 bar H 2 S / Na + / Cl - [START_REF] Carrier | Contribution à l'étude thermodynamique de solutions d'électrolytes sous haute pression, haute température[END_REF] 298 K < T < 623 K Ca 2+ / Cl -/ CO 2 [START_REF] He | The carbonic acid system and calcite solubility in aqueous Na-K-Ca-Mg-Cl-SO 4 solutions from 0 to 90°C[END_REF] T < 363 K H 2 S / Na + / Cl - [START_REF] Carrier | Contribution à l'étude thermodynamique de solutions d'électrolytes sous haute pression, haute température[END_REF] 298 K < T < 623 K Ca 2+ / Cl -/ CO 2 [START_REF] He | The carbonic acid system and calcite solubility in aqueous Na-K-Ca-Mg-Cl-SO 4 solutions from 0 to 90°C[END_REF] T < 363 K Na + / Cl -/ CO 2 [START_REF] Duan | An improved model calculating CO 2 solubility in pure water and aqueous NaCl solutions from 273 to 533 K and from 0 to 2000 bar[END_REF] 273 K < T < 533 K, P< 2000 bar

Results and discussion 4 Solubility predictions

Once Soreide and Whitson's model and Pitzer's model provide respectively the fugacity and activity coefficients of components in the gas phase and in the liquid phase, the Ensemble Henry's law (eq. 1) is applied to assess the solubility of H 2 S and CO 2 in different conditions. Note that in the proposed approach, we do not perform a rigorous thermodynamic flash calculation, since the fugacity equality constraint in both phases is not checked for water. It is a limitation of this actual approach, but the results presented further show that this approximation does not significantly affect the accuracy of the model in the considered temperature and pressure ranges. The predictions for H 2 S solubility in water were compared to Ng et al. [START_REF] Ng | Water content of natural gas systems containing acid gas[END_REF] experimental data. These authors performed measurements of H 2 S solubility in water in the case of different CH 4 / C 3 H 8 / H 2 S / CO 2 gas mixtures at various temperatures. Figure 1 shows the concentration of dissolved H 2 S versus the total pressure of 75% CH 4 -C 3 H 8 (95:5 mole ratio) and 25 % H 2 S-CO 2 . The calculation of fugacity coefficients of gases was performed assuming propane to show similar interactions properties than methane with the others components. The new model gives excellent prediction of H 2 S with this set of experimental data. NaCl solution versus the total pressure at 40°C, 80°C and 160°C. Our new model also presents a good agreement with the experimental data, even in highly concentrated solution and on a wide range of temperature and pressure. For highly concentrated brine solutions, these results show that the application of the ensemble Henry's law combining Dhima's constants, Soreide and Whitson' model and Pitzer's model provides accurate predictions of CO 2 and H 2 S solubilities even under high temperature, high pressure. It is therefore appropriate for in situ pH calculation, as detailed in the next part of the paper.

pH predictions

The calculation method of pH is classically derived from the electroneutrality equation as described elsewere [START_REF] Crolet | pH measurements under high-pressures of CO 2 and H 2 S[END_REF]. Once H + concentration is determined, pH is calculated from H + activity according to :

  () log log . m H H H pH a m         (eq. 6)
In this expression, a H+ is the activity of H + , γ (m) H+ is the activity coefficient of H + on the molality scale and m H+ is the concentration of H + in mol / kg of water. Comparisons with different sets of experimental data are presented from Figure 3 to Figure 5. Crolet and Bonis [START_REF] Crolet | pH measurements under high-pressures of CO 2 and H 2 S[END_REF] reported pH measurements under 1 bar of H 2 S at 25°C and for different sodium chloride concentrations. Our model presents an excellent correlation with the experimental data up to 2 mol.L -1 NaCl (Figure 3). Another set of experimental data for highly concentrated solutions was proposed by Hinds al. [START_REF] Hinds | Technical note : measurements of pH in concentrated brines[END_REF], at 1 bar CO 2 and 25°C. As illustrated in Figure 4 the new model predicts the experimental results within less than 0,05 pH unit in the whole range of ionic strength from zero to more than 4 mol.L -1 NaCl. Validity of the new model at high pressure and temperature was evaluated from measurements by Crolet and Bonis [START_REF] Crolet | pH measurements under high-pressures of CO 2 and H 2 S[END_REF] and Meyssami et al [START_REF] Meyssami | Prediction of pH in model systems pressurized with carbon-dioxide[END_REF]. Figure 5.a illustrates measurements performed at 74°C in 0.5 mol.L -1 NaCl solutions. Figure 5.b presents measurements in pure water at 32°C. For both applications, an excellent prediction is obtained up to 50 bar of CO 2 .

Conclusion 5

The in-situ pH of formation water is one of the most important parameters for material selection in the oil and gas production. Most predictions tools currently employed were designed in the 80's and were adapted to temperatures lower than 150°C, maximum pressures of 50 bar and slightly concentrated solutions (up to 1 mol/L). However, the increasing interest of HPHT fields requires more accurate models with extended validity domain. For this purpose, we applied a calculation method based on the ensemble Henry's law for solubility calculations and taking into account the activity of chemical species for chemical equilibrium constants calculation.

The impact of high pressures on gas solubility was considered by fugacity coefficients calculation using Soreide and Whitson's model. The effect of high concentrations of salts on the activity of chemical species in the liquid phase was modeled using Pitzer's formalism.

Even though only sparse experimental data is available in the literature at high temperature, high pressure and in concentrated solutions, all comparisons between the new model and the measurements gave an excellent agreement. Due to the lack of experimental data for temperatures above 75°C and pressures higher than 50 bar, the validity of our model was not checked in these conditions. However considering the validity domains of individual elements of the model taken from the literature, we expect accurate predictions of in-situ pH in highly concentrated brines (up to 5 mol/L) under high pressure (around 1000 bar) and high temperature (to 250°C). This represents a considerable improvement of the actual tools, typically limited to 150°C and 50 bar. (0) , β (1) , β (2) , C Φ , ψ Pitzer's model parameters

Nomenclature

Figure 2 :

 2 Figure 2: Predictions of CO 2 solubility in NaCl 4 mol / kg of water at different temperatures

Figure 3 :

 3 Figure 3: Influence of NaCl concentration on pH at 25°C under 1 bar of H 2 S

Figure 4 :

 4 Figure 4: Influence of NaCl concentration on pH at 25°C under 1 bar of CO 2

Figure 5 :

 5 Figure 5: Influence of CO 2 pressure on pH (a) at 74°C in NaCl 0.5 mol.L -1 [36] and (b) at 32°C in pure water [38].

Predictions of H 2 S solubility in pure water under pressure of 75% CH 4 -C 3 H 8 (95:5) and 25% H 2 S-CO 2 (3:1) gas mixture at 49°C.

			0,04						Ng et al. [34]
									New model
	H 2 S
	of dissolved	fraction)	0,02
	Concentration	(molar	0,00
			0	100	200	300	400	500	600	700
									Pressure (bar)
					75% CH 4	-C 3	H 8	(95:5) and 25% H 2	S-CO 2	(3:1)
	Figure 1: The new model predictions for CO	2 solubility in water were compared to Rumpf et al. [35]
	experimental data. These authors measured CO	2 solubility in highly concentrated NaCl solutions under
	pressure of pure CO 2 at different temperatures. Figure 2 shows the concentration of dissolved CO 2 in 4
	mol.kg -1


J Kittel

S Gonzalez

E Lemaire

L Raynal

Corrosion in post-combustion CO 2 capture plants -comparisons between MEA 30% and new processes

CO 2 capture and storage (CCS) is one of the main option for CO 2 mitigation. Post-combustion capture processes using alkanolamines are considered as one of the preferred options for CCS, since their industrial viability was already demonstrated. Absorption -desorption by 30% monoethanolamine (MEA) represents the reference technology for post-combustion CO 2 capture. However, the cost of CO 2 removal with this MEA 30% process remains very large, and needs to be reduced. In CCS field, a lot of research program aim at developing new solvents, with a particular focus on the energy consumption needed at regeneration step. Minimizing solvent degradation is also often considered as the second immediate priority. Solvent degradation is usually caused by chemical reaction with oxygen, which is always present in post-combustion fumes. Other impurities, such as NO 2 or SO 2 , may also react with the solvent and form undesirable by-products. This paper is divided in two sections. In the first part, we propose a review of CO 2 capture with monoethanolamine, which represents the reference situation. Focus is given to corrosion aspects, including laboratory results and pilot plant studies. In the second part of the paper, we present two ways of improving CO 2 capture processes that are currently studied at IFPEN. The first one is an evolution of the reference 30% MEA process, using a higher MEA concentration in order to reduce the energy consumption. The second one uses a newly developed solvent, with improved performances. For both cases, specific corrosions tests will be presented, both in laboratory conditions and in pilot plant operations.

Introduction -post-combustion CO 2 capture with alkanolamines

Post-combustion CO 2 capture with 30 % mass monoethanolamine represents the most widely accepted reference process. It is based on the high chemical affinity of the alkaline amine with CO 2 , following the global reaction: The main reaction products are the amine carbamate and the protonated amine. This chemical equilibrium is greatly influenced by temperature. It is shifted to the left at high temperature and high pressure. This property is put into advantage in the industrial process, which consists of successive absorption -desorption in a loop system.

Figure 1 provides a simplified process flow description of this type of treatment unit. The flue gas which contains high concentration of CO 2 enters in the treatment plant at the bottom of the absorber. Lean MEA is introduced at the top of the column, and chemical reaction between the amine and CO 2 takes place. At the liquid output at the bottom of the absorber, the solvent is enriched in acid gas: we speak of rich amine. At the top of the absorber, the gas has been stripped of its CO 2 .

The rich amine is then pre-heated by a heat exchanger then fed into the top of a regeneration column. In this part of the unit, the solvent is raised to high temperature by steam, which releases the dissolved CO 2 . At the liquid output of the regenerator, the solvent is hot and contains less acid gas: we speak of lean amine. The solvent is then cooled by the heat exchanger and sent back to the top of the absorber to start a new cycle. Pure CO 2 is collected at the top of the regenerator. Before sequestration, it has to be compressed to 100 bar for transportation. This compression step also represents an important penalty in terms of energy consumption.

Section 1 -Reference process with 30% MEA

Performance indicators for the 30% MEA process

30% MEA represents the reference case for CO 2 capture process. In most recent pilot plant studies, this solvent was tested for comparison with newly developed solutions. An excellent review of research programs ongoing worldwide on CO 2 absorption was proposed last year by Wang et al. [START_REF] Wang | Post-combustion CO2 capture with chemical absorption: A state-of-the-art review[END_REF].

A complete description of the performances of CO 2 capture process with MEA far exceeds the goals of this paper. However, a few key features must be given, in order to explain the research roads described in the second section of the paper.

The first major parameters for efficient CO 2 post-combustion capture are chemical reactivity and reaction kinetic between the solvent and CO 2 . Indeed, Post-combustion fumes are produced at a high flow rate and with a low CO 2 partial pressure. The solvent needs to be extremely reactive to allow a high level of CO 2 transfer from the gas phase to the liquid phase. High reaction kinetics is also an important feature: the lower the kinetics, the larger the contact time between the gas and liquid phase to achieve the target CO 2 removal. Thus, decreasing the solvent kinetics can only be mitigated by increasing the height of the absorber tower, or by using more efficient contactors, which a huge impact on capital costs. Regarding this specific aspect, MEA represents an excellent solvent. CO 2 loading, α, in the solvent at equilibrium with 100 mbar CO 2 is close to 0.5 mol CO2 /mol MEA. The reaction kinetics is also extremely fast, allowing reasonable dimensions of industrial equipments for the absorption.

The second performance indicator for solvent ranking is the regeneration energy, i.e. the energy needed to release the absorbed CO 2 . This energy is usually provided by steam generated in the reboiler, and constitutes a direct penalty for the power plant performance. MEA 30% is known to be very demanding in terms of energy consumption. Indeed, about 3.7 GJ/t CO2 are required for solvent regeneration [START_REF] Knudsen | Experience with CO2 capture from coal flue gas in pilot-scale: Testing of different amine solvents[END_REF][START_REF] Abu-Zahra | CO2 capture from power plants: Part I. A parametric study of the technical performance based on monoethanolamine[END_REF], representing the most important part of the operating costs.

The last two important parameters regarding industrial feasibility of chemical absorption is the solvent stability and corrosivity to construction materials. MEA is known to react quite rapidly with O 2 contained in the flue gas, with two major consequences: i/ degraded solvent becomes less reactive, and needs to be periodically cleaned and replaced; ii/ some degradation products are corrosive to steel. Pilot plant studies conducted during the EU funded Castor project have shown that the degradation of the 30% MEA process corresponded to 1.4 kg MEA degraded / t CO2 captured [START_REF] Lemaire | New IFP optimized first generation process for post-combustion carbon capture: HiCapt+[END_REF]. At the industrial scale, this degradation corresponds to the consumption of approximately two batches of solvent per year, representing a cost of 7 M€/year for a CO 2 capture plant in a 600 MWe power plant. Corrosion is also a major drawback of most gas treating units using amines. Corrosion can be caused either by the carboxylic acids (i.e. acetic acid, formic acid and oxalic acid formed by oxidative degradation), or by the amine itself.

A brief summary of actual knowledge of corrosion reactions in 30% MEA is given below.

Focus on corrosion

Several papers presenting laboratory corrosion measurements in MEA solutions have been published in the last years [START_REF] Fleury | Corrosion in amine solvents used for the removal of acid gases[END_REF][START_REF] Kittel | Corrosion in alkanolamine used for acid gas removal: From natural gas processing to CO2 capture[END_REF][START_REF] Soosaiprakasam | Corrosion and polarization behavior of carbon steel in MEA-based CO2 capture process[END_REF][START_REF] Veawab | Corrosion behavior of carbon steel in the CO2 absorption process using aqueous amine solutions[END_REF][START_REF] Veawab | Identification of oxidizing agents in aqueous amine-CO2 systems using a mechanistic corrosion model[END_REF]. It appears that carbon steel corrosion rate is mainly influenced by temperature and CO 2 loading in the solvent, as illustrated in Figure 2.

Experiments at various MEA concentrations between 20 and 40% showed that the increase of amine concentration had a detrimental impact on corrosion rate, but in a much lower magnitude than temperature and CO 2 loading [START_REF] Kittel | Corrosion in alkanolamine used for acid gas removal: From natural gas processing to CO2 capture[END_REF][START_REF] Veawab | Corrosion behavior of carbon steel in the CO2 absorption process using aqueous amine solutions[END_REF].

The impact of heat stable salts on MEA corrosivity was also reported [START_REF] Fleury | Corrosion in amine solvents used for the removal of acid gases[END_REF][START_REF] Tanthapanichakoon | Electrochemical investigation on the effect of heat-stable salts on corrosion in CO2 capture plants using aqueous solution of MEA[END_REF]. Among the most common heat-stable salts, oxalic acid seems to have the largest impact on corrosion. However, the mechanism is more related to decreasing the protectivity of corrosion scale by a chelating effect, rather than to an electrochemical activity.

From all laboratory studies, it appears that carbon steel grades cannot be used safely in all operating conditions encountered in a MEA capture unit, with predicted corrosion rates exceeding 100 µm/year [START_REF] Kittel | Corrosion in alkanolamine used for acid gas removal: From natural gas processing to CO2 capture[END_REF]. Thus, many studies are dedicated to corrosion inhibitors [START_REF] Soosaiprakasam | Inhibition performance of copper carbonate in CO2 absorption process using aqueous MEA[END_REF][START_REF] Tanthapanichakoon | Polarization behavior and performance of inorganic corrosion inhibitors in monoethanolamine solution containing carbon dioxide and heat-stable salts[END_REF]. Unfortunately, most efficient inhibitors, such as salts of heavy metals, often adversely affect amine degradation. It also appears from most laboratory studies that austenitic stainless steel grade 316L seems to perform well [START_REF] Kittel | Corrosion in alkanolamine used for acid gas removal: From natural gas processing to CO2 capture[END_REF][START_REF] Tanthapanichakoon | Electrochemical investigation on the effect of heat-stable salts on corrosion in CO2 capture plants using aqueous solution of MEA[END_REF]. All these corrosion results have been compared, at various levels, with field data from pilot plants.

Excellent correlation between laboratory evaluation and pilot plant data were found for carbon steel. Experiments in the Castor pilot plant operated by Dong Energy in Denmark or in the pilot unit of the International Test Centre for CO 2 Capture (ITC) in Canada showed corrosion rates exceeding several hundreds of µm/year in the hot rich amine sections [START_REF] Kittel | Corrosion in MEA units for CO2 capture: Pilot plant studies[END_REF].

Good performance of stainless steel grade AISI 316L was also demonstrated by specific corrosion experiments in the Castor pilot plant [START_REF] Kittel | Corrosion in MEA units for CO2 capture: Pilot plant studies[END_REF], as well as in the RWE pilot plant in Niederaussem, Germany [START_REF] Moser | Material testing for future commercial post-combustion capture plants. Results of the testing programme conducted at the Niederaussem pilot plant[END_REF].

Discussion -roadmap for lowering the cost of 30% MEA process

Considering MEA 30% as the benchmark, new generation processes for post-combustion CO 2 capture must provide significant improvements for the key performances. With that aim, CO 2 capture cost reduction is the main driving force for all R&D activities in the domain. Two main types of costs must be considered with equal importance. Operating costs (OPEX) include power needs of ancillary equipments (pumps, blowers, compressors,...), reboiler heat for solvent regeneration, solvent reclaiming for the removal of degradation products, and solvent make-up to replace degraded amine. Recent multi-scale simulations of the costs of CO 2 capture showed that the OPEX represents approximately from half to 2/3 of the total cost of avoided CO 2 depending if one considers an integrated plant or not [START_REF] Raynal | CO2 capture cost reduction: use of a multiscale simulations strategy for a multiscale issue[END_REF][START_REF] Raynal | From MEA to demixing solvents and future steps, a roadmap for lowering the cost of post-combustion carbon capture[END_REF]. According to the latter studies, the energy consumption directly used for the regeneration of 30% MEA, estimated at 3.7 GJ/t CO2 , is by far the first item in terms of cost since it represents from 30 to 60% of the total CO 2 capture cost. This explains why a lot of research programs tend to develop new solvents requiring less energy consumption for the regeneration.

On the other hand, capital costs must not be forgotten. CAPEX is mainly influenced by the size of the equipments, and more precisely the absorber and the regenerator. Thus, decreasing the reaction kinetics has a direct impact on column height and cost. Material selection is also an important parameter defining the final cost.

In the next section of this paper, two different alternatives currently investigated at IFPEN are described.

The first one consists in an optimization of the MEA process, by increasing the solvent concentration, resulting in improvements of energy needs.

The second one is a more radical evolution, and uses a new generation demixing solvent, allowing a significant reduction of regeneration energy, as well as a high temperature regeneration at high pressure.

Section 2 -Research in progress for high performance CO 2 capture

Optimization of energy consumption with MEA -Hicapt +(TM) process

As discussed in the previous part of this paper, CO 2 post-combustion capture with MEA presents many drawbacks. The two more important are the energy consumption for the regeneration and the oxidative degradation. Corrosion is also an important operational problem, but there exists efficient metallurgical solutions.

In spite of these difficulties, MEA remains one of the most widely investigated solvent for CO2 post-combustion capture. Indeed, it is cheap, easily available, non toxic and highly effective due to its high capacity and fast kinetics. Several large scale projects, e.g. Castor pilot plant or the Enel pilot plant in Brindisi [START_REF] Lemaire | Results of the 2.25 t/h post-combustion CO2 capture pilot plant of ENEL at the Brindisi coal power plant and last R&D developmnets for Hicapt+ process[END_REF] have demonstrated the good operability, flexibility, stability and reliability of this process during long test runs.

Based on all the knowledge acquired during pilot tests with 30% MEA, it was found that increasing the amine concentration to approximately 40% could result in significant reduction of energy demand for the regeneration. This feature was verified in the pilot plant operated by ENEL in the Brindisi coal fired power plant [START_REF] Lemaire | Results of the 2.25 t/h post-combustion CO2 capture pilot plant of ENEL at the Brindisi coal power plant and last R&D developmnets for Hicapt+ process[END_REF]. Two successive test runs were conducted with 30% and 40% MEA, resulting respectively in a regeneration energy of 3.4 GJ/t CO2 and 3.0 GJ/t CO2 . Unfortunately, it is known also that increasing MEA concentration usually results in a strong aggravation of the degradation. This was verified in laboratory experiments, showing that MEA consumption was doubled when the concentration was raised from 30% to 40% At this level, operation of the process would not be economically viable: it would require extremely frequent reclaiming of the solvent, and MEA consumption would represent up to four batches per year, which represents an important cost even for the cheap MEA. At the same time, the associated cost of degraded solvent disposal could become a problem. An important part of our research on MEA based process was therefore dedicated to minimizing the oxidative degradation [START_REF] Martin | New Amines for CO2 Capture. IV. Degradation, Corrosion, and Quantitative Structure Property Relationship Model[END_REF][START_REF] Delfort | MEA 40% with improved oxidative stability for CO2 capture in post-combustion[END_REF][START_REF] Lepaumier | New Amines for CO2 Capture. I. Mechanisms of Amine Degradation in the Presence of CO2[END_REF][START_REF] Lepaumier | New Amines for CO2 Capture. II. Oxidative Degradation Mechanisms[END_REF]. More than 200 candidate oxidation inhibitors belonging to more than fifteen different chemical families were tested. Some molecules with excellent properties were found, as illustrated in Figure 3. These oxidation tests were performed at 80°C under continuous air flow. Under these conditions and without inhibitor, the degradation of MEA proceeds extremely fast, forming NH 3 as typical ultimate product. This gaseous component can easily be monitored in the gas stream by FTIR analysis, giving a good indication of degradation process. It is quite obvious from these results that several chemical species present excellent oxidation inhibition performances. The direct results of this research is the Hicapt+ (TM) process developed by IFPEN and PROSERNAT. It uses a 40 % MEA solvent with an efficient oxidation inhibitor, thus combining high energy efficiency and low degradation.

Unfortunately, the more efficient oxidation inhibitors interact with the usual stainless grades used in the conventional 30 % MEA processes These inhibitors have a strong reducing capability, which may affect the passive layer of 304L and 316L. Small pilot plant studies conducted at IFPEN have shown uniform corrosion rates close to 0,5 mm/year for 316L coupons in the hot rich solvent. For this reason, it was not possible to test the Hicapt +(TM) process at pilot scale to date, since most available pilot units are made of 304L and 316L.

In order to solve this corrosion problem associated with the oxidation inhibitor, two types of solutions were investigated and gave satisfactory results. Efficient corrosion inhibitors were found, enabling the maintain of stainless steel passivity in the presence of the oxidation inhibitor. However, the more efficient corrosion inhibitors were strong oxidants, and they appeared to interact with the degradation inhibitor efficiency. This solution was thus abandoned. Investigations were then dedicated to find corrosion resistant materials. Solutions with a weak economical impact were found, and should be presented in more details in the near future. Small scale pilot plant with appropriate materials are also currently under construction at IFPEN, in order to demonstrate the efficiency of the solution in real operating conditions.

New generation of solvent -DMX (TM) process

Although MEA represents the most easily accessible solvent for short term commercial applications, its global performances remain extremely penalizing with regard to the cost of avoided CO 2 . New generation processes must be developed to significantly lower the overall costs, and allow the largest dissemination of post-combustion technologies. A lot of research is in progress worldwide, with the objective to find new amine or blend of amines with improved performances. The great number of daily publications on this field does not allow to propose an exhaustive list of references, but a recent review of some of the most important research programs worldwide was proposed recently by Wang et al. [START_REF] Wang | Post-combustion CO2 capture with chemical absorption: A state-of-the-art review[END_REF].

In order to focus our research on the most important parameters, an analysis of process and design impact on the cost of avoided CO 2 was realized [START_REF] Raynal | From MEA to demixing solvents and future steps, a roadmap for lowering the cost of post-combustion carbon capture[END_REF].

Cost repartition between OPEX and CAPEX was first established for the reference case using 30 % MEA for a non integrated plant. The main findings are presented in Table 1. Based on this analysis, individual impact of key parameters for the development of new solvents were discussed. The main characteristics of the solvent can be divided in four main domains:

-the thermodynamic performances, defining the cyclic capacity, the heat of reaction and the stripping energy, -the kinetic performances, -the thermal and chemical degradation, -the operational performances, including volatility, foaming, corrosion, viscosity, ... Cross analysis with cost impact allowed to determine the most efficient targets for new solvents:

-good thermodynamic properties, with a particular focus on the cyclic capacity, -low thermal and chemical degradation for possible regeneration under pressure, -good kinetics. With these objectives in mind, IFPEN performed intensive high throughput screening (HTS) for thermodynamic and degradation evaluations [START_REF] Martin | New Amines for CO2 Capture. IV. Degradation, Corrosion, and Quantitative Structure Property Relationship Model[END_REF][START_REF] Porcheron | High Throughput Screening of CO2 Solubility in Aqueous Monoamine Solutions[END_REF][START_REF] Porcheron | High throughput screening of amine thermodynamic properties applied to post-combustion CO2 capture process evaluation[END_REF]. A very interesting class of solvents was discovered in this investigation, which are capable of two-phase separation when they are heated and loaded with CO 2 . This process in now developed by IFPEN and PROSERNAT under the name DMX (TM) . In this class of solvents, the CO 2 captured concentrates in one of the two liquid phases, and only this portion has to be sent to the stripper. This allows substantial reduction of the energy requirement for the regeneration of the solvent. In a first step, fine tuning of demixing solvents properties had to be determined, as a function of CO 2 loading and temperature [START_REF] Aleixo | Physical and chemical properties of DMXTM solvents[END_REF]. It was particularly important to make sure that two-phase separation occurred only at high temperature, i.e. could not appear in the absorber where it would strongly affect gas -liquid transfer. evaluation with the best candidates showed that the reboiler duty could be reduced down to 2.3 GJ/t CO2 , which is a strong gap with 30% MEA (3.7 GJ/t CO2 ) or with 40% MEA (3.0 GJ/t CO2 ).

Once the more promising DMX TM amine was chosen, degradation and corrosion tests were performed. These tests are not easy to perform, since phase separation is expected in closed reactors at high temperature. Weight loss coupons were then placed at two distinct heights in the reactor, where thermodynamic calculations predict two phases. In order to evaluate the solvent behavior in extreme conditions, corrosion tests were performed at 160°C and under a high pressure of CO 2 (27 and 55 bar). Under these conditions, the global CO 2 loading is respectively 0.3 and 0.7 mol CO2 /mol amine . Most of the absorbed CO 2 is in the dense phase. Corrosion rates determined by weight loss measurements after three weeks exposure are presented in Table 2. The chemical composition of the steel coupons is given in Table 3. These corrosion rates are extremely promising for the development of the process using this solvent. For comparison, the corrosion rate of carbon steel in CO 2 rich 30 % MEA at 110°C exceeds 0.5 mm/year.

From similar experiments, it was shown also that this solvent exhibited far better resistance to thermal and oxidative degradation than MDEA, which is already considered to have an excellent degradation behavior [START_REF] Raynal | From MEA to demixing solvents and future steps, a roadmap for lowering the cost of post-combustion carbon capture[END_REF][START_REF] Raynal | The DMX process: An original solution for lowering the cost of post-combustion carbon capture[END_REF].

Thanks to these excellent properties, high temperature regeneration of the solvent can be considered. As a consequence, gaseous CO 2 at the stripper outlet can be obtained at a pressure of several bar, minimizing the cost of compression to 110 bar required before CO 2 transport and storage.

The next steps for the development of this process is to verify that all promising performances seen in the laboratory and in small scale pilot units will be confirmed on large scale pilot experiments. This pilot plant demonstration represents one of the objectives of the EU funded Octavius project which started in March 2012, coordinated by IFPEN and with 16 other partners.

Conclusions

The present paper describes some of the research in progress at IFPEN on 2 postcombustion capture with chemical solvents. At present time, monoethanolamine represents the benchmark solvent for this application. The main performances of this process were resumed, showing that the cost of avoided CO 2 remained quite high. It was shown also that this solvent was quite corrosive to carbon steel, and required to use stainless steel as construction material.

Then, two different new generation processes were presented. The first one consists mainly of an optimisation of the usual MEA process. The energy consumption is lowered by approximately 15%, thanks to an increase of MEA concentration from 30% to 40%. The consequences of this higher concentration are discussed in terms of degradation and corrosion. It is shown that the risks of oxidative degradation increase considerably, but can be significantly reduced by efficient oxidation inhibitors. Corrosion needs also to be taken into account. It was shown that oxidation inhibitor could interact with the passive layer of usual austenitic stainless steel grades, allowing fast uniform corrosion. Solutions with more resistant materials were found to avoid this problem.

The second process uses a radically different type of amine. Furthermore, it presents the very particular trend to separate in two liquid phases at high temperature and when it is loaded with CO 2 . Then, only the CO 2 rich liquid phase needs to be heated in the regenerator for CO 2 stripping, which represents a net benefit for energy consumption. This new solvent also presents excellent thermal resistance up to 160 -180°C, as well as resistance to oxidative degradation. High temperature regeneration allows to produce CO 2 at several bars, allowing substantial compression gains. Finally, corrosion by this new solvent appeared to be much less than with conventional MEA, or even with usual amines used in gas treatment (DEA or MDEA). Corrosion rates as low as 10 µm/year were observed for carbon steel even at the highest temperature expected for the process. Therefore, less costly construction materials should be employed, and the use of expensive proprietary corrosion inhibitors will be avoided.

Figure 1 :

 1 Figure 1: Simplified diagram of a MEA CO 2 capture unit

Figure 2 :

 2 Figure 2: Impact of temperature and CO 2 loading on the corrosion rate of carbon steel in 30% MEA [6]

Figure 3 :

 3 Figure 3: Effect of oxidation inhibitors on NH 3 evolution from MEA 40% [19]

Table 1 : Cost repartition for the 30 % MEA process

 1

	CAPEX (€/t CO2 )	OPEX (€/t CO2 )	TOTAL (€/t CO2 )
	Stripper	4.7	44.5	49.2
	Compressor	5.1	5.1	10.2
	Absorber	4.7	/	4.7
	Solvent	0.2	2.4	2.6
	Washing tower	2.2	/	2.2
	Heat exchanger	2.0	/	2.0
	Pumps	1.0	0.4	1.4
	Blower	0.1	1.0	1.1
	Miscelaneous	0.7	0.3	1.0
	TOTAL	20.7 (28 %)	53.7 (72 %)	74.4
	Steam price: 21.7 €/t
	Electricity price: 45 €/MWh

Table 2 : Results of corrosion experiments in the DMX TM solvent at high temperature Corrosion rate (µm/an)

 2

	Test condition	Steel grade	light phase	dense phase (CO 2 rich)
	160°C / 27 bar	AISI 1018 AISI 316L	1 ± 2 µm/an 3± 2 µm/an	1 ± 2 µm/an 3 ± 2 µm/an
	160°C 55 bar	AISI 1018 AISI 316L	9± 2 µm/an 8 ± 2 µm/an	8 ± 2 µm/an 8 ± 2 µm/an

Table 3 : Chemical composition of weight loss coupons (wt. %)

 3

		C	Si	Mn	Ni	Cr	Mo	N	S	P	Al	Cu
	AISI 316L	0.021 0.448 1.335 10.295 16.861 2.161 0.047 0.002 0.029
	AISI 1018	0.186 0.223 0.560	0.084	0.106	0.013		0.022 0.009 0.015 0.261