Natural Gas Sweetening with Ionic Liquids A Selectivity Analysis. Report Number MSc Thesis by A. Amplianitis February 2014

Natural Gas Sweetening with Ionic Liquids A Selectivity Analysis Report Number 2607 MSc Thesis by A. Amplianitis February 2014

Natural Gas Sweetening with Ionic Liquids A Selectivity Analysis by A. Amplianitis In partial fulfillment of the requirements for the degree of Master of Science in Sustainable Energy Technology (SET) at the Delft University of Technology, to be defended publicly on Thursday February 13, 2014 at 14:00. Supervisor: Prof.dr.ir. Thijs Vlugt Thesis committee: Dr.ir. Theo de Loos Dr.ir. Klaas Besseling Ir. Mahinder Ramdin Report number: 2607 An electronic version of this thesis is available at http://repository.tudelft.nl/.

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Abstract Global natural gas consumption has grown significantly over the years rendering natural gas one of the most important energy sources of the future. Although the known natural gas resources are significant, half of the known gas fields are estimated to contain more than 2% CO 2 rendering them sub-quality reserves. Apart from carbon dioxide, natural gas also contains other sour gasses like hydrogen sulfide (H 2 S). These impurities need to be removed from natural gas because in the presence of water, acids that can corrode pipelines and other processing equipment are formed. Furthermore, as CO 2 provides no heating value it has to be removed in order to meet gas quality specifications before distribution to customers. Gas processing to remove acid gasses is referred to as natural gas sweetening. Among the materials investigated as new potential solvents for CO 2 absorption processes, ionic liquids are one category of solvents that may in the future offer an alternative to amines and low capacity physical solvents. CO 2 solubility and selectivity data in ILs are essential if we are to judge the separation performance of these solvents. The goal of the present thesis was to experimentally determine the solubility of methane in two phosphonium based ionic liquids for which CO 2 solubility data were already available. In that way the ideal selectivity of CO 2 /CH 4 in those ILs was determined. Furthermore, mixed gas solubilities were experimentally determined in ternary systems of IL+CO 2 +CH 4. As the acquired experimental data are not enough in order to calculate real selectivities, the present work focuses on comparing the mixed gas bubble-point pressures to the sum of the single gas bubble-point pressures acquired from measurements in binary systems. The ultimate goal is to determine whether or not there is significant interaction between the two gases that can eventually result in real selectivity deviating from ideal selectivity. Finally, solubility of CO 2 in an imidazolium-based IL is experimentally determined. iv

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Acknowledgements First of all I would like to thank Prof.Thijs Vlugt as well as the entire Engineering Thermodynamics group for giving me the opportunity to be part of a highly skilled team and work in the forefront of chemical engineering research. I would like to give my special thanks to Prof. Theo de Loos for his indispensable guidance and help week after week on all the theoretical and technical issues that arose during my time in the thermolab. I would like to express my outmost gratitude to Mahinder Ramdin. Without his guidance and uninterrupted help, this dissertation would not have been possible. My experimental work in the Engineering Thermodynamics laboratory would not have been possible without the valuable help and continuous guidance of Eugene Straver. Furthermore I want to thank Mariette de Groen and María Teresa Mota Martínez for their support and presence in the thermolab. I would like to thank my dear friends Georgiana, Myrto, Lina, Despoina, Spyros, Christina and Constantinos for their continuous moral support. Last and most importantly, I would like to thank my parents who enabled my studies abroad by supporting me both morally and financially. vi

Table of contents ABSTRACT IV ACKNOWLEDGEMENTS VI 1. INTRODUCTION 1 2.1 NATURAL GAS SWEETENING 4 2.1.1 LIQUID-PHASE PROCESSES 4 2.1.2 DRY-BED PROCESSES 7 2.1.3 MEMBRANES 8 2.1.4 CRYOGENIC FRACTIONATION 8 2.2 IONIC LIQUIDS 9 2.3 CO 2 /CH 4 SELECTIVITY 10 3. EXPERIMENTAL PROCEDURE 12 3.1 PHASE BEHAVIOR THEORY 12 3.2 EXPERIMENTAL METHOD 13 3.3 MEASUREMENTS AND CORRECTIONS 16 3.2 IONIC LIQUIDS USED 17 4. MODELING THE PHASE BEHAVIOR OF BINARY SYSTEMS 18 4.1 PENG-ROBINSON EQUATION OF STATE 18 4.2 VAPOR-LIQUID EQUILIBRIA (VLE) 20 4.3 MATLAB 22 5. RESULTS 23 5.1 BINARY SYSTEMS 23 5.1.1 THTDP-DICYANAMIDE 23 5.1.2 THTDP-PHOSPHINATE 27 5.1.3 EMIM-PHOSPHATE 31 5.2 TERNARY SYSTEMS 33 TERNARY SYSTEMS SUMMARY 39 6. CONCLUSION AND FUTURE DIRECTION 40 6.1 CONCLUSION 40 6.2 FUTURE DIRECTIONS 42 vii

BIBLIOGRAPHY 43 APPENDIX A 46 CALCULATION OF THE SECOND VIRIAL COEFFICIENTS OF PURE GASSES AND GAS MIXTURES. 46 APPENDIX B 48 EXPERIMENTAL MEASUREMENTS 48 viii

1. Introduction Global natural gas consumption has grown significantly over the years rendering natural gas as one of the most important energy sources of the future. 1 It is thought to be the most environmentally friendly fossil fuel, since burning it leads to negligible SO 2 emissions, low NO x levels and less than half of the CO 2 emitted when burning coal or oil. 1 In 2010, natural gas supplied 23.81% of the world s energy demand and the volume of the consumed natural gas increased by 7.4% over 2009 levels. 3 The volumes of natural gas traded as pipeline and liquefied natural gas (LNG) in 2012 is shown in Table1. Although the known natural gas resources are significant, half of the known gas fields are estimated to contain more than 2% CO 2 rendering them sub-quality reserves. 2,4 The increased demand for natural gas has inevitably led to a re-evaluation of the development potential of those sub-quality gas reserves that had been previously considered economically unviable. Apart from carbon dioxide, natural gas also contains other sour gasses like hydrogen sulfide (H 2 S). These impurities need to be removed from natural gas because in the presence of water, acids are formed that can corrode pipelines and other processing equipment. 2 Furthermore, as CO 2 provides no heating value it has to be removed in order to meet gas quality specifications before distribution to customers. Those specifications usually stipulate a maximum level of 3% by volume CO 2 in natural gas transmitted to customers by pipeline. 2,5 As most of the significant natural gas resources are located far from the large and established natural gas markets, significant volumes of natural gas need to be transported as LNG. 2 The specifications of CO 2 removal from natural gas to be processed in a cryogenic plant to produce LNG are far more stringent than those for typical gas pipelines. Carbon dioxide concentration needs to be less than 50 ppm before the sweetened gas enters the cryogenic processes within the plant in order to avoid the formation of dry ice. 2,5 Top natural gas exporters Top natural gas importers 1 2 3 4 5 6 7 8 9 Russian Federation Norway Qatar Canada Algeria Netherlands Indonesia Malaysia U.S. Pipeline LNG Total Pipeline LNG Total 186.5 95.9 19.2 92.4 36.5 53.3 9.9 1.5 30.3 13.4 4.7 75.8 0.0 19.3 0.0 31.4 30.5 1.6 199.9 100.6 94.9 92.4 55.8 53.3 41.3 32.0 32.0 U.S. Japan Germany Italy U.K. France South Korea Turkey Spain 93.3 0.0 92.8 66.3 35.0 35.0 0.0 28.8 8.9 12.2 93.5 0.0 9.1 18.7 13.9 44.4 7.9 27.5 105.5 93.5 92.8 75.3 53.6 48.9 44.4 36.7 36.4 10 Australia 0.00 25.4 25.4 Ukraine 33.0 0.0 33.0 Table 1. Volumes of natural gas traded as pipeline and LNG (billions of cubic meters) by top natural gas exporting and importing countries in 2012. 2,3

The development therefore of the sub-quality natural gas reserves together with the increased LNG production and commercialization have presented new challenges to gas processing that require more efficient approaches to the conventional absorption and separation technologies that are most commonly used for CO 2 removal from natural gas. Various technologies have been practiced both on laboratory and industrial scale for natural gas sweetening. These processes involve chemisorption/physisorption, membrane separation or molecular sieves, amine physical absorption, carbamation, amine dry scrubbing and mineral carbonation. 6 Currently, the most vastly used methods for natural gas sweetening purposes are amine-based processes. Alkanolamines such as monoethanolamine (MEA), diethanolamine (DEA), and methyldiethanolamine (MDEA) based methods are vastly used and account for over 95% of all gas sweetening in the United States. 6,7 There are however certain disadvantages for the commercial use of these solutions. Main issues are the loss of amine reagents, transfer of water into the gas stream during the desorption stage, degradation of the amine reagent to form corrosive byproducts and high-energy consumption during regeneration. 6 The regeneration step may account for the 70% of the total operating costs of the capture process. 6 Issues like that have to be overcome with new, more efficient and less costly methods. Among the materials investigated as new potential solvents for CO 2 absorption processes, ionic liquids (ILs) are one category of solvents that may in the future offer an alternative to amines and low capacity physical solvents. ILs are commonly defined as organic salts with melting temperatures of less than 373K 2. Properties of ILs such as extremely low vapor pressure, high chemical/thermal stability, no flammability and in certain cases zero toxicity may render them useful replacements for volatile organic solvents. 2,6 Furthermore the physical and chemical properties of ILs can be enhanced and modified by altering both their cationic and anionic moieties. 7 In that way, IL properties can be tailor-designed to meet the specific needs for natural gas sweetening processes: high solubility of CO 2, increased CO 2 selectivity, high thermal stability and low toxicity. The fact that CO 2 is inherently soluble in many ionic liquids has prompted many researchers to explore the vast synthetic landscape provided by ILs. Initially most of the studies focused on ILs with imidazolium-based cations because of their observed affinity towards CO 2. 6,7 Several of those ILs have been studied with regard to their CO 2 solubility and in order to understand phase behavior of CO 2 -IL pairs. 6,8 Combinations of different anions and cations however, give a huge number of possible ILs that need to be studied. Furthermore, as separation processes involve mixtures of two or more components that need to be separated, apart from solubility data, selectivity data are also essential if we are to judge the separation performance of a solvent. Although there is great availability of CO 2 solubility data in the literature, selectivity data of CO 2 in ILs are scarse. 7 Furthermore these selectivity data mainly refer to ideal selectivities (i.e. the ratio of pure gas solubilities) and not real selectivities. Real or actual selectivities cannot always be accurately determined from pure gas solubilities with the assumption of ideal mixing. 7 However since 2

measuring mixed-gas solubilities is significantly more difficult there are almost no mixture data available in the literature. The selectivities that are mainly relevant for CO 2 capture from natural gas are CO 2 /CH 4 and CO 2 /H 2 S. 6,7 With all the above in mind, the goal of the present thesis was to experimentally determine the solubility of methane in two phosphonium based ILs for which CO 2 solubility data were already available. In that way the ideal selectivity of CO 2 /CH 4 in those ILs was determined. Furthermore mixed gas solubilities were experimentally determined in ternary systems of IL+CO 2 +CH 4. As there is no single straightforward way to calculate real selectivities, the present work focused on comparing the mixed gas bubble-point pressures to the sum of the single gas bubble point pressures acquired from measurements in binary systems. The ultimate goal was to determine whether or not there is significant interaction between the two gases that would eventually result in real selectivity deviating from ideal selectivity. Finally solubility of CO 2 in an imidazolium-based IL was experimentally determined. The experimental work involved in the present thesis can be summarized in the next points: Experimental study of the gas-liquid equilibrium of ionic liquid 1-Ethyl-3-methylimidazolium diethyl phosphate ([emim]-[phosphate]) with CO 2. Experimental study of the gas-liquid equilibrium of the phosphonium-based ionic liquids trihexyl-tetradecyl-phosphonium-dicyanamide([thtdp][dca]) and trihexyl-tetradecylphosphonium-bis(2,4,4-trimethyl-pentyl)-phosphinate ([thtdp][phosphinate]) with methane and subsequent modeling of the two binary systems with the Peng-Robinson EoS. Experimental study of the gas-liquid equilibrium of the ternary systems of the two phosphonium-based ionic liquids, mentioned above, with gas mixtures of CO 2 and CH 4. The content of this thesis is divided in five chapters. After this brief introduction, a detailed literature review of the various gas separation technologies currently employed is presented in Chapter 2. In Chapter 3 the experimental procedure and set-up is presented. Chapter 4 is dedicated to binary systems modeling using Equations of State. Chapter 5 summarizes all the experimental and modeling results for binary and ternary systems. Finally Chapter 6 outlines the conclusions that were drawn and presents future research directions. 3

2. Literature study 2.1 Natural Gas Sweetening As discussed previously, raw natural gas contains large amounts of acid gases such as carbon dioxide (CO 2 ) and hydrogen sulfide (H 2 S). Those gasses are referred to as acid gasses, because they dissolve in water to form weak acids. 2,6 Furthermore, natural gas can also contain other contaminants such as carbonyl sulfide (COS), mercaptans (R-SH) and carbon disulfide (CS 2 ). 9 Many issues arise from the presence of these contaminants. Acids formed by the dissolution of acid gasses to water can corrode pipelines. Combustion of sulfur-based contaminants produces serious pollutants responsible for acid rain. Furthermore, those sulfur compounds are inherently poisonous to humans and animals and are corrosive to metals and other material used for the transportation and handling of natural gas. 2,9 Another important issue is that since CO 2 is non-flammable, its presence reduces the heating value and therefore the sale value of natural gas. 2 It is therefore clear that the removal of acid gasses and sulfur compounds is a major issue in natural gas processing. The major challenge in natural gas sweetening processes is to remove all the contaminants to as low a level as possible. Nowadays, there are several methods and processes deployed to sweeten natural gas. However, since the concentration of CO 2 and H 2 S in the raw natural gas to be processed, as well as the permitted acid gas levels in the final product may vary significantly, no single process is deemed markedly superior for all circumstances. 2,9 Over the next pages, an overview of the most commonly used processes for natural gas sweetening will be presented. 2.1.1 Liquid-phase processes Liquid-phase processes can be classified into three categories based on the nature of the liquid solvent that is being used. Namely there are chemical solvents, physical solvents and hybrid solvents that are mixtures of amines together with physical solvents. Chemical Solvent Processes In sweetening processes that use chemical solvents, absorption of the acid gasses is achieved by using alkanolamines or alkaline salts of various weak acids such as sodium and potassium salts of carbonate. Regeneration of the solvent is achieved by reducing the pressure or by application of high temperatures, which results in the desorption of the acid gases from the solvent. These processes are suitable for removing H 2 S and CO 2 from raw gas, but will still not remove organic sulfur components such as carbon disulfide. 9 Furthermore chemical solvents are specifically suitable when contaminants 4

at low partial pressures need to be removed to very low concentrations. Their use is also advantageous because of the low co-absorption of hydrocarbons. 2 Amine Processes Chemical absorption processes using aqueous amine solutions is a mature technology used in the ammonia process, steam reforming and natural gas sweetening processes. 7 Amines are compounds derived from ammonia (NH 3 ) by replacing one or more hydrogen atoms with another hydrocarbon group, usually an alkyl or aromatic group. 2 Depending on the degree of substitution of hydrogen atoms by organic groups, amines are categorized as primary, secondary or tertiary. Replacement of a single hydrogen atom results in a primary amine such as monoethanolamine (MEA) or diglycolamine (DGA). Replacement of two hydrogen atoms results in a secondary amine such as diethanolamine (DEA) or diisopropylamine (DIPA). Finally replacement of all three hydrogen atoms gives tertiary amines such as methyl-diethanolamine (MDEA). 9 The theoretical maximum CO 2 loading of primary and secondary amines is limited to approximately 0.5 mole of CO 2 per mole of amine. In the case of tertiary amines, the chemical mechanism is different allowing a theoretical maximum of 1 mole of CO 2 per mole of amine. Furthermore tertiary amines require a lower heat of regeneration. 2,7,9 Regardless of the aqueous amine solution used as the sweetening agent, the general process flow diagram for an amine sweetening plant does not change much. The sour gas will nearly always enter the plant through a scrubber to remove any free liquids and entrained solids. It enters then the absorber through its bottom and flows upward in intimate countercurrent contact with the amine solution that absorbs acid gasses. Once at the top, the sweetened gas exits the absorber and passes through an outlet separator in order for entrained amines to be recovered. Due to the presence of water in the amine solution, the sweetened gas is now saturated with water so a dehydration step is necessary before the gas is sold or fed to a cryogenic plant for LNG production. The rich amine solution leaves the bottom of the absorber containing 0.20-0.81 mole of acid gas per mole of amine. 9,11 Before entering the top of the stripper, in order to be regenerated, the amine solution goes through a flash tank to recover hydrocarbons that might have dissolved in it. In the stripper the amine comes to contact with steam, which results in the amine-co 2 separation. A stream of lean amine is removed from the bottom of the stripper and once cooled, returns to the absorber. Acid gasses leave the stripper from the top and can then be vented, incinerated, or compressed for re-injection into a suitable reservoir for enhanced oil recovery. 12 Major disadvantages of the conventional amine gas treating processes are: Large amounts of energy required for the amine regeneration Relatively low CO 2 loading capacity of amines which requires high solvent circulation rates and large high-pressure absorber columns 5

Corrosive nature of amine solutions induce high wear of the processing equipment Degradation of the amines to organic acids Co-absorption and subsequent loss of hydrocarbons Volatility of the amine solvents causes environmental pollution Potassium Carbonate Process The potassium carbonate process is another chemical solvent process for treating gas streams. It employs an aqueous solution of potassium carbonate (K 2 CO 3 ) to remove CO 2 and H 2 S. This process requires high partial pressure of CO 2 and usually a two-stage process is needed to sufficiently remove acid gasses to acceptable low levels. 9 Besides that though, the process flow steps are very similar to those of an amine process. It is a particularly useful process for removing large quantities of CO 2. Advantages of using carbonate solutions are the high chemical solubility of CO 2 in the carbonate/bicarbonate system and the low solvent cost. Main disadvantage is the fact that potassium carbonate is highly corrosive. Physical Solvent Processes In these processes, organic solvents are employed and acid gasses are removed from gas streams by physical absorption instead of chemical reaction. The high solubility of acid gasses in organic solvents (at high pressures) is the driving force for these processes. Solubility usually increases as the temperature decreases and as pressure increases. Regeneration of the solvent is later achieved either by heating or by pressure reduction. Compared to chemical solvents, the heat required to regenerate physical solvents is much less, since the heat of absorption of acid gasses for physical solvents is significantly lower. Furthermore, as most of the physical solvents absorb water, the required capacity for dehydration of the processed gas is much less compared to aqueous amine processes that saturate the sweet gas with water. Main weakness of physical solvents remain the issue of relatively low acid gas absorption capacities (at low pressures) which makes physical solvent processes competitive with amine processes only when the feed gas is available at high pressures. 13 An advantage of the use of physical solvents however, is their less corrosive nature, compared to aqueous amine solutions, whose handling requires more expensive and corrosion resistant materials. An ideal physical solvent should have a high selectivity for acid gasses, very low vapor pressure to minimize solvents losses, low viscosity, high thermal stability and should not be of corrosive nature. Although many different compounds have been commercially used as physical solvents for gas treating, none of them is markedly superior for all circumstances. 13 Selection of the 6

right solvent comes down to feed gas composition, process objectives and the special characteristics of the solvent. Physical solvents that have been commercially used for treating natural gas are: dimethyl-ether of polyethylene glycol (DEPG), propylene carbonate (PC), N-methyl-2-pyrrolidone (NMP) and methanol (MeOH). 13 DEPG is used in the SELEXOL process licensed by Dow. It is a process that can remove CO 2 simultaneously with H 2 S and water. SELEXOL together with the RECTISOL process, form the leading physical absorption technologies for treating feed gas with very high CO 2 concentrations. 2 RECTISOL is a process licensed by Lurgi GmbH that employs methanol as the physical solvent and is used to remove H 2 S, COS and bulk CO 2. NMP is employed in the PURISOL process (also licensed by Lurgi GmbH) to simultaneously remove H 2 S, CO 2, RSH and H 2 O although this particular solvent, is highly selective for H 2 S. Finally, propylene carbonate (PC) is used in the Fluor process (licensed by Fluor Daniel Inc.) to remove CO 2, H 2 S, COS, CS 2 and H 2 O from natural gas, achieving sweetening and dehydration of natural gas in one step. Hybrid Solvents Hybrid solvent processes employ mixtures of amines and physical solvents in order to take advantage of the best characteristics of both. Depending on the combination of the physical solvent and the amine that is being used, nearly complete removal of H 2 S, CO 2, and organic sulfur compounds is possible. Other advantages are higher acid gas loading, lower energy requirements for regeneration, lower corrosion rates, and lower foaming tendency. The most widely known hybrid solvent process is the Shell SULFINOL process, which applies a mixture of sulfolane, water, and diisopropanolamine (DIPA) or methyldiethanolamine (MDEA), Sulfinol-D, and Sulfinol-M, respectively. This process is used to selectively remove H 2 S, COS, RSH, and other organic sulfur compounds for pipeline specifications, while co-absorbing only part of the CO 2. 14 16 2.1.2 Dry-bed processes Dry-bed processes use a fixed bed of solid material to remove acid gases either through ionic bonding (physical adsorption) or through chemical reactions. When the bed is saturated with acid gases, it needs to be regenerated or replaced. In the case of physical adsorption, regeneration of the adsorbent is achieved by one or more simple temperature or pressure swing cycles. When the adsorbed component reacts chemically with the bed material, the process is called chemisorption and desorption is generally not possible. In this case the bed material needs to be replaced. Using porous solid adsorbents for purification of gas mixtures is a process technology used in the production of hydrogen, separation of oxygen and nitrogen from air streams and the capture of 7

odorous pollutants from industrial processes. For natural gas treating, adsorption-based separation is used to remove water, sulfur compounds, mercury and heavy hydrocarbons. 17 2.1.3 Membranes Although membrane separation technologies have been applied in the natural gas industry to remove CO 2, N 2 and H 2 S since 1980, they still account for less than 5% of the market for new natural gas processing equipment installed. 18 Various membrane technologies have been deployed in the natural gas industry. Of those technologies however, processes to capture CO 2 have been the most widely used and currently CO 2 capture is the only natural gas separation process for which membrane processes are competitive with the conventional amine technology. 18 Membrane systems for natural gas processing employ polymeric membranes with the industry standard currently being cellulose acetate. 13,18 These membranes are of the solution-diffusion type, in which a thin layer of cellulose acetate lays on top of a thicker layer of a porous support material that provides the required mechanical strength. 18 Gas separation is achieved by selective permeation of the gas constituents in contact with the membrane. High-pressure gas is fed to one side of the membrane while the other side is maintained at much lower pressure. After the gases have dissolved in the membrane material, they move across the membrane barrier under the imposed partial pressure gradient. The advantages of membrane technologies which makes them highly attractive process separation units are the ability to separate chemical species without a phase change, low thermal energy requirements, simple process flow schemes with few pieces of rotating equipment, compact plant footprints and convenient start up and shutdown procedures. These features of membranes systems are potentially attractive for remote, unmanned and footprint conscious sites. The main limitation of their use is the significant loss of hydrocarbons in the effluent stream. This issue however, can be handled by a hydrocarbon recovery stage downstream the membrane separation stage. Furthermore, while membrane systems perform well at reduced feed flow rates, their performance drops when design flow rates are exceeded. To tackle that, additional membrane modules must be added in parallel which increases the overall cost. 13 2.1.4 Cryogenic Fractionation Cryogenic fractionation involves cooling the gases to a very low temperature so that the CO 2 can be liquefied and separated. This technology requires substantial energy to achieve the low temperatures needed. Furthermore, issues arise by the formation of CO 2 solids during cryogenic 8

distillation. To overcome those problems, two technological approaches have been pursued: (1) extractive distillation by the addition of a heavier hydrocarbon to alter the solubility of components in the column (Ryan/Holmes process) and (2) controlled freezing and re-melting of the solids (Controlled Freeze ZoneTM and CryoCells processes). 2 Main advantage of this method over amine-based absorption processes, is the fact that the acid gas components are discharged as a highpressure liquid stream that can be easily pumped for geo-sequestration or for use in enhanced oil recovery operations, while yielding a high-quality methane product. 19 Furthermore, highly corrosive aqueous amine solvents are avoided and process footprint is reduced, which may be important consideration for offshore or floating production facilities. 20 2.2 Ionic Liquids Room-temperature ionic liquids (RTILs) are organic salts that melt below 100 C and have many interesting properties that make them excellent candidates for use in a variety of applications. RTILs have negligible vapor pressures; are thermally and chemically stable and non-flammable. 6,7 The negligible vapor pressures of RTILs along with their desirable gas solubility, enables them to be used for various gas separation applications. CO 2 capture from flue gas as well as for natural gas sweetening purposes is a process field where the use of ILs as potential solvents has attracted the interest of many researchers over the last decade. Primarily CO 2 solubility in ILs has been the object of most of the relevant research. The fact that the solubility and the selectivity of CO 2 in RTILs can be readily tuned by tailoring the structures of the cation and/or anion has given immense depth in the research possibilities. Figure 1 shows the most commonly used anions and cations for IL synthesis. Figure 1. Commonly used anions and cations for IL synthesis. 7 9

The solubility of CO 2 in different ILs has been widely studied by several researchers. 8,21 24 However, for gas separation processes, apart from solubility, gas selectivity is also a major factor that dictates the choice of the most suitable solvent. Although dozens of CO 2 solubility data are available in the literature, selectivity data of CO 2 in ILs are scarcely reported. In the case of natural gas sweetening processes solubility of hydrocarbons like methane and ethane in ILs should also be considered. Nevertheless research on methane solubilities in ILs is far less extensive than the available research on CO 2 solubilities. 2.3 CO 2 /CH 4 selectivity In natural gas sweetening processes, the following selectivities are mainly relevant: CO 2 /N 2, CO 2 /H 2, CO 2 /CH 4. In this part however, only work relevant to CO 2 /CH 4 selectivity is mentioned with the majority of the research focusing mainly on ideal selectivities (i.e., the ratio of pure gas solubilities). Determining real or actual selectivities in mixtures of gases cannot always be accurately done from pure gas solubilities with the assumption of ideal mixing. Measuring mixed-gas solubilities is significantly more difficult, therefore almost no mixture data can be found in the literature. Although ideal selectivity of CO 2 /CH 4 in ILs is comparable to that in conventional physical solvents, real CO 2 /CH 4 selectivity in ILs is expected to be lower than the ideal selectivity, since CH 4 solubility in many ILs increases for increasing temperature while this is the opposite for CO 2. 7 In their work, Anderson et al. 25 measured the solubility of methane, among other gasses, in the IL [hmpy][tf2n]. The solubility of these gases (298 K) decreased by the following order: SO 2 > CO 2 > C 2 H 4 > C 2 H 6 > CH 4 > O 2 > N 2. Similar gas solubility trends (i.e., CO 2 > C 2 H 4 > C 2 H 6 > CH 4 > O 2 ) were observed in the ILs [hmim][tf2n], [bmim][pf6], and [bmim][tf2n]. 25 27 In most of the ILs studied, N 2 and O 2 solubilities are much lower compared to CO 2 therefore the relevant selectivities are high enough. Hydrocarbons however, such as methane, show moderate solubilities in ILs, thereby reducing the CO 2 /hydrocarbon selectivity. Applying regular solution theory, Camper et al. 28 33 showed that the physical solubility of gases in ILs was well correlated with the liquid molar volume of the IL and that ideal selectivities for CO 2 /N 2 and CO 2 /CH 4 should increase as the molar volume of the IL decrease. Finotello et al. 33 measured CO 2, N 2, and CH 4 solubilities in [emim][tf2n],[emim][bf4], [hmim][tf2n] and [mmim][meso4]. Their results show that, as temperature increases, the solubility of CO 2 decreases in all RTILs, the solubility of CH 4 remains constant in [emim][tf2n] and [hmim][tf2n] but increases in [mmim][meso 4 ] and [emim][bf4]. Also, CO 2 /CH 4 ideal solubility selectivity increases as temperature decreases. 10

Kumelan et al. 34 measured the solubility of methane and of xenon in the ionic liquid 1-n-butyl- 3-methylimidazolium methyl sulfate ([bmim][ch 3 SO 4 ]) with a high-pressure view-cell technique based on the synthetic method. Among his findings is that CH 4 becomes less soluble in [bmim][ch 3 SO 4 ] with rising temperature. Jacquemin et al. 35 reported experimental values for the solubility of carbon dioxide, ethane, methane, oxygen, nitrogen, hydrogen, argon and carbon monoxide in 1-butyl-3-methylimidazolium tetrafluoroborate, [bmim][bf4]. They conclude that carbon dioxide is the most soluble gas with mole fraction solubilities of the order of 10-2. Ethane and methane are one order of magnitude more soluble than the other five gases that have mole fraction solubilities of the order of 10-4. In their work, Carvalho and Coutinho 36 measured CH 4 solubilities in imidazolium, phosphonium, and ammonium ILs. Among the findings was the fact that an increase in temperature had a small or even negligible impact on CH 4 solubility. Furthermore, the solubility of CH 4 was shown to relate to the polarity of the IL. Althulith et al. 37 present experimental measurements of the CH 4 solubility in [emim][fap]. Which is then compared to solubilities in other ILs. This comparison shows that the solubility of CH4 in the various ILs slightly decreases in the order: [hmim][tf2n] > [emim][fap] > [bmim][tf2n]. Furthermore solubilities of CH 4 are much lower compared with CO 2 solubilities, thus making the relevant ILs suitable for separating CO 2 from natural gas. Bara et al. 38 measured the solubility and ideal selectivities of the gas pairs CO 2 /N 2 and CO 2 /CH 4 in imidazolium-based ILs functionalized with oligo(ethylene glycol). They showed that the CO 2 solubility in these oligo(ethylene glycol) functionalized ILs were similar to their corresponding alkyl analogues, but N 2 and CH 4 solubilities were lower corresponding to a higher ideal selectivity for the two gas pairs. Similar results were reported by Carlisle et al. 39 for nitrile-functionalized ILs. The nitrile-functionalized ILs exhibited lower CO 2, N 2, and CH 4 solubilities, but showed improved CO 2 /N 2 and CO 2 /CH 4 selectivities compared to alkyl-substituted analogues. 11

3. Experimental Procedure 3.1 Phase Behavior Theory According to Gibbs phase rule for non-reacting systems (eq. 3.1), the number of variables that may be independently fixed in a system at equilibrium, is the difference between the total number of variables that characterize the intensive state of the system and the number of independent equations that can be written connecting those variables. 40 This difference gives the degrees of freedom (F) of the systems. For a system containing N chemical species and π phases, F is evaluated: = 2 + (3.1) A unary system (N = 1) must have at least one phase (π = 1) so the maximum degrees of freedom are two. Therefore a single P-T diagram can describe phase behavior of pure substances. A typical P-T diagram of a pure substance is shown in Figure 2. At the triple point, the unary system (N = 1), with three phases (π = 3) reduces to a single point. 40 Figure 2 P-T diagram of a pure substance For a mixture of two components the degrees of freedom increase. A binary mixture (N = 2) with at least one phase (π = 1) has a maximum of three degrees of freedom (F), namely: pressure, temperature and mole fraction. These three dimensions can be projected in a three-dimensional space P-T-x projection. 12

3.2 Experimental method Experiments were carried out in a Cailletet apparatus, schematically shown in Figure 3. This apparatus allows the measurement of phase equilibrium according to the synthetic method. 41 In this method, a mixture of fixed overall composition is injected into an equilibrium cell, which is then fully sealed off throughout the whole experiment. That way, phase behavior is constrained to two degrees of freedom, namely temperature and pressure. For each data point measured, the temperature of the sample remains fixed at a set value and the pressure is slowly varied until phase change is observed visually through the Cailletet transparent glass tube. Solubility experiments were executed on binary systems of ionic liquids plus gas and pseudo-binary systems of ionic liquid plus a gas mixture of CO 2 and CH 4. In both cases bubble-point pressures were determined as the pressures at which the last bubble of gas disappears in the liquid. Figure 3. Schematic representation of the Cailletet apparatus: A, autoclave; B, magnets; C, capillary glass tube; D, drain; E, motor; F, metal stirrer; G, platinum resistance thermometer; H, rotating hand pump; Hg, mercury; I, thermostat liquid in; L, line to dead weight pressure gauge; M, mixture being investigated; Ma, manometers; O, thermostat liquid out; Or, hydraulic oil reservoir; P, closing plug; R, Viton-O-rings; S, silicone rubber stopper; T, mercury trap; Th, glass thermostat; V, valve. 41 13

The ionic liquids used are highly hygroscopic. It is therefore essential that the amount of water dissolved in the IL samples is as low as possible. In practice however, IL samples are considered acceptable if water content is below the limit of 1200 ppm. To achieve that, the samples would have to spend at least 24 hours at a temperature of 80 C in the vacuum oven. After that period, the exact water content was measured using the Karl Fischer titration method. If the water content is within acceptable limits, then the following steps are performed: An empty Cailletet tube (equilibrium cell) is weighed with an accuracy of 10-4 grams. Using a thoroughly washed and dried syringe, an arbitrary amount of ionic liquid is injected in the tube, which is then weighed again to calculate the exact mass of the IL sample (typically between 100 and 200 mg with an accuracy of 10-4 grams) Once the magnetic stirrer ball is inserted in the tube, the tube gets connected to a vessel of calibrated volume at the gas rack. (Figure 4) The air and water in the sample are evacuated under high vacuum while the IL sample is kept frozen by repeatedly being immersed in liquid nitrogen. The moles of gas in the calibrated vessel are calculated using the virial EoS. (Low pressure ensures insignificant error by the use of the virial EoS 40 ) By adjusting the pressure of the gas or gas mixture in the calibrated volume vessel, the desired mass percentage of gas in the final mixture is achieved. The gas rack used for filling the Cailletet tube has an accuracy of ±0.1 mbar which leads to an uncertainty in the composition of u(x)=±0.001. Figure 4. Schematic representation of the gas rack. 42 14

The gas partial pressure needed to achieve the desired gas mass percentage in the final sample is calculated using the virial EoS (truncated to two terms for application at low pressures 40 ): = () (3.1) Where: T : Temperature of the gas in the calibrated vessel R : Gas constant B(T) : Gas or gas-mixture second virial coefficient, which is only temperature dependent. 40 (Detailed calculation of the second virial coefficients Appendix A) The molar volume of the gas or gas mixture is calculated: = "##"$ Where n are the moles of gas needed to achieve the desired mass fraction of gas in the sample. (3.2) Once the desired mass percentage of gas is achieved, the tube is sealed with mercury and then transferred to the autoclave. The autoclave is connected to a hydraulic oil system in which pressure is generated by means of a screw-type hand pump. Pressure measurements are done using a dead weight pressure gauge with an accuracy of 0.05 bar. 41 The temperature of the sample is kept constant at a set value by means of a thermostat liquid, which is circulated through a glass jacket that surrounds the Cailletet tube. A thermostat bath is used to maintain the temperature of the thermostat liquid to the desired value with an accuracy of 0.01K 41. The temperature of the thermostat liquid is recorded with an accuracy of 0.02 K using a platinum resistance thermometer. Finally proper agitation of the sample is achieved using a stainless steel ball that is inserted in the tube along with the sample. The ball is forced to move up and down by two reciprocating magnets. 15

3.3 Measurements and corrections Once the Cailletet tube is connected to the autoclave, it gets covered with the glass jacket. The jacket is then connected to the thermostat-bath and the flowing water within it, regulates the sample temperature. After the sample reaches the desired temperature of the measurement, the pressure is slowly varied using the dead-weight gauge. Bubble-point pressure for a specific temperature is then determined as the pressure at which the last bubble of gas disappears. For the measured bubble-point pressure to be correct, certain corrections need to be applied: Atmospheric pressure correction: The dead weight gauges always measure relative to the atmospheric pressure. Therefore a barometer is employed with an accuracy of 0.1 mbar. Gravity correction for the laboratory latitude (52 ):,"##$%$& = "#$%&"' 9.81247 9.80665 = "#$%&"' 1.00059 (3.3) Temperature correction for the influence of the temperature ( C)%on the piston area:,"##$%$& =,"##$%$ [1 20 2.3 10 ] (3.4) Pressure transmitting fluid correction. In this case the pressure from the mercury column is calculated using the temperature dependent density of mercury: (3.5) " = 12.281 + 13595 " = " h The actual experimental pressure is finally calculated: "#$% =,"##$%$& + "# " 16

3.2 Ionic liquids used Experiments were performed with the following ILs: 1-Ethyl-3-methylimidazolium diethyl phosphate ([emim]-[phosphate]) Trihexyl-tetradecyl-phosphonium-dicyanamide([thtdp]-[dca]) Trihexyl-tetradecyl-phosphonium-bis(2,4,4-trimethyl-pentyl)-phosphinate ([thtdp][phosphinate]) The ILs were purchased from SIGMA ALDRICH. In the table that follows are the information provided by the company for these ILs. In Figures 5 and 6, the chemical structures of the ILs are presented. Thtdpdicyanamide Thtdpphosphinate Emimphosphate Empirical Formula (Hill Notation) C 34 H 68 N 3 P C 48 H 102 O 2 P 2 C 10 H 21 N 2 O 4 P Molecular Weight (g/mole) 549.90 773.27 264.26 Density (g/ml at 20 C) 0.90 0.895 - Assay 95.0% 95.0% 98.0% Table 2. Ionic liquid properties 43 Figure 5. Chemical structures of [emim]-[phosphate] (left) and [thtdp]-[dicyanamide] (right). Figure 6. Chemical Structure of [thtdp]-[phosphinate]. 17

4. Modeling the phase behavior of binary systems 4.1 Peng-Robinson Equation of State Using equations of state (EoS) is the most common method for phase equilibria correlation in mixtures at high and low pressure. Cubic EoS, derived from the van der Waals EoS, are the most common and industrially important EoS. Among them, the Peng Robinson EoS has proven to combine the simplicity and accuracy required for the prediction and correlation of volumetric and thermodynamic properties of fluids. 44,45 The original Peng-Robinson EoS 46: = + + ( ) (4.1) Applying this equation at the critical point yields: ( ) = 0.45724 ( ) = 0.07780 (4.2) At temperatures other than the critical: = (, ) = ( ) (4.3) Where (, ), is a dimensionless function of reduced temperature ( ) and acentric factor (ω) that equals unity at the critical temperature., = 1 + 1. ( (4.4) = 0.37464 1.54226 0.26992 The extension of the PR EoS to mixtures, requires the use of mixing rules45: "# = " "# = " (4.5) Where " = 1 " = " " = 1 ( + )(1 " ) 2 = = 0 (4.6) 18

, are the molar fractions of the components of the binary system and, are the binary interaction parameters. To make the required calculations, numerical values of the critical temperature and pressure (, ) as well as molecular weight and acentric factors (ω) are needed for CO 2, CH 4 and the ILs While these properties are well known for common substances such as carbon dioxide and methane, they are not readily available for ILs. One reason for this is the fact that most of them start to decompose at low temperatures and in some cases even at temperatures close to their boiling point. As a result no experimental data exists and these properties must be estimated. Group contribution methods correlate structural molecular properties with mathematical functions representing a chemical property of a molecule. A group contribution method expresses the thermodynamic property of a chemical compound as a function of the sum of contributions of smaller groups of atoms constituting the molecule. A large variety of group contribution methods have been proposed in the past years, differing in their field of applicability and in the set of experimental data they are based on. Since the first group contribution methods were developed by Riedel in 1949 and Lydersen in 1955, a large number of methods have been studied to achieve the most reliable results. Valderrama et al. 47 combined the best results of the Lydersen method with the best results of the Joback-Reid method to propose a modified Lydersen-Joback-Reid method. This method gives good results for molecules of high molecular weight such as ILs. In Table 3, critical properties and acentric factors of CO 2 andch 4 are reported. Table 4 reports critical properties and acentric factors for the ILs Thtdp-dca and Thtdp-phosphinate. These values were calculated using group contribution methods in the MSc thesis of T. Olasagasti 42 Methane (CH 4 ) Carbon Dioxide (CO 2 ) Molecular Weight (g/mole) 16.043 44.01 Critical Temperature (T c ) 190.6 K 304.1K Critical Pressure (P c ) 4.599 MPa 7.376 MPa Acentric factor (ω) 0.012 0.239 Table 3. Critical properties and acentric factors of CO 2 andch 4. Thtdp-[dca] Thtdp-phosphinate Empirical Formula C 34 H 68 N 3 P C 48 H 102 O 2 P 2 (Hill Notation) Critical Temperature (T c ) 1525.5K 1878.8K Critical Pressure (P c ) 0.765 MPa 0.551 MPa Acentric factor (ω) 0.5822-0.13188 Table 4. Critical properties and acentric factors of ILs [thtdp]-[dca] and [thtdp]-[phosphinate]. 19

4.2 Vapor-liquid Equilibria (VLE) When thermodynamics is applied to vapor/liquid equilibrium, the goal is to find by calculations, the temperatures, pressures and compositions of phases in equilibrium. To achieve that, appropriate models for the behavior of the system in vapor-liquid equilibrium are needed. The two simplest models are Raoult s and Henry s law 40. The mathematical expression for Raoult s law is: = "# ( = 1,2,, ) (4.8) In which% and are respectively the liquid and vapor phase mole fraction, "# the saturation pressure of the pure component and the total pressure on the system. Two major assumptions are required in order to reduce VLE calculations to Raoult s law. Namely that the vapor phase is an ideal gas and that the liquid phase is an ideal solution. The first assumption means that Raoult s law only applies for low to moderate pressures and the second assumption implies that it can have approximate validity only when the species that comprise the system are chemically similar. It is however valid for any species present at a mole fraction approaching unity, provided only that the vapor phase is an ideal gas 40. Application of Raoult s law to species%requires a value for "# at the temperature of application, and thus is not appropriate for a species whose critical temperature is less than the temperature of application. In this case, Henry s law is the appropriate VLE model which in essence defines the solubility of the supercritical gas in the solvent. For a species present as a very dilute solute in the liquid phase, Henry s law states that the partial pressure of the species in the vapor phase is directly proportional to its liquid phase mole fraction. Thus, = (4.9) Where is the Henry s constant. Values of come from experiments. Ideal solution behavior is described by the Lewis/Randal rule: " = (4.10) The solid line in Figure 7 represents experimental values of. At = 1 the line becomes tangent to the Lewis/Randal rule. In the other limit, 0 becomes zero. The ratio / is therefore indeterminate in this limit, and application of l Hôpital s rule yields: 20

lim = = (4.11) Equation 4.11 defines Henry s constant, as the limiting slope of the curve at = 0. The equation of this tangent line expresses Henry s Law: = (4.12) of. 40 Which is applicable in the limit as 0, but is also of approximate validity for small values Figure 7. Composition dependence of liquid-phase fugacities for species i in a binary mixture. 21

4.3 Matlab The implemented code, calculates the bubble point of a mixture using the fundamental relation of any phase equilibrium calculations; the equality of fugacities of each species in each phase, that is: = = (4.7) This condition is solved using the algorithm presented in the following figure: Figure 8. Bubble point pressure algorithm for specified temperature and liquid composition. EoS. Fugacity coefficients for both the liquid and gas phase are calculated using the Peng-Robinson 22

5. Results 5.1 Binary systems In this section the experimental results of the bubble-point pressure measurements will be reported for the three binary systems that were studied. First P-T diagrams of the experimental data are presented followed by P-x diagrams that include both experimental and modeling results. The solubilities of CH 4 and CO 2 in each studied IL are then compared. The experimental data on the systems [thtdp]-[dca]+co 2 and [thtdp]-[phosphinate]+ CO 2 were available in the MSc Thesis work of T.Z. Olasagasti. 42 5.1.1 Thtdp-dicyanamide The solubility of CH 4 in the IL [thtdp]-[dca], was determined at temperatures ranging from 302.13 to 363.48 K and pressures up to 11.57 MPa by measuring the bubble-point pressures at various compositions of CH 4 in the IL. In Figure 9, the pressure in plotted versus the temperature for this system. It can be concluded that the bubble-point pressures of CH 4 in the IL increases almost linearly with temperature, indicating that the solubility decreases with increasing temperature. 15 1.5wt%CH4 P"[MPa]" 10 5 1.25wtCH4 1wt%CH4 0.75wt%CH4 0.5wt%CH4 0.35wt%CH4 0.25wt%CH4 0 290 305 320 335 350 365 380 395 T"[K]" Figure 9. P-T diagram of the system thtdp-dicyanamide + CH 4. The lines plotted are meant as a guide to the eye. Among the three different ILs that were used for solubility measurements, [thtdp]-[dca] was the less viscous. Due to the relatively low viscosity, the magnetic stirrer was able to operate at higher speeds thus helping the system to reach an equilibrium state faster after small changes in pressure. 23