Ionic Liquid as extractant in Liquid Liquid Extraction

Transcription

1 Chapter 2 Ionic Liquid as extractant in Liquid Liquid Extraction

2 2.1 Liquid-liquid extraction It s a simple procedure in which, different types of solute dissolved in one liquid phase, is preferentially removed or extracted, by using an immiscible liquid or solvent in which, one of the solute has high affinity in comparison to others, for the these immiscible solvent used for extraction. It is also known as solvent extraction. In such operation, the solution which is to be extracted is called feed and the liquid with which the feed is contacted is the solvent. These forms biphasic layer, in which solvent rich phase is known as extract and the residual liquid remains after the extraction is known as raffinate. More complicated system may use two solvent to separate the components of a feed. Through put from 100,000 m 3 /h or/and even higher can be treated with extractor of reasonable size. Although energy consumption for the normal extraction process itself is almost negligible, the attached steps for the recovery of solvent require more or less energy, depending on the nature of the components and the difficulty of separation. Often, not only the extract phase but also the raffinate phase has to be processed by washing, distillation, or another follow-up treatment. The complete extraction process with solvent regeneration and raffinate treatment needs a quite complex plant with the corresponding investment cost. Technically, Extraction or separation of dissolved chemical component X from liquid phase A is accomplished by bringing the liquid solution of X into contact with a second phase, B, given that phases A and B are immiscible. Phase B may be a solid, liquid, gas, or supercritical fluid. In Liquid Liquid extraction, phase B is always a Liquid. A distribution of the component between the immiscible phases occurs. After the analyte is distributed between the two phases, the extracted analyte is released and/ or recovered from phase B for subsequent extraction procedures or for instrumental analysis. The theory of chemical equilibrium, the reversible distribution reaction as X X (2.1) A B And the equilibrium constant expression referred to as the Nernst distribution law is, 51

3 K X X B (2.2) D A Where, the brackets denote the concentration of X in each phase at constant temperature (or the activity of X for non ideal solution). By convention, the concentration extracted into phase B appears in the numerator of equation 2.2. The equilibrium constant is independent of the rate at which it is achieved. The analyst s function is to optimise extracting conditions so that the distribution of solute between phases lies far to the right in Equation 2.1, and the resulting value of K D is large, indicating a high degree of extraction from phase A into phase B. Conversely, if K D is small, less chemical X is transferred from phase A into phase B. If K D is equal to 1, equivalent concentration exist in each phase. More solvent combination are miscible than immiscible, and more solvents are immiscible with water than with any other solvent. Solvents miscible with water in all proportion include acetone, acetonitrile, dimethyl acetamide, N, N-dimethylformamide, dimethyl sulphoxide, 1, 4-dioxane, ethyl alcohol, glyme, isopropyl alcohol, methanol, 2-metoxyethanol, N- methylpyrrolidone, tetrahydrofuran, and trifluoroacetic acid. Another consideration when selecting an extraction solvent is its density. This will help us to determine the position of the layer. Solubility of the solvent also plays a major role in extraction. Solvents may form two visibly distinct phase when mixed together, they are often somewhat soluble in each other and will, in fact, become mutually saturated when mixed with each other. As in many separation processes, the pressure and temperature conditions play a large role in the effectiveness of the separation. In order for a good split of the feed the pressure and temperature must be such so as to ensure that all components remains in the liquid phase. The process will be adversely affected if one or more of the component is allowed to become a vapour, or the extraction may not occur at all if large enough portions of components are allowed to vaporise. In addition, the temperature should be high enough that the components are all soluble with one another. If extremes in temperature are present, finding a suitable solvent for extraction can be problematic. This is however, generally not the case, since one of the biggest in the extraction process is that it can be done at ambient pressure and temperatures. 52

4 In many applications, a separation process is desired where an extreme temperature will destroy the desired products such as the pharmaceutical industry. For these applications, extraction is ideally suited, since the only temperature requirement is that dictated by the solubility. At this point the biggest challenges would be finding a suitable solvent for extraction. Temperature plays a smaller role in extraction than in other separation process. It is only dependent upon the temperature of the streams fed in to the column. There is not a heating requirement for the process and Heat of enthalpy is generally insignificant. For these reason, extraction can be considered as an isothermal process. Pressure also plays only a small role in extraction. When combined with the temperature consideration it is only necessary that the mixture remain in the two-phase liquid region. The fact that extraction process can be run at isothermal and isobaric conditions is quite beneficial to the phase stability of the system. Phase stability from a thermodynamic standpoint is temperature and pressure dependent and since these are not changing the stability of the phase will not change. Activity coefficients are the most important physical property in the extraction process. The reason for this is that these are used to determine the miscibility of the solute in both of the solvents involved. While there are many different equations available to determine a particular activity some are better than others for extraction purpose. When working with liquid liquid systems the NRTL and the UNIFAC models are the most accurate in predicting the activities of the liquids involved according to Medersema et al. (2008). Although better than such predictive models such as Van Laar or Margules they still fall short of perfection. Once a predictive model has been plotted on a diagram it will most likely be necessary to fix the exact equilibrium line experimentally for the most accurate data. The activity coefficients also determine the partition factor which will determine whether or not a good separation is possible. Viscosity is a property that cannot be overlooked, its presence in two different areas, flooding and choice of equipments. Flooding is a phenomenon that can occur in extraction just as it can for other unit operation. Viscosity is also valuable in the determination type of system to use for extraction. Components having a high viscosity cannot be used in spray or packed columns. 53

5 2.1.1 Limitations One must consider the extremes conditions which can be used as separation process 1. Suitable solvent Solvent, partially soluble with the carrier used. Feed components immiscible with the solvent. Solute is soluble in the carrier and at the same time completely or partially soluble in the solvent. Different densities than the feed components for a phase separation to facilitate and maintain the capacity of the extractor is high. Extremely high selectivity, for the solute for the solvent to dissolve the maximum amount, of solute and the minimum amount of the carrier. Large distribution coefficient to reduce the theoretical number of stages. Low viscosity increases the capacity of the extraction column and does not allow for the settling rate of dispersion to be slow. Chemically stable and inert toward other component of the systems. Low cost, non toxic and non-flammable 2. Equipment Interfacial tension and viscosity High interfacial tension and viscosity leads to more power being supplied to maintain rapid mass transfer throughout the extraction process. Low interfacial tension and viscosity leads to the formation of an emulsion. 3. Temperature preferred to be higher since solubility increases, but temperature not higher than the critical solution temperature 4. Pressure for condensed systems must be maintained below the vapour pressure of the solution such that a vapour phase will not appear and interrupt liquid equilibrium 5. Separation may only occur for composition in the region between the feed composition and that apex of the carrier used. 54

6 2.2 What is Naphtha cracker? Cracking is the process whereby complex organic molecules such as kerogens or heavy hydrocarbon are broken down into simpler molecules such as light hydrocarbons by the breaking of carbon-carbon bonds in the precursor. The rate of cracking and the end products are strongly dependent on the temperature and presence of catalysts. Cracking is the breakdown of large alkane into smaller, more useful alkanes and alkenes. The thermal cracking method (also known as Shukhov cracking process ) was invented by Russian engineer Vladimir Shukhov and patented in 1891 in the Russia Empire, patent no , Nov. 27; 1891.This process was modified by American engineer William Merriam Burton and patented as U.S. patent on June 8, In 1942, the delegation of the American Sinclair Oil Corporation paid a visit to Shukhov. Sinclair Oil protested the personal right appropriated by the Rockefeller Standard Oil concern on the discovery of oil cracking. It indicated that Burton s patent used by the Standard Oil concern was the modified patent of Shukhov proved to the Americans that the Burton s method was just the slightly changed modification of his 1891 patent. Number of flammable liquid mixture of hydrocarbon i.e., a component of natural gas condensate or a distillation product form petroleum, coal tar or peat boiling in a certain range and containing certain hydrocarbon are normally refers to as Naphtha. It contains volatile aromatic similar to gasoline; its colour differs from colourless to reddish brown. Naphtha is a generic term applied to refined, partly refined or unrefined petroleum products and liquid products of natural gas which distil below 240 O C; the volatile fraction of the petroleum, which is used as a solvent or as a precursor to gasoline. In fact, not less than 10% of material should distil below 75 o C ; not less than 95% of the material should distil below 240 o C under standard distillation conditions by James G. Handbook of petroleum product analysis, by James G Speight (John Wiley & sons Inc, New Jersey, USA ) Naphtha resembles gasoline in terms of boiling range and carbon number, being a precursor to gasoline. Naphtha is used as automotive fuel, engine fuel, and jet-b. Broadly, naphtha is 55

7 classified as Light Naphtha and Heavy Naphtha. Light naphtha is used as rubber solvent, lacquer diluents, while heavy naphtha finds its application as varnish solvent, dyer naphtha, and cleaner naphtha. Volatility, solvent properties purity and odour determine the suitability of naphtha for a particular use. The use of naphtha as an incendiary get in warfare, and as an illuminant dates back to 1200 AD. Naphtha is characterised as lean (high paraffin content) or rich (low paraffin content). The rich naphtha with higher proportion of naphthene content is easier to process in the platforming unit. A rich naphthene produces greater volumetric yield of reformate than does a lean charge. The wide range of naphtha available from ordinary paraffinic straight run to highly aromatic type and the varying degree of volatility offer products suitable for many uses. Naphtha solvents may belong to categories such as, special boiling spirits having distillation range o C, White spirit boiling within o C, and high boiling petroleum fractions. In aromatic complexes, naphtha is converted into basic petrochemical intermediates; Benzene, toluene and xylene (BTX). Petroleum naphtha is by far most popular feedstock for aromatic production. Most ethylene cracker feeds contain 10-25% of aromatic components, depending on the source of the feed. The aromatic compounds are not converted to olefins and even small amounts are formed during the cracking process in the cracker furnaces stated by Zimmermann et al. (2005) therefore, they occupy a part of the capacity of the furnaces and they put an extra load on the separation section of the C 5 aliphatic compounds. If a major part of the aromatic compound present in the feed to the cracker could be separated up stream of the furnaces, it would offer several advantages: higher capacity, higher thermal efficiency and less fouling. The current processes for the separation of aromatic and aliphatic and aliphatic hydrocarbon are either suitable for aromatic concentration of 20% or more, such as extraction, extractive distillation and azeotropic distillation, or show low aromatic/ aliphatic selectivity and/or have low capacities, Therefore it was whether and how the performance of adsorption, membrane separation and extraction could be drastically improved at low aromatics content by using new separating agents. 56

8 Figure 2.1: Flow diagram for Naphtha cracking, Medersema (2005) Figure 2.1 shows a flow scheme of naphtha cracker. Firstly metal contaminants in the feed are hydrogenated and removed by molecular sieves. The remaining feed stream is heated up to 60 o C and enters the furnace section where aliphatic hydrocarbon are converted by the high temperature ( o C) into smaller molecules, preferably ethylene and propylene. After the furnace section the product stream had to be cooled down in several steps. Bold lines in Figure 2.1 illustrate aromatic hydrocarbon flow. Presence of aromatic compounds in the feed to the cracker also has a negative influence on thermal efficiency. Because aliphatic hydrocarbon form azeotropes with aromatic compounds present, it become necessary to remove aromatic compounds before it further processes in cracker. Some polar solvents such as sulpholane, NMP (N-methyl pyrrolidine), ethylene glycol, N-formyl morpholine (NFM) involves additional cost and energy in the process. For an economic feasibility operation, the amount of aliphatic compounds in the aromatic product stream should be as low as possible. On the other hand, total removal of the aromatic compound from the feed is not necessary because aromatics are formed during the cracking 57

9 process. Therefore, the recovery of the aromatics hydrocarbon is set to be at least 98% and the purity of the generated hydrocarbon is also set to be at least 98%. If the aromatic hydrocarbon can be removed from the feed to the naphtha cracker, the energy efficiency of the cracker will be improved. Since not the complete cracker will be affected by the removal of the aromatic hydrocarbon, an energy saving of 5-10% is expected. 58

10 2.3 Various processes used for separation in naphtha cracker The separation of aromatic hydrocarbon such as benzene, toluene, ethylbenzene and xylenes, i.e. from C 4 TO C 10 aliphatic hydrocarbon mixture is challenging since these hydrocarbon have boiling points in a close range and several combinations from azeotropes. The separation of benzene and cyclohexane is one of the most important and most difficult processes in the chemical industry. Since benzene and cyclohexane form close boiling point mixture at the entire range of their composition, the separation is difficult by means of a standard distillation process. Since distillation is not feasible option for the separation of aromatic hydrocarbon from C 4 - C 10 aliphatic due to their boiling in a close range and the formation of several azeotropes, separation process based on differences in interaction or affinity must be used. Examples of few processes are extraction, adsorption, membrane permeation, extractive distillation. For each process option, the criterion with regard must be fixed. The capacity in an extraction process is determined by the distribution coefficient of the aromatics and the capacity of a membrane process is determined by the flux rate through the membrane. The benchmark for the selectivity (> 30) is the selectivity obtained in the sulpholane extraction process, which is the most used process for separating aromatic and aliphatic hydrocarbons. The recovery of the aromatics needs not to be complete because of the formation of aromatics during the cracking process. Therefore recovery of the aromatics needs not to be complete because of the formation of aromatics during the cracking process. Therefore, the value of 98% is an arbitrary value, just as the requirements of both product streams. For process like extraction and extractive distillation, the investment cost scale with the capacity to the power while the adsorption / desorption and membrane processes are scaled up with a power of The higher scale up factor for adsorption /desorption and membrane processes is caused by the fact that the capacity of these processes is directly related to the amount of adsorbent or membrane area. Generally it can be stated that adsorption and membrane process are most suitable for small scale processes than for large scale ones. Unless the unit price of the adsorbent or membrane is very low. 59

11 2.3.1 Membrane permeation According to Garcia et al. (2000) for membrane separation of aromatics and aliphatic hydrocarbon a large number of polymeric membrane are possible. Some of the studied systems are the separation of benzene and cyclohexane or toluene and n-octane. In recent years, pervaporation separation, and to a lesser extent also vapour permeation, have emerged as relatively simple alternatives to many water/organic and organic/water separation applications, Wang et al. (2001) and Kita et al. (2001), Pervaporation and vapour permeation are especially attractive in azeotropic and close boiling point separation applications, since these processes are not based on the relatives volatilities of components, but on the difference in sorption and diffusion properties of the feed substance as well as the permselectivity of the membrane. If pervaporation or vapour permeation with high selectivity membrane produces permeates and retentate streams with a high purity, no additional separation processes are required. On the other hand, application of pervaporation or vapour permeation in organic separation with organic membrane is still very limited because of their stability, a low selectivity and/ or low flux rates. Polymeric membrane shows low flux rate, and have limited thermal stability, especially in the presence of organic solvents. Therefore, use of polymeric membranes for organic/organic separation in rare in industrial processes. For these reasons, polymeric membrane will not be considered for this separation. In comparison to polymeric membranes, some zeolite, like faujasite with a low Si/Al ratio, exhibit relatively high separation factors varying from 20 up to 260 for the separation of benzene/hexane and benzene/ cyclohexane, depending on the temperature. For organic/organic separation, zeolite membranes are to be preferred because of their robustness in organic solvents and their application at higher temperatures. Another advantage of using zeolite membrane is that these can be easily regenerated by calcinations. Desorption in zeolite membrane is not difficult as in But the application of a zeolite membrane installation will only become economically feasible if the membrane module cost price is reduced and higher aromatic fluxes are obtained. Therefore, zeolite membrane must have higher flux by reducing the zeolite layer 60

12 thickness of the membrane, while maintaining selectivity and technologies to achieve reduction in the manufacturing costs of zeolite membrane modules Adsorption A suitable adsorbent for the separation of aromatic and aliphatic and hydrocarbon from a process stream such as naphtha, which contains 10-25% aromatics, must absorb the aromatic hydrocarbon. Most zeolite, because of the presence of the exchangeable cations, is polar adsorbent. Molecules such as water or ammonia, CO 2, N 2 or aromatic hydrocarbon are therefore adsorbed more strongly than non-polar species of comparable molecular weight was stated by Ruthven (1998). Suitable zeolite for the adsorption of aromatic compounds have a low Si/Al rario, for instance X and Y types. However, since aromatic hydrocarbon are srongly adsorbed on zeolite, desorption is a difficult step. Desortion by pressure sing is difficult when the process streams are in the fluid phase and with thermal swing desorption the activity of the zeolite decreases with timw. Steam stripping is not suitable because zeolite with a low Si/Al ratio quickly loose their stability in contact with water or steam. Suitbale compounds for desorption by displacement are polar or polarisable compounds, which musst have a large interaction with the zeolite than the aromatic compounds, for instance alcohols, like methanol, ethanol, proponal or propylene glycol. A disadvantages is, of course, that a second separation step is required to remove the dispalcement liquid. Other alternative way of desorbing benzene from a NaX zeolite is described by Bellows et al. (1993) i.e by hydrogenation of benzene to cyclohexne with a feed containing a mixture of aromatic and aliphatic hydrocarbons, the aromatic hydrocrbon can be hydrogenated and be desorbed in a similar way. The disavantages of this method is that the aromatic compounds are converted to cyclo alkanes and that they are not avilable as feedstock for other compounds. However due to the foreseen difficulties with desorption in comparison with zeolite membrane, the adsorption/ desorption option was not pursued further. 61

13 2.3.3 Extraction and extractive distillation The conventional processes for the separation of aromatic and aliphatic hydrocarbon are liquid extraction, suitable for the range of 20-65% aromatic content, extractive distillations, for the range of 65-90% aromatics, and azeotropic distillation for high aromatic content, > 90%. For the extraction of aromatic hydrocarbon, a solvent system must be designed. The solvents selected must preferably extract the aromatic hydrocarbon from the feed, since these are present in a relatively low concentration. For solvent selection, the following are of importance in extraction processes; Group selectivity must be high Selectivity usually decreases with increasing temperature Selectivity can be increased by addition of an anti solvent Selectivity decreases as more and more hydrocarbon are dissolved in solvent. These are several commercial extraction processes available for the selective separation of aromatics from naphtha or gas condensates. Typical solvents used are polar components such as sulpholane used by Chen et al. (2000), N-methyl pyrrolidone (NMP) by Krishna et al. (1987), N-formyl morpholine (NFM) by (Krupp Udhe, Snamprogetti), ethylene glycols by Yorulmaz et al. (1985), propylene carbonate by Ali et al.(2003), furfural (Texaco) or DMSO (IFP).Where as Hamid et al. (1996) describe the usual solvent feed ration is 1.5-4, which means that a large amount of solvent will be required for the feed to the ethylene crackers. The raffinate stream from the extractor contains aliphatic hydrocarbons and a small amount of the solvent, which can be washed out by water. The aliphatic product stream can then be processed further. The extract contains aromatic and some light paraffins, which can be recovered and recycled to the extraction column as backwash. In commercial processes, solvent recovery can be carried out by Flashing or evaporation Flashing followed by stripping Extractive stripping followed by distillation. 62

14 No extraction process is in operation for removing aromatic compounds anywhere. The only way to improve the economic feasibility of extraction technology at low aromatics content is the development of new solvents systems that exhibit a dramatically higher aromatic distribution coefficient and / or a higher aromatic/aliphatic selectivity than sulpholane. Literature information indicated that aqueous cyclodextrin solution and ionic liquid could have the potential to fulfil these requirements. The separation of aromatic compounds from aliphatic hydrocarbons with aqueous solution of cyclodextrin (CD) is an option, because cyclodextrin can incorporate several organic compounds and the separation of the aqueous solution of complexed cyclodextrin derivatives are highly soluble in water by Szejtli (1998) Therefore, aqueous cyclodextrin solution could be used as extraction solvents for the separation of aromatic compounds from cracker feeds stated by Uemasu et al. (2004). Room temperature ionic liquid (RTIL) are liquids that are composed entirely of ions and in this sense alone resemble the ionic melt which may be produced by heating normal metallic salts such as sodium chloride to high temperature (e.g. NaCl is liquid above 800 o C). In fact, ionic liquid can now be produced which remain liquid at room temperature and below (even as low as -96 o C). Ionic liquid possess of a number of properties, which may be of importance in their application as extractive media in liquid/liquid extraction processes. They are liquid at room temperature and, in fact, they have an enormous liquid temperature range of 300 o C, which is large than that of water (100 o C) and this offers the potential for considerable kinetic control of extractive processes. They are good solvents for a wide range of inorganic, organic and polymeric materials. The composition of ionic liquid may be adjusted enabling control of their acidity or basicity. Several ionic liquids are known which is neither air nor water sensitive or miscible with water, thus enabling the concept of liquid/liquid extraction from aqueous media. Aromatic hydrocarbon are reported to have low activity coefficients at infinite dilution in several ionic liquids, while aliphatic hydrocarbons show high activity coefficients in the same ionic liquid stated by Huddleston et al.(1998). This means that ionic liquid can indeed be used as extractants for aromatic hydrocarbons from a mixture of aromatic and aliphatic hydrocarbons, the use of aromatic cations and possibly aromatic anions will have a positive effect on the extraction of aromatic hydrocarbons from the mixture. 63

15 The constituents of ionic liquids (being ionic) are constrained by high Coulombic forces and thus, exert practically no vapour pressure above the liquid surface. This property may allow the development of novel recovery schemes for certain organic species in relation to normal liquid/liquid extraction in which product recovery could be affected by distillation or pervaporation. The application of ionic liquid for extraction processes is promising because of their non-volatile nature explained by DuPont et al. (2000). This facilitates solvent recovery using techniques as simple as flash distillation or stripping. 64

16 2.4 Literature survey Rogers et al. (1998) were the first to report that ionic liquids can be considered as a replacement for volatile organic compounds in their paper titled Room temperature ionic liquids as novel media for clean liquid-liquid extraction. The partitioning of simple substituted benzene derivatives like aniline, toluene, salicylic acid etc. between water and the [RTIL BMIM] PF 6 was carried out. The partition coefficient derived was compared to that obtained by using 1-octanol as an extraction solvent. The effect of ph on the partition coefficient was also measured. Selvan et al. (2000) reported the liquid - liquid equilibrium data for the systems, - heptane+toluene+1-ethyl-3-methylimidazoliumriiodide [EMIM]I 3 & heptanes +toluene+1- utyl-3-methylimidazolium Triiodide [BMIM]I - 3. Tie line compositions were measured and values for Selectivity S were calculated. The analysis of both the phases was done using Gas Chromatography (GC). The column contained a precolumn to trap the IL as it does not elute in the GC. There was no IL present in the upper organic layer for all the different concentrations tried out. A system involving an IL as an extractant could potentially reduce the costs of solvent recovery as the IL will have to be recovered only from the extract. Non Random Two Liquid (NRTL) thermodynamic model was used to correlate the experimental data. The tie line compositions fitted well with the NRTL model. Holbrey et al. (2003) reported the formation of inclusion compounds in ionic liquid-aromatic compound mixtures. Benzene, toluene and xylenes are remarkably soluble but rarely completely soluble in IL s. The lower IL phase displayed a low viscosity in comparison with the IL neat, which suggests the formation of a clathrate having non stoichiometric but reproducible compositions. NMR and neutron scattering data were used to prove the formation of liquid clathrates. This paper gave a significant bearing to investigations regarding the extraction of aromatics from aliphatics using IL s. Meindersma (2005) investigated the extraction of aromatics from naptha using IL s. Separation of aromatics from the C 4 to C 10 aliphatics in naphtha is difficult as all the components have close boiling points and many of the combinations form azeotropes. An alternative, greener and more efficient extraction process was developed using IL s as extraction solvents in place of Sulpholane, which is commonly used by the petrochemical industry for this process. The values for Selectivity S and partition co toluene-heptane model was used to simulate the actual industrial separation involved. The 65

17 ionic liquids 1-ethyl-3-mehtylimidazolium ethylsulphate and 3-methyl-N-butylpyridinium tetrafluoroborate were utilised for the separations of toluene from heptanes at 40 toluene to heptanes in the feed. These values were also determined as a function of temperature. It was found that these IL s gave a higher selectivity than sulpholane. The mole fractions of toluene and heptane in the raffinate were obtained using GC while a mass balance was used to calculate the mole fractions of the components in the extract. Urszula et al. (2007) in this work, a systematic study of the impact of different factors on the phase behaviour of alkoxy-imidazolium-based ionic liquids with hydrocarbons has been presented. All systems examined showed upper critical solution temperature (UCST) behaviour, with low solubility of the ionic liquid in n-alkanes and cycloalkanes and high solubility in aromatic hydrocarbons. An increase in the alkyl chain length of n-alkane, or at benzene ring resulted in an increase in the UCST. The choice of anion was shown to have large impact on the UCST of the system: by changing the anion [BF 4 ] to [Tf 2 N], the solubility dramatically increased. The relative hydrocarbon affinity for different anions observed was [BF 4 ] <[Tf 2 N].The experimental results of LLE have been correlated using the binary parameters of NRTL equation. The average root-mean-square deviation of the equilibrium mole fraction for all the calculated values was Ternary systems 1-hydroxyethyl-3-methylimidazolium based ionic liquids with dichloromethane or 1, 2-dichloroethane, and 2-propanol were investigated by Vesna et al. (2007). For the first time, phase splitting in a (IL + dichloromethane) system was found and the related temperature composition phase diagrams for all aforementioned ILs were determined. In the case of 2-propanol, partial miscibility was observed only for ILs containing the [PF 6 ] and [BF 4 ] anions. For the ternary mixture ([C 2 OHmim] [PF 6 ] + dichloromethane + 2-propanol), a moderate pressure effect on the cloud point with a negative pressure temperature slope was detected. Mara et al. (2008) calculated the predictive capability of COSMO-RS, a predictive model based on unimolecular quantum chemistry calculations, was evaluated for the description of the liquid liquid equilibrium (LLE) and the vapour liquid equilibrium (VLE) of diverse binary mixtures of water and ILs. The effect of the ions conformers on the quality of the predictions 66

18 was assessed and the quantum chemical COSMO calculation at the BP/TZVP level derived from the lowest energy conformations was adopted. While the LLE predictions degrade with increasing the hydrophilic IL anion character, in general a good qualitative agreement between the model predictions and experimental VLE and LLE data was obtained. Anjte et al. (2010) studied the Liquid - liquid equilibrium data for ternary systems of several aromatic and aliphatic hydrocarbons with the ionic liquid 3-methyl-N-butylpyridinium dicyanamide were determined at T = K and K and atmospheric pressure. As aromatics benzene, cumene and p-xylene have been chosen, as paraffins n-hexane and n- nonane were used. A logical order in the extraction capacity of 3-methyl-N-butylpyridinium dicyanamide for the different aromatics is obtained: benzene > p-xylene > cumene. Gonzalez et al. (2010) studied the separation of toluene from aliphatic hydrocarbons (heptane, or octane, or nonane) was analysed by solvent extraction with 1-ethyl-3- methylpyridinium ethylsulphate ionic liquid, [EMpy][ESO 4 ]. Liquid (LLE) data for the ternary systems {heptane (1) + toluene (2) + [EMpy] [ESO 4 ] (3)}, {octane (1) + toluene (2) + [EMpy] [ESO 4 ] (3)}, and {nonane (1) + toluene (2) + [EMpy] [ESO 4 ] (3)} were obtained by measurements at T = K and atmospheric pressure. The selectivity, % removal of aromatic, and solute distribution ratio, obtained from experimental equilibrium results, were used to determine the ability of [EMpy] [ESO 4 ] as a solvent. The experimental results for the ternary systems were correlated with the NRTL model. Finally, the results obtained were compared with other ionic liquids and other solvents. Dominguez et al. (2011) studied the extraction of toluene from cyclic hydrocarbons (cyclohexane, or methylcyclohexane, or cyclooctane, or cyclohexene) was analysed by liquid extraction with 1-butyl-3-methylimidazolium methylsulphate ionic liquid, [Bmim] [MSO 4 ], as solvent. The experimental (liquid + liquid) equilibrium (LLE) data were determined at T = K and atmospheric pressure. Solubility curves were obtained by the cloud point method and tie-line compositions were determined by density measurement. An analysis of the influence of different cyclic hydrocarbons on the extraction was performed. 67

19 Raquel et al. (2012) studied various IL such as Liquid -liquid equilibrium data for the ionic liquids 1-ethyl-3-methylimidazolium bis(trifluoromethylsulphonyl)imide, [EMim][NTf 2 ], 1- propyl-3-methylimidazolium bis(trifluoromethylsulphonyl)imide, [PMim][NTf 2 ], 1-butyl-3- methylimidazolium bis(trifluoromethylsulphonyl)imide, [BMim][NTf 2 ], and 1-hexyl-3- methylimidazoliumbis(trifluoromethylsulphonyl)imide, [HMim][NTf 2 ], mixed with ethanol and heptane were studied at T = K and atmospheric pressure. The ability of these ionic liquids as solvents for the extraction of ethanol from heptane was evaluated in terms of selectivity and solute distribution ratio. Moreover, density and refractive index values over the miscible region for the ternary mixtures were also measured at T = K. Finally, the experimental data were correlated with the Non Random Two Liquids (NRTL) and Universal Quasi Chemical (UNIQUAC) thermodynamic models, and an exhaustive comparison with available literature data of the studied systems was carried out. Garcia et al. (2012) the use of binary mixture of ionic liquids N-butylpyridinium tetrafluoroborate ([bpy] [BF 4 ]), and 1-butyl-4-methylpyridinium bis (trifluoromethylsulphonyl) imide ([4bmpy] [Tf 2 N]) in the liquid liquid extraction of toluene from n-heptane has been investigated at K and atmospheric pressure. The experimental capacity of extraction and selectivity for this binary mixture has proved to be intermediate to those corresponding to the pure ionic liquids, and they can be predicted using a logarithmic linear model of solubility. Furthermore, the results showed that the use of binary mixture of {[bpy][bf 4 ] + [4bmpy][Tf 2 N]} at a mole solvent composition around 0.7 for [bpy][bf 4 ] improves both the capacity of extraction of toluene and the selectivity with respect to those of sulpholane, the organic solvent taken as a benchmark. Thus, this mixed ionic liquid could be likely to be used in the extraction of aromatic from aliphatic in replacement to sulpholane. 68

20 2.5 Experimental Procedure Materials and Methods Methanol (AR grade), n-heptane (AR grade), toluene rectified (AR grade), potassium hydrogensulphate (AR grade), and ethyl acetate (AR grade) were obtained from S. D. Fine Chemicals Limited, Mumbai. 1-methylimidazole (99%) and n-butyl chloride (99%) was obtained from Spectrochem Private Limited, Mumbai. 1-methylimidazole and n-butyl chloride was distilled once before use. All the ionic liquids were prepared in the laboratory following standard procedures as mentioned later in this section. Contech CA series precision balance (accuracy 10 3 g) was used to measure the mass of all chemicals used in the extraction process. The densities of the raffinate were measured by using a DA-100 Densitometer (accuracy of 10 3 g/ml) manufactured by Kyoto Electronics, Japan. Solvent were removed during the preparation of the ionic liquid using a rotary evaporator (Yamato RE300). Spectrascan UV2700 of Chemito make were used for obtaining the UV-Visible spectrum of the raffinate in order to prove the absence of IL Synthesis of [BMIM]Cl 1-butyl-3-methylimidazolium chloride ([BMIM]Cl) was prepared following a method described in literature by Huddleston et al. (1998) using 1-methylimidazole and n-butyl chloride. [BMIM]Cl was used to prepare 1-butyl-3-methylimidazolium hydrogen sulphate [BMIM]HSO 4 following the procedure described in literature by Singh et al. (2005) H 3 C N N + HSO 4 CH 3 Figure 2.2 Structure of [BMIM]HSO 4 69

21 2.5.2 Synthesis of Dicationic Ionic Liquid [MIM] 2 C n Br 2, were prepared by following the procedure described by Hua et al. (2010) bis-(3- methy-1-limidazole)-propylene hydrogen sulphate salt [C 3 (MIM) 2 ][(HSO 4 ) 2 ] and bis-(3-methy- 1-limidazole)-hexylene hydrogen sulphate salt [C 6 (MIM) 2 ] [(HSO 4 ) 2 ] were prepared by modifying the anion of reported [MIM] 2 C 3 Br 2 and [MIM] 2 C 6 Br 2 salt. [MIM] 2 C n Br 2 salt (0.01 mol) was taken in a round bottom flask with 0.02 mol of KHSO 4 anhydrous salt dissolved in 60 ml of methanol. The mixture for stirred for 24 hr at room temperature. After the reaction is completed, the residual potassium bromide was filter out and excess solvent was evaporated under vacuum. The mixture was wash repeatedly with dichloromethane, which was further dry under vacuum, to yield colourless viscous liquid. N + N CH 3 H 3 C N N + 2HSO 4 - Figure 2.3 Structure of [C 3 (MIM) 2 ] [(HSO 4 ) 2 ] 70

22 N + N CH 3 H 3 C N N + 2HSO 4 - Figure 2.4 Structure of [C 6 (MIM) 2 ][(HSO 4 ) 2 ] Experimental procedure The extraction of toluene from n-heptane was carried out in a 25 ml round bottom flask (J-Sil borosilicate glass). Various ratios of toluene to n-heptane were taken in a round bottom flask and 1 ml of the ionic liquids [BMIM][HSO 4 ], [HMIM][HSO 4 ], [C 3 (MIM) 2 ][(HSO 4 ) 2 ] and [C 6 (MIM) 2 ][(HSO 4 ) 2 ] were separately added to it. The volumes of the chemicals were accurately measured using a micro-pipette. The mixture was stirred with a magnetic stirrer in constant temperature water bath maintained at 30º C at atmospheric pressure. For the first batch, the n-heptane-rich (raffinate layer) was analysed using Gas Chromatography after every 30 minutes and the areas under the toluene peaks were found to be almost constant after a period of 1 hour. Thus, it were concluded that the time required to reach liquid-liquid equilibrium was about 1 hour. All the further batches was stirred for more than 1 hour and then allowed to settle. The amount of toluene present in raffinate was analysed with a Gas Chromatograph. A Chromline Equipment GC-100 Gas Chromatograph with SE-30 packed column and FID detector were used for this analysis. Whereas toluene present in the extract was analysed by using a pre-column 5M Z-Guard column (which retains ionic liquids), the organic layer was analysed with the GC column using FID detector. 71

23 2.6 Results and Discussion A component balance for toluene, n-heptane and IL were done to determine x(1) E, x(2) E and x(3) E, which represent the mole fractions of toluene, n-heptane and IL in the extract, respectively. Once the mole fractions of all the components in the extract as well as the raffinate are known, the partition coefficient and selectivity for toluene are determined using the equations mentioned below. fraction of toluene in the raffinate as stated by Wilson et al. (1964) x1 E... 1 x1 R Selectivity, S is defined as shown below. x1 E x1 R S... 2 x 2 E x 2 R The purpose of this work is to explore the possibility of using the monocation and dication imidazole-based ionic liquids as solvents for the extraction of toluene from n-heptane. With this aim, the liquid-liquid equilibrium for the ternary systems {n-heptane (1) + toluene (2) + [BMIM][HSO 4 ] (3)}, {n-heptane (1) + toluene (2) + [HMIM][HSO 4 ] (3)}, and {n-heptane (1) + toluene (2) + [C 6 (MIM) 2 ] [(HSO 4 ) 2 ] (3)}, {n-heptane (1) + toluene (2) + [C 3 (MIM) 2 ] [(HSO 4 ) 2 ] (3)} are experimentally determined. ed for the operation. It is observe from Table 2. toluene in the feed increases. 72

24 Table 2.1 Experimental (liquid-liquid) equilibrium data in mole fraction for ternary systems and selectivity (s) values alkane-rich phase ionic liquid-rich phase S x(1) x(2) x(1) x(2) x(3) {toluene(1) + heptane(2) + [HMIM][HSO 4 ](3)} {toluene(1) + heptane(2) + [BMIM][HSO4](3)}

25 {toluene (1) + heptane (2) + [C 3 (MIM) 2 ][HSO 4 ](3)} {toluene(1) + heptane (2) + [C 6 (MIM) 2 ][HSO 4 ](3)} Thus, the effect of the chain length of the alkane on imidazole ring on the extraction process is also studied. Figure 2.5 sh of toluene in raffinate. It is observe all the four IL systems used. 74

26 Mole fraction x(1) R Figure 2.5 Plot of the solute distribution ratio for the ternary systems {toluene (1) + heptane (2) + IL (3)} systems at T= K as a function of the mole fraction of toluene in alkanerich phase. {toluene (1) + heptane (2) + [HMIM][HSO4] (3)}, {toluene (1) + heptane (2) + [BMIM][HSO4] (3)}, {toluene (1) + heptane (2) + [C 3 (MIM) 2 ][HSO4] (3)} and {toluene (1) + heptane (2) + [C 6 (MIM) 2 ][HSO4][HSO4]} Similarly, Figure 2.6 shows variation of selectivity (S) with mole fraction of toluene in raffinate phase. Values obtained for all four IL systems decreased with an increase in mole fraction of toluene in raffinate, at lower toluene concentrations. Among all ILs studied geminal dicationic IL showed - present in raffinate. 75

27 S x(1) R Figure 2.6 Plot of Selectivity for the ternary systems {toluene (1) + heptane (2) + IL (3)} systems at T= K as a function of the mole fraction of toluene in alkane-rich phase. {toluene(1) + heptane (2) + [HMIM][HSO4] (3)}, {toluene(1) + heptane (2) + [BMIM][HSO4] (3)}, {toluene(1) + heptane (2) + [C 3 (MIM) 2 ][HSO4] (3)} and {toluene(1) + heptane (2) + [C 6 (MIM) 2 ][HSO4][HSO4]} 76

28 2.7 Data Correlation Thermodynamic models such as van Laar, Wilson, Universal Quasi-Chemical (UNIQUAC) and Universal Quasi-Chemical Functional Group Activity Coefficient (UNIFAC) models have been traditionally using to evaluate interactions between components in solutions. It is known that the Wilson et.al. (1964) and van Laar et al. (1968) models do not describe liquid-liquid equilibrium satisfactorily. UNIFAC, which is a group contribution method, requires evaluation of the binary interaction parameters. These values are well documented for n-heptane and toluene but are not available for ILs such as [BMIM][HSO 4 ], [HMIM][HSO 4 ], [C 6 (MIM) 2 ][(HSO 4 ) 2 ], and [C 6 (MIM) 2 ] [(HSO 4 ) 2 ]. The usage of UNIQUAC requires the determination of binary interaction parameters for a particular data set. This involves a lengthy optimisation process and a large number of observations. Therefore, none of these models is consider suitable for this particular ternary system. Renon et al. (1968) proposed one of the most popular models for the ionic liquid-toluene-nheptane system, is the NRTL model. The NRTL model has satisfactorily correlated for such data by Selvan et al. (2000), Meidersema et al. (2006), and Pereiro et al. (2008). However, such correlated data for the systems described in this work have not yet being, evaluated. An attempt has being made to model the data for the four systems studied in this work using the NRTL equation. A liquid phase activity coefficient for a species in solution as given by the NRTL model is: m m x G j ji ji x G x G j 1 m j ij r rj rj ln r i m m ij m x G j 1 x G x G k ki k kj k kj k 1 k 1 k 1 where, g g ij jj... 4 ij RT G exp... 5 ij ij ij 77

29 where g is the energy parameter; x represents the mole fraction; ij represents the non parameter; i is the component whose activity coefficient is being evaluated; m is the total number of components; j, k are dummy variables. ij for mixtures consisting of nonpolar and polar non-associated species can be assumed to be equal to 0.3. Therefore, for the cases studied here, the same value for the non randomness parameter has been considered. It has been assumed here that the ILs are completely in the extract and therefore, only two equilibrium equations arise, given by (6) and (7). x 1 R 1 R x 1 E 1 E...6 x 2 R 2 R x 2 E 2 E... 7 In order to obtain a solution to the set of equations (1) through (7), an initial guess value for ij was taken from guess value of Selvan et al. (2000) ij in (3) and (5). Using the values o R x1 nrtl and E x 2 nrtl are found. Equations (6) and (7) are modified into (8) and (9) as shown below in order to evaluate x1 R. nrtl z E x 1 R x 1 R nrtl exp z E x 2 R 2 R x 2 E 2 E exp nrtl The objective function (OF) used to minimise the difference between the experimental and calculated data are given by equations (10), (11), (12). The root mean square deviation (RMSD) is given by equation (13). 78

30 OF E E E x 1 R x 1 R 1 nrtl exp E x 2 R 2 R x 2 E 2 E 2 nrtl exp RMSD OF 24 d The optimum values of the binary interaction parameters calculated are those that yield a ij are shown in Table 2.2 Table 2.2 Values of NRTL binary interaction parameters for the ternary systems at T= K Guess value for solving NRTL IL1 IL2 IL3 IL The natural logarithm of the liquid phase activity coefficient for both toluene and n-heptane evaluated for all four ILs at the optimum values of the binary interaction parameters are shown in Table 2.3, Table 2.4, Table 2.5, and Table

31 Table 2.3 Activity Coefficients for toluene and heptane in alkane rich phase and IL rich phase calculated by NRTL model for [HMIM][HSO 4 ] IL1 ln (y 1 ) R ln (y 2 ) R ln (y 1 ) E ln (y 2 ) E E E E E E E E E E E E E