INTEGRATED IONIC LIQUEFACTION PROCESS

Abstract: A method of separating oil product and solid product from a single stage carbonaceous liquefaction process stream which utilizes polar solvents and alkali and alkaline earth components in an ionic liquefaction process after which non-­ condensible gases are removed from the process stream; immiscible water is removed from the process stream; alkaline-containing solids are removed from the process stream and the products present in the process stream are separated and recovered and subjected to additional stabiliza­ tion if necessary.

Description: INTEGRATED IONIC LIQUEFACTION PROCESS This invention relates to the liquefaction of carbonaceous materials and in particular to a method of producing useful products, including liquid products which can be used directly in petroleum-like refining processes, or as a feed stock for further chemical synthesis, or as a low sulfur fuel oil or the like, and, pseudo-plastic products with melting points in the range of from about 1000C to 2000C which are solid products at ambient temperature and are characterized by their reduced sulfur and ash content which make them useful as fuel and coke substitutes. One of the principal approaches to coal liquefaction and solvation in the past has employed reactions producing free radicals through thermal bond rupture. For many kinds of coal used this typically reguired temperatures above about 350 C for enough free radicals to form through thermal bond rupture of carbon-carbon bonds, carbon-oxygen bonds, carbon-nitrogen bonds and carbon-sulfur bonds, to react with other compounds or hydrogen in order to form lower molecular weight compounds than the complex materials present in the coal. In some typical prior processes the free radicals formed were stabilized and hydrogenated by hydrogen atoms through hydrogen transfer from solvent donor molecules such as l,2,3,4-tetrahydronaphthalene or 9,10-dihydrophenanthrene, often used in the presence of minor amounts of coal solubilizing agents. After such react ions the solids are generally separated by distillation, or by solvent deashing. For such systems to be effective, it is important that the mixtures be suitable for either effective distillation or deashing without excessive coking. In processes where even minor amounts of highly polar solvents are employed either in the liquefaction reactor or subsequently, distillation problems are encountered and when alkali compounds are present, severe coking problems can arise. In one embodiment, the present invention uses a mixture of carbonaceous material, polar solvent, and alkaline compounds to produce liquid and solid products by a process which will be referred to herein as ionic liquefaction. The resulting products, both liquids and solids are then separated in various process embodiments from the unreacted carbonaceous material. The products are separated and stabilized in a manner which minimizes the deletereous effects normally encountered in the distillation, separation, and stabilization of mixtures containing polar organic liquids and alkaline compounds. It is therefore an object of this invention to produce and separate liquid and solid products from carbonaceous materials, in liquefaction processes which employ mechanisms other than the formation of free-radicals. This invention provides a method for liquefaction that is characterized by reaction conditions which do not favor thermal rupture free radical mechanisms and uses primarily ionic chemical mechanisms at lower temperatures and pressures than previously employed to produce useful, novel liquid and solid products. The invention as described contemplates an integrated process for the liquefaction of carbonaceous material starting from the preparation of the feed material to the separation of products, including the recovery and use of recycle streams rich in phenolics. It is a further object of the present invention to produce a novel pseudo-plastic material that is a solid at room temperature, which has reduced sulfur and ash content and which is useful as a fuel, coke or petrochemical feedstock. These and other objects of the invention together with the features and advantages thereof will become apparent from the following detailed specification when read in conjunction with the accompanying drawings in which like reference numerals refer to corresponding parts, and: Fig. 1 is a schematic flow chart of a typical integrated ionic liquefaction process useful with the present invention for coal; Fig. 2 is a schematic representation of a liquefaction subsystem useful with the flow chart of Fig. 1; Fig. 3 is a schematic representation of a separation subsystem useful with the present invention; Fig. 4 is a schematic representation of a distillation subsystem useful with the present invention; Fig. 5 is a schematic representation of a hydrogenation subsystem useful with the present invention; ; Fig. 6 is a schematic of an acid hydrolysis subsystem useful with the present invention; Fig. 7 is a schematic of another acid hydrolysis subsystem useful with the present invention; Fig. 8 is a schematic flow chart of another typical ionic liquefaction process using acid hydrolysis useful with the present invention; Fig. 9 is a schematic flow chart of another typical ionic liquefaction process using coking useful with the present invention Fig. 10 is a schematic flow chart of another integrated ionic liquefaction process useful for the production of a pseudo plastic solid fuel and coke substitute; Fig. 11 is a schematic illustration of a liquid-liquid solvent extraction system useful in the process shown in Fig. 10; and Figs. 12 to 14 are a series of graphic representations showing carbon monoxide conversions and temperature profile plotted against time for a series of experimental runs according to the teachings of the present invention. As used herein, the term carbonaceous material, includes solid, semi-solid and liquid organic materials which are susceptible to the described treatment methods. Examples of solid carbonaceous materials which may be used in connection with the practice of the present invention include coal, such as anthracite, bituminous, subbituminous and lignite coals, as well as other solid carbonaceous materials, such as wood, lignin, peat, solid petroleum residuals, solid carbonaceous materials derived from coal, and the like, depending on the products sought. Examples of semi-solid and liquid carbonaceous materials include coal tars, tar sand, asphalt, shale oil, heavy petroleum oils, light petroleum oils, petroleum residuals, coal derived liquids and the like. Ionic liquefaction as used herein is intended to mean the chemical process described herein, which is characterized by polar solvent solubilization of the polymeric structure of carbonaceous materials susceptible to the described treatment methods, in the presence of alkali and alkaline earth compounds in amounts which favor ionic reactions involving the solubilizee carbonaceous material and ionic species such as phenoxide, hydroxide, and formate ions, and favor stabilization of the ionic species to produce distillable products, low-sulfur fuel oils, and low-ash reduced-sulfur pseudo-plastic, normally solid products useful as fuel, coke, or petrochemical feedstock. In addition to ionic reactions, it is believed that ionic liquefaction, as described herein, may change the apparent molecular weight and other physical properties of the solubilized carbonaceous material by reducing the extent of hydrogen bonding between carbonaceous material molecules. Because of the reactive nature of ionic species remaining after solubilization and ionic reaction, the product mixture must be further processed in order to stabilize the desired products to be able to recover liquid and solid usable products including a recycle stream rich in phenolics which can be used in the ionic liquefaction reactor. This processing involves removal of alkaline salts, and further stabilization of organic ionic species by hydrogenation, acid hydrolysis, solvent extraction or coking. The term alkaline as used generally and herein is synonymous with basic which includes, without limitation, alkali metal and alkaline earth compounds. A base can be an aqueous solution which contains OH ions, or any substance which accepts protons, or any substance which is an electron pair donor. Typical cat ions are the light metals of Groups IA and IIA of the Periodic Table. Preferred cations are Na and K Typical anions include OH , C03-, HC02, and HCO As applied to ionic liquefaction a polar solvent or polar solvents means a solubilizing agent selected from the group consisting of aromatic alcohols, phenols, polycyclic phenols and substituted phenols, and mixtures thereof. Typically, such solvents do not have an hydrogen. Liquid mixtures of solubilizing polar solvents used in the ionic liquefaction process of the present invention typically will contain greater than 50% by weight of such polar solvents. More particularly, in accordance with a preferred embodiment of the invention, useful hydrocarbons are obtained from carbonaceous materials by contacting the carbonaceous materials with a solvent/solute system consisting of (a) an organic phase comprising a solubilizing agent containing more than about fifty percent (50%) by weight of a phenolic-type solvent such as an aromatic phenol, polycyclic phenol, substituted phenol or a mixture thereof, and (b) an inorganic phase comprising an aqueous solution of one or more alkali or alkaline-earth metals. More particularly, the solvent/solute system comprises one or more solubilizing agents selected from aromatic phenolics, e.g. phenols, and polycyclic and/or substituted phenols, typically of from 6 to 15 carbon atoms, e.g. o-cresol, m-cresol, p-cresol, catechol, resorcinol, naphthol, and mixtures and derivatives thereof. Although not essential to the practice of this preferred embodiment of the invention other organic constitutents such as aromatic alcohols1 polycyclic aromatic hydrocarbons, partially hydrogenated and/or fully hydrogenated polycyclic aromatic hydrocarbons, typically having from 1 to 4 carbon rings, and more preferably from 2 to 3 carbon rings e.g. naphthalene, anthracene, phenanthrene, Tetralin (tetrahydronaphthalene), gamma-picoline, isoquinoline, dihydronaphthalene, Decalin (decahydronaphthalene), 9,10-dihydroanthracene, 9,10-dihydrophenanthrene, and mixtures and derivatives thereof, also may be included in the solvent/solute system. Synthesis gas, as that term is used herein, means a gas primarily comprised by carbon monoxide and hydrogen. Other gaseous components present in small concentrations can include carbon dioxide, light hydrocarbon gases1 and some impurities such as nitrogen and still be effective in the process described herein. In addition, small amounts of water vapor may also be present. It has been shown that solubilization of coal and other carbonaceous materials can be achieved using a variety of coal derived solvents and other organic solvents. For example, U.S. Patent No. 4,133,646 teaches the advantages of using minor amounts of phenolic recycle solvents in coal liquefaction. Similar advantages are taught by Kamiya et al., "Effect of Phenolic Compounds on Liquefaction of Coal in the Presence of Hydrogen-Donor Solvent", Fuel, Vol. 57 (November 1978), pp. 681-685; and by Sams, et al., "Internal Rearrangement of Hydrogen During Heating of Coals with Phenol", Fuel, Vol. 60 (April 1981), pp. 335-341. It is also known that the use of various alkali bases, e.g. NaOH, Na2CO3 and Nans03, are useful in carbonaceous liquefaction processes. See for example, Donovan et al., "Oil Yields from Cellulose Liquefaction", Fuel, Vol. 60 (October 1981), pp. 899-902 and Araya et al., "Study of Treatments of Subbituminous Coals by NaOH Solutions", Fuel, Vol. 60, (December 1981), pp. 1127-1130. Ross and Blessing have described how coal may be solubilized and hydrogen added to the organic product by aqueous base 11Hydroconversion of a Bituminous Coal with CO-H20", Fuel, Vol. 57 (June 1978), p. 379. They have also described how alcohols having an a -hydrogen atom are effective hydrogen donor solvents when catalyzed by alkaline compounds (U.S. patent No. 4,298,450). They state, however, that alcohols not slaving an a hydrogen are not effective solvents. The unexpectedly high solubilization and liquefaction achieved in the present invention through the synergistic effect in a liquefaction reactor of a solvent/solute system combining a phenolic solvent, water and added amounts of an alkali or alkaline-earth metal compound, with or without the addition of synthesis gas, depending on the reactants selected, has made possible the investigation and discovery of other promising opportunities for enhancing the liquefaction and the hydrogen to carbon ratio of the resultant products, which will be more fully described hereinafter. Appell et al., in their paper entitled "On the Mechanism of Lignite Liquefaction with Carbon Monoxide and Water", Chem. and Ind.1 Vol 47, (1967),.p. 1703, describe how using freshly powdered low-rank coal and a selected solvent will produce a 72t yield of a benzene-soluble oil when using operating pressures near 5000 psig and temperatures in excess of 365 C. Appell et al., also describes the use of a solvent comprising alpha-naphthol (a phenol), phenanthrene (a polycyclic aromatic hydrocarbon) and water in the presence of naturally occurring amounts of alkali or alkaline-earth metal compounds at similar operating conditions. In addition testing is described involving the addition of K2C03 in water as a solvent with the conclusion that the addition of K2CO3 increases the extent of the water gas shift reaction but is not believed to significantly improve hydrogen uptake by the coal during liquefaction under the conditions enìptoyed. Farcasiu et al., U.S. Patent No. 4,133,646 (1979) teach that an improved liquefaction process can be obtained using a two-stage liquefaction process wherein a donor solvent in combination with minor amounts of phenolic compounds is reacted with coal and hydrogen at 600 - 8500F. The unreacted coal was recovered by filtration, phenolic compounds were recovered by distillation, or extraction, and the resulting substantially phenolic free distillation residue was subjected to various upgrading treatments such as delayed coking and hydrotreating. Farcasiu et al, did not obtain the synergistic effect realized by the present invention using solvents containing greater than 50% by weight of phenolic compounds with added alkaline compounds in a single stage liquefaction reaction. Farcasiu et al, similarly also did not recognize that phenolic compound recovery from the product stream must be preceded by alkaline compound removal. Likewise, Farcasiu et awl does not show how alkaline compound separation may be controlled by varying the water level during the added alkaline compound removal step. Further, there is no disclosure in Farcasiu et al of how the process could be advantageously improved by adding synthesis gas to the first stage reaction, either as a reactant in the formate ion chemistry described hereinafter, or to produce a hydrogen enriched stream for upgrading operations. Beside the foregoing, Farcasiu et al does not disclose that residues containing more than about 50% by weight phenolics may be upgraded by hydrotreating or coking, but instead indicates that the residue should be substantially phenol free. The foregoing and other art, in summary, has not recognized that a liquefaction process for coal or other carbonaceous material can be substantially improved by the use of an organic solvent containing greater than 50% by weight of phenolic compounds in amounts between about 1.5 to 5.0 times the weight of carbonaceous materials1 in combination with between about 25 to about 400 parts by weight of alkali for every 1000 parts by weight of carbonaceous material, and between about 25 to about 400 parts by weight of water for every 1000 parts by weight of carbonaceous material; when the carbonaceous material, the organic solvent and the solvent/solute pair are reacted together at temperatures less than about 360 C and pressures between about 300 psia (2.06 MPa) to about 2500 psia (17.2 MPa). Further improvement can be obtained by the presence of synthesis gas having a ii2/CO ratio between about 0.5 and 2.0 in amounts between about 0.16 to about 1.25 m3/Kg of carbonaceous material when combiiied with stabilization of the reaction product before further upgrading, after removal of the alkali and water. More particularly, in accordance with a preferred embodiment of the invention the solvent/solute systems useful in the practice of the invention are solubilizing mediums comprising organic and inorganic fractions or constituents which may solubilize a portion of the carbonaceous material and/or may otherwise enhance liquefaction of the carbonaceous material. As noted supra, the organic fractions of the solvent/solute systems comprise one or more solubilizing agents selected from the group consisting of aromatic phenolics1 e.g. phenols, and polycyclic and/or substituted phenols, typically of from 6 to 15 carbon atoms, e.g. o-cresol, m-cresol, p-cresol, catechol, resorcinol, naphthol, and mixtures and derivatives thereof. Although not essential to the practice of this embodiment of the invention the solvent/solute systems in many instances will include other organic constituents. Suitable organic constituents include aromatic alcohols, polycyclic aromatic hydrocarbons, partially hydrogenated and/or fully hydrogenated polycyclic aromatic hydrocarbons, typically having from 1 to 4 carbon rings, and more preferably from 2 to 3 carbon rings e.g. naphthalene, anthracene, phenanthrene, acenapllthene, l-methylnaphthalene, 2-methylnaphthalene, Tetralin (tetrahydronaphthalene), gamma-picoline, isoquinoline, dihydronaphthalene, Decalin (decahydronap?lthalene), 9,10-dihydroanthracene, 9,10-dihydrophenanthrene, and mixtures and derivatives thereof. During initial phases of operation, some of the above mentioned solubilizing agents and/or other organic constituents will be present: then in subsequent Operation, the organic constituellts will be carbonaceous material-derived phenols of the type; and polycyclic aromatic hydrocarbons of the type, or derivatives related to the type, described hereinbefore. Particularly useful organic phase solubilizing agents and/or other organic fraction constituents have a boiling point above 50 C, more preferably of from about 1000C to about 4600C, and most preferably of from about 1500C to about 4000C. In the practice of this latter embodiment, the solubilizing agent is typically from about 50 to 100 weight percent of the organic fraction of the solvent/solute system. Suitable inorganic fraction constituents of the solvent/solute system include water, and alkali and/or alkaline-earth metal compounds, and their derivatives. The water content can be from about 5 parts to about 60 parts per 40Q parts by weight of the solvent/solute system, more usually about 15 parts to about 40 parts per 400 parts by weight of the solvent/solute system. Suitable examples of alkali and alkaline-earth metal compounds include hydroxides, carbonates, bicarbonates, nitrates, sulfates, sulfites, sulfides, formates and other salts, mixtures thereof 1 and the like, although other compounds may be employed for the purpose. Specific examples incude NaOH, Na2CO3, NaRCO31 Na2SO4, NaNO31 KOH, K2C03, KEiCO2, CaC03, mixtures thereof and the like. Presently preferred species are NaOH, KOH and Na2CO3 in from about 1 part to about 40 parts per 400 parts by weight of the solvent/solute system, more usually 1 to about 15 parts per 400 parts by weight of the solvent/solute system. It is understood that the amount of alkali or alkaline-earth metal compound present for purposes of the present invention is an added amount, i.e. an amount in excess of the amount which would be present from the various naturally occurring alkali ot alkaline-earth metal compounds. Further, it is understood that the alkali or alkaline-earth metal compound content will be maintained at the desired level in a recycle solvent stream. As will be seen in Example VI et seq., the combination of organic and inorganic fractions and constituents provide a beneficially synergistic effect on solubilizing of carbonaceous material. The amount of the solvent/solute system required in the reaction mixture or slurry is dependent upon the amount and nature of the carbonaceous material to be treated. Generally, it is preferred to employ up to about 500 parts of the solvent/solute system to 100 parts of carbonaceous material, more preferably at least about 350 parts of solvent/solute to 100 parts of carbonaoeous material, and most preferably at least 150 parts of solvent/solute system to lC parts of carbonaceous material According to this embodiment, carbonaceous material is solubilized in the solvent system - alkali medium to form a reaction mixture or slurry. Frequently, the reaction conditions are water gas shift reaction conditions, as hereinbefore described. The reaction mixture is heated to a sufficient temperature, typically below about 4O00C., and pressure to obtain enhanced solubilizing of the carbonaceous material for production and/or conversion of hydrocarbon liquids, as hereinbefore defined, from the carbonaceous material. For most purposes1 it is contemplated that sufficient temperature levels for the solvent/solute system are from about 1000C to a temperature below about 400 C under the reaction conditions employed, more preferably from about 140 C to about 380 C, and most preferably from about 260 C to about 360 C, at a pressure of at least about 300 psia (2.06 MPa), more preferably from about 500 psia (3.45 MPa) to about 2500 psia (17.2 MPa), and most preferably from about 500 psia (3.45 MPa) to about 1500 psia (10.35 MPa). It has been found that under the foregoing reaction condition, relatively short periods of time result in the production of the desired product. Although sufficient times are dependent upon the nature of the carbonaceous material, the reaction conditions employed, and the like, for the most purposes, it is contemplated that reaction times of at least about 1 minute, more preferably from about 10 minutes to about 120 minutes, and most preferably from about 15 minutes to about 60 minutes are sufficient to result in enhanced solubilizing and the production and/or conversion of hydrocarbon liquids. As will be appreciated by those skilled in the art, the solvent/solute systems containing coal or othercarbonaceous material solubilized according to the present invention, may be further treated as described herein. In ionic liquefaction the reaction mechanisms of the chemistry is substantially different than the previously employed reaction conditions which favored thermal rupture-free radical chemical reactions. At the lower temperatures used for ionic liquefaction, the predominant chemistry can be termed as solvation-ionic chemistry, involving solubilization of the coal polymer by polar phenolic solvents, followed in situ by attack upon the coal structure by ionic species such as phenoxide and formate ions. Without being limited to any specific theory, it is believed that the primary points of attack are at the carbon adjacent to oxygen containing functional groups present in the coal. The esters, ketones, and ethers present, are all sites for nucleophilic substitution. hydroxide and carboxylic acid containing functional groups are considered to be substantially unreactive to nucleophilic substitution. Thep7ienoxide ion mechanism for an ether group is: Initiation EMI15.1 Reaction EMI15.2 Hydroxide Regeneration EMI15.3 Solvent Regeneration EMI15.4 The parallel for mate ion mechanism V to VII is EMI16.1 This mechanism is believed to be enhanced at the mild conditions employed because of the intimate contact made possible by the phenolic solvent sQlubilization where there is competition for available ionic species in the presence of the water gas shift reaction VIII to X. EMI16.2 The summation of VIII, IX and X is the net water gas shift reaction XI: EMI16.3 Ionic liquefaction, therefore, offers many process advantages over conventional thermal rupture free radical liquefaction. The temperature regime for ionic liquefaction is typically under 360 C. At these temperatures thermal bond rupture is not the primary mechanism because it proceeds slowly, The lower temperature employed has the effect therefore of reducing the amount of methane, ethylene, ethane, and acetylene produced from 20% to less than 1% of the MAF carbonaceous material System pressure for ionic liquefaction can be obtained from the vapor pressure of the solvents alone or by the addition of external pressurized gases. Sufficient pressure is preferably maintained to ensure that a majority of the solvents are in the liquid phase. This is typically 3.45 - 10.35 MPa (500 - 1500 psia). The ionic liquefaction mechanism described herein to produce a liquid, can, in addition, use synthesis gas which is composed of carbon monoxide and hydrogen instead of requiring expensive pure hydrogen. This synthesis gas is the fuel for the production of pure hydrogen, therefore the production of pure hydrogen is not needed to obtain the desired improvement in hydrogen to carbon ratios possible in the liquid products of ionic liquefaction. In free radical liquefaction, a typical mechanism is: Thermal Rupture EMI17.1 <tb> XII. <SEP> Coal------ <SEP> 2R <tb> Hydrogen Donation EMI17.2 Tetralin Dihydronaphthalene (Tetrahydronaphthalene) Solvent Regeneration EMI17.3 In order to promote the Solvent Regeneration reaction, (equation XIV) free radical liquefaction is typically performed using high purity hydrogen at pressures between 13.8 - 17.2 MPa (2000 - 2500 psia). The ionic liquifaction described herein, at the conditions employed, will produce polar products having low molecular weights. The free-radical mechanism will produce mainly non-polar materials. The free-radical type mechanism can, in addition1 lead to polymerization reactions which produce undesirable high molecular weight materials. These type of reactions are not favored and therefore avoided in the ionic liquefaction described herein. The presence of highly polar phenolic compounds1 alkali compounds, and the ionic forms thereof, along with the oxygenated compounds derived from the carbonaceous material, during the ionic liquefaction process described herein, can, however, lead to potential processing problems downstream from the liquefaction reactor. Some oxygenated compounds, including many phenols, are often thermally unstable, especially in the presence of coking promoters such as ionic alkaline species. These compounds are concentrated in the bottoms during distillation operations. It is, therefore, important to minimize the amount of alkaline compounds present in the distillate feed stream. A major proportion of the alkaline compounds should, therefore, be removed before distillation is begun. In addition, the proposed phenoxide ion chemistry shows a solvent incorporation step as an inherent part of the reaction mechanism. Under the conditions typically used in ionic liquefaction, i.e., relatively low temperature and low pressure, it is difficult for the solvent regeneration reaction IV to proceed to completion. Therefore, to regenerate solvent and obtain additional product, a portion of the liquefaction product liquids should preferably undergo further reaction in the presence of hydrogen under conditions which will break the ether bond between the phenoxide ion and the coal derived organic species. The reaction may be performed at conditions which are severe enough to break the ether bond, but are not severe enough to saturate the aromatic ring, or remove the oxygen atom as water. The hydrogenation, acid treatment or solvent extraction steps will also serve to stablize the product by reducing the concentration of the most unstable ionic species and reduce the ash content of the product. Because the ionic liquefaction products produced by the liquefaction processes described are typically high in reactive oxygenated species there can be a tendency for the products to oxidize and/or polymerize with time or with thermal treatment. In the case where additional solvent recovery and product upgrading are necessary, specific sequential processing steps are then needed. A process incorporating these necessary steps is shown in Figure 1. Referring to the schematic diagram in Fig. 1, the feed preparation at (A), comminutes the carbonaceous material, stream (1), by conventional' means such as hammermills or ball mills or comparable equipment: and adds a water-alkali mixture stream (4); and recycle polar solvent streams (2) and (3), containing greater than 508 by weight of phenolic species. The comminution process may be accomplished either dry or wet. If performed wet, then the recycle polar solvent may be used as the wetting agent, if proper precautions are taken. The carbonaceous feed is preferably comminuted to 100 percent minus 74 micron (200 mesh) particle size, more preferably to 100 percent minus 147 microns (100 mesh) particle size, and most preferably to 100 percent minus 350 microns (40 mesh) particle size but in any event must be in a form which will enable the requisite solubilization for the ionic liquefaction to proceed. Using 1000 parts by weight of stream (1) carbonaceous material as an example, the preferred amount of polar recycle solvent for the required solubilization to proceed, recycle streams (2) plus (3) is between 1500 and 3500 parts by weight depending on the prepared form of the carbonaceous material1 with 2000 parts by weight of solvent the most preferred amount. The polar recycle solvent contains preferably greater than about 50% by weight of phenolic compounds, and more preferably greater than 60% by weight phenolic compounds. The preferred amount of alkaline material in water-alkali mixture stream (4) is selected to be enough to produce the desired results. It has been found under the conditions disclosed herein that between about 25 parts and 400 parts by weight per 1000 parts by weight of carbonaceous material is effective with the more preferred amount being between about 25 and 150 parts by weight, and depending on the kinds and amounts of phenolic compounds employed, the preparation of the carbonaceous material and the conditions selected, the most preferred amount is about 50 parts. The amount of water in stream (4) should preferably be sufficient to maintain the alkaline material in the ionic form in solution at the herein described ionic liquefaction conditions, and sufficient to allow the water gas shift reaction (equation XI), to produce the hydrogen required for solvent regeneration and ionic species stabilization in the hydrogenation (E) step. The amount of water in water-alkali mixture stream (4) is between about 25 and 400 parts by weight, more preferably between about 50 and 250 parts by weight, and most preferably between about 100 and 200 parts by weight. Make-up alkali may be provided to stream (4) via a make-up alkali stream (4A). The liquefaction section (B) is presented in more detail in Fig. 2. A high pressure slurry pump (K) can be used to pump the slurry, stream (5) from the preparation section (A) of Fig. 1 and to bring the slurry to the desired system pressure. The slurry then passes tllrougll tile slurry-hot product heat exchanger (L) to provide initial slurry heating through waste heat recovery from the product slurry stream in the separation section (C) of Fig. 1. The slurry is then mixed with high pressure synthesis gas, stream (6), from the gasification/gas processing section, sections (G) and (ill) of Fig. 1. If synthesis gas is used, the composition of the synthesis gas stream should primarily be composed of carbon monoxide and hydrogen, although small amounts of impurities such as nitrogen and carbon dioxide can be present and still achieve the most preferred results. 1lse preferred i2/CO ratio under the conditions described is between about 0.5 and 2.0; a more preferred ratio is between about 0.5 and 1.5; and the most preferred ratio is between about 0.75 and 1.25. Other concentrations under different reaction conditions may also be effective. Preferred gas treatment rates are between about 1.25 m3/kg carbonaceous material and 0.16 m3/kg (40,000 SCF/ton to 5,000 SCF/ton), the most preferred gas treatment rates are between about 0.47 and 0.16 m3/kg (15,000 to 5,000 SCF/ton). Referring to Fig. 2, the reaction mixture, stream (26), is then brought to the desired temperature in a gas-fired tube heater (M), operated under turbulent flow conditions to optimize heat transfer. Process derived low or medium calorific value gas from gas processing sections (E1) and (I), lines HF and IF, is used to fire the preheater. The reaction mixture stream (27), is then transferred to the liquefaction reaction (N) and held at the desired reaction temperature long enough for the initial ionic liquefaction process steps to take place to the desired extent. The solubilization and ionic reaction steps, but not the solvent regeneration or ionic species upgrading, take place in this reactor. The preferred reactor design is a high length to diameter ratio reactor with internal baffles to provide sufficient mixing of the polar solvent, alkali1 and carbonaceous materials for the reaction mechanisms to proceed efficiently. Depending upon the desired product distribution, and desired reactant residence time distribution, other reactor designs can1 of course, be.. selected for use. Temperature and pressure may be adjusted for the optimum conversion of individual carbonaceous feed materials. Preferred reactor (N) residence times are from about 2 to about 120 minutes, and most preferably from about 15 to 45 minutes depending on reactor design. The preferred temperature range is from about 250 C to 350 C. (482 F to 662 F). It is important that a system pressure is selected and maintained at such a predetermined level which will keep sufficient water in the liquid phase to maintain ionic alkaline species in the liquid phase, rather than as salts, during a substantial portion of the reaction time. The preferred pressure range is between about 3.45MPa to l7.24MPa (500 psia to 2500 psia), and the most preferred pressure range is about 5.52MPa to 8.28MPa (800 psia to 1200 psia) for achieving the foregoing liquid phase. The reaction products exiting reactor (N), stream (7) of Figs. 1 and 21 are then separated into component streams in the separation system, section (C) of Fig. 1. A more detailed schematic of this system is shown by Fig. 3. The reaction products stream (7) is fed to gas-slurry separator(S). The gas-slurry separator(s) serve to separate the majority of the slurry product from gaseous products. The separator pressure is maintained at a lesser or reduced pressure than the liquefaction reactor1 preferably between about 0.69 and 3.45 MPa (100 - 500 psia), and the temperature is kept at conditions sufficient to keep most of the volatile organic compounds in the liquid phase, but much of the remaining water as vapor. The preferred temperature range is 200 to 300 C (392 to 572 F); the most preferred temperature range is 200 to 250 C (392 to 482 0F). The slurry phase residence time is preferably less than 30 minutes, more preferably less than 15 minutes, and most preferably less than 5 minutes. The vapor, stream (28), from the gas-slurry separator is passed to a water condenser (P) where the majority of the water is condensed, along with residual organic compounds, such as polar solvents and (C4+) hydrocarbons. Typical concentrations of water vapor in the gas phase entering the condenser will be between about 25 to 45 mole percent and leaving the water condenser the water vapor in the gas phase will have been reduced to between about 0.01 and 0.05 mole percent. The other components of the gas stream are typically carbon monoxide, carbon dioxide, C1 through C3 hydrocarbons, COS, H2S, NI13 and possibly small amounts of other gases, such as nitrogen. The preferred H2/CO ratio of this gas stream should generally be greater than about 2 to 1, and most preferably greater than about 9 to 1. The water rich stream, stream (31), from the condenser (P) is preferably withdrawn through, stream (33), from the separation system, and may be used for plant cooling water or the like. however, water may be added back to the slurry from the gas-slurry separator (stream 32), if needed, to improve the slurry filtration characteristics. The slurry stream (24), goes to the feed slurry-hot product heat exchanger (K), then through a pressure let-down valve and through stream (30), to the filtration apparatus (Q). Leaf filters, candle filters1 centrifuges, hydroclones, or comparable equipment can be used for the filtration. Solids may also be separated by processes such as solvent deashing, or critical solvent extraction. The purpose of the filtration step is to separate unreacted carbonaceous material and alkaline salts from the ionic liquefaction product liquids. Due to the hygroscopic nature of the alkaline salts, much of the water present will also be separated with the filter cake. Tile preferred temperature for filtration is preferably between about 150 and 100 C (302 and 212 F). The pressure is maintained at a sufficient level to obtain efficient filtration, preferably between about 0.34 and 1.03 MPa (50 psia to 150 psia). In the filtration step, the solids content of the liquid is typically reduced to less than about 1.0 percent by weight, the mineral matter content to less than about 0.5 percent by weight, and the alkaline species content to less than about 0.25 percent by weight. The values obtained are dependent upon factors such as the degree of comminution used in feed preparation, the ionic liquefaction conditions selected, and the design of the filter equipment design. The separation train should preferably be operated to reduce the alkaline compound content of the filtrate stream (10) of Fig. 1 below about 0.25 percent by weight, more preferably below about 0.15 percent by weight, and most preferably below about 0.10 percent by weight. Alkaline compound concentrations above this level are not desirable since that can lead to coking problems, and to unwanted organic ionic compound precipitation in downstream processing steps. Liquids, stream (10) of Fig. 1, with alkaline compound contents of about 0.25 percent by weight and lower may successfully be further processed by distillation (D) and hydrogenation (E)1 although the upper limit is dependent upon the exact nature of the product. The solids rich filter cake, stream 8 is used as feed to a gasification system (G) for production of synthesis gas and recovery of alkaline compounds (J). The gas stream, stream (9B) of Fig. 1, is sent to gas processing (I) for upgrading before use as hydrogen rich gas in subsequent hydrogenation (E) operations. The liquids from the separation system are sent, stream (10) of Fig. 1, to a distillation tower (D) where the crude ionic liquefaction product is fractionally distilled. The products from the distillation are water, stream (11) of Fig. 1, a recycle solvent stream, stream (2) of Fig. 1, and a crude product1 stream (12) of Fig. 1. A schematic of a typical distillation column is given in Fig. 4. The distillation operation has three primary purposes. First, excess water is removed from the crude product stream. Second, a polar recycle solvent stream containing greater than about 50 weight percent phenolic materials is recovered. Third, a concentrated ionic compound stream is produced. The production of the concentrated crude stream serves to reduce the size of downstream stabilization operation. This operation is typically accomplished by taking a 2000C (3920F) cut of the incoming feed. The distillation column is typically operated at or near atmospheric pressure preferably at about 0.10-0.15 MPa (14.5-21.75 psia). The condenser is operated as a partial condenser to separate the overhead product into an organic rich liquid phase and a water rich vapor phase. Using typical ionic liquefaction feed ratio of about 2 parts solvent to 1 part carbonaceous feed, in stream (5) of Fig. 1, the overheads to bottom ratio will be about 1 to 1, although the exact value will vary depending on the bottoms viscosity desired. The physical properties of the distillation feed and bottoms streams are a function of the carbonaceous feed and the ionic liquefaction conditions. In a process using a feed comparable to a Texas lignite and about a 2 to 1 solvent to feed ratio, the preferred viscosity of the distillate feed is between about 1 x 10-5 and 1 x 10-4 m2/s at 38 C (10 and 100 centistokes at 100 F). The preferred density of this stream is between about 1000 and 1200 kg/m3 (1.0 and 1.2 grams/ml). The overheads organic liquid preferred viscosity is between about 5 x 10-6 and 2 x 10-5 m2/s at 38 C (5 to 20 centistokes at 100 F). The preferred density of this stream may be between about 950 and 1100 kg/m3 (0.95 ad 1.1 grams/ml). The preferred viscosity of the crude is between about 5 x 10-4 and 2 x 10-3 m2/s at 38 C (500 and 2000 centistokes at 100 F). The preferred density of the crude is between 1100 and 1200 kg/m3 (1.1 and 1.2 grams/ml). The crude will normally contain at least about 40% phenolic compounds. Three processes may be used in the ionic liquefaction process described to upgrade the resultant crude. The first embodiment, shown in Fig. 1, is hydrogenation. Alternatively, tile crude may be stablized by acid hydrolysis or by coking. A schematic of a typical hydrogenation and distillation train is shown in Fig. 5, a schematic of a typical acid hydrolysis train is shown in Fig. 6, and a schematic of a typical process incorporating coking is shown in Fig. 9. Referring to the process schematic shown in Fig. 5, the hydrogenation train may consist of four major processing units: (1) crude preheaters (R) and (S ), (2) a crude hydrogenation reactor (T), (3) a gas-oil phase separator (U), and (4) a catalyst regeneration-alkali recovery unit (V). The incoming crude stream (12), is first preheated in a waste heat recovery preheater (R) through heat exchange with the refined crude oil, stream (17) from the gas separator. The remainder of the heating is accomplished in a gas-fired tube heater (S'). The preheated crude stream is then passed to the hydrogenation reactor (T) where it is reacted with a hydrogen-containing gas which is introduced into the hydrogenation reactor (T) as stream 13. The preferred operating condition for the hydrogenation reactor (T) is in a slurry phase catalytic hydrogenation mode. The preferred operating conditions permit separation of the refined crude and catalyst using a solids disengaging zone, and a catalyst withdrawal operation. Preferred hydrogenation conditions are those which are severe enough to break ether bonds and upgrade ionic species, but which are not severe enough to favor saturation of aromatic rings or the removal of organic oxygen as water. Preferred operating temperature is between about 343 and 4540C (650 and 8500F, and most preferably between about 343 and 4000C (650 and 7520F). The preferred pressure range is between about 6.9 and 13.8 MPa (1000 and 2000 psia). The preferred catalyst types are standard hydrogenation catalysts such as Co-Mo/Al203 and Ni-Mo/A1203, typically about 1/32, 1/16 or 1/8 inch extrudate form. Preferred catalyst loadings are about 0.01 to 1.0 kg cat/kg oil/hr, more preferred loadings are about 0.01 to 0.5 kg cat/kg oil/hr, the most preferred loadings are about 0.05 to 0.15 kg cat/kg oil/hr. Preferred hydrogen treatment rates are about 178 to 1424 m3H2/m3 oil (1000 to 8000 SCFI12/BB1 oil): more preferred hydrogen treatment rates are about 178 to 712 m3H2/m3 oil (1000 to 4000 SCFn2/BB1 oil. Preferred hydrogen comsumption is less than about 3.90kg H2/m3 oil (1.3G lb H2/BB1 oil), and more preferably less than 3 about 1.94 kg ,/m oil (0.68 lb El2/BBl oil). It is desirable to hold the crude at the temperature selected for sufficient time to permit the solvent regeneration and ionic stablization reactions to occur. Preferred residence times are between about 10 and 90 minutes, more preferably between about 15 and 60 minutes, and most preferably between about 30 and 45 minutes. In the catalyst regeneration-alkali recovery unit (V) the catalyst rich slurry is first degassed, and the gases, stream (42), mixed with other exit gases, stream (39), from the hydrogenation train, and passed via stream (18) to gas processing unit (I). The catalyst and refined oil are separated by standard solid-liquid separation devices such as filters or hydroclones. Residual refined oil and soluble alkali salts are recovered by hot water leach. The leachate is recycled to the liquefaction reactor, unit (B) of Fig. 1. Coke is burnt off the catalyst by fluid bed combustion in the presence of added air or oxygen, spent catalyst is removed, stream (15), makeup catalyst is added, stream (41B) and the regenerated catalyst, stream (16), is recycled to the hydrogenation reactor (T)(Fig. 5). Refined ionic liquids in stream (17) are separated from the treatment gases in a gas liquid separator (U). The gases in stream (39), are sent to gas processing for upgrading by removal of acid gases and light hydrocarbons. The refined liquids from the hydrogenation train, depending on the conditions employed, will preferably have been converted to greater than 80 wt percent oils, more preferably than 90 wt. percent oils, and most preferably greater than 95 wot. percent oils, as defined by pentane solubility and ASTM D1160-77 distillation results. The preferred viscosity of the refined oil is between about 5 x 10-6 and about 1 x 10-4 m2/s at 38 C (5 to 100 centistokes at 100 F), and more preferably between about 5 x 10 6 and 2 x 10 m2/s (5 to 20 centistokes at 100 0F). The preferred density of the refined oil is between about 1000 and 1100 kg/m3 (1.0 to 1.1 grams/ml). The refined oils, stream (17) of Fig. 1 and stream (38) from catalyst regeneration alkali recovery unit (V), are then sent to a distillation unit, unit (F) for final processing as will be described in detail hereinafter. The second upgrading embodiment is acid hydrolysis, as shown in Fig. 6. In this process alternative, hydrogen is added to the ionic organic species in the crude through ionic reactions with acid. In one method, carbonic acid produced in the acid hydrolysis train is used as the hydrogen source, although small amounts of other acids alone or in combinations can be used, such as sulfuric acid, hydrochloric acid, formic acid, acetic acid, and carbamic acid, which will enhance the rate and extent of ionic hydrogenation. Referring to Fig. 6, the acid hydrolysis train consists of five units, (1) a gas absorber tower (AA) to produce carbonic acid, (2) a crude stabilizer (BB), (3) a liquid phase separator (CC), (4) a high-oxygen product distillation tower (DD), and (5) low-oxygen product distillation tower (EE). In tlle preferred operation, carbonic acid is produced through gas absorption in a counter-current gas absorber (AA). Water is fed in, stream (101) at the top of gas absorber tower (AA), and carbon dioxide or carbon dioxide rich gas from the ionic liquefaction reactor or other source is fed in the bottom of the gas absorber tower (AA) stream (9B). The absorber is operated in such a manner to produce carbonic acid, stream (103), with a concentration preferably between 1650 mole/m3 (0.103 lb. mole/ft3) and 40 mole/m3, more preferably between 1650 mole/m3 and 1000 mole/m3, which is passed to crude product stabilizer (BB). The absorber is preferably operated in the temperature range 16 to 66 C (60 to 150 F), and more preferably in the range 16 to 32 C (60 to 900F). The preferred operating pressure is in the range of about 0.34 to 10.35 MPa(50 to 1500 psia)1 and more preferably in the range 6.90 to 10.35 MPa (1000 to 1500 psia). The tower may be operated as a bubble cap column, although other designs such as packed columns, venturi scrubbers, or spray towers may be used. Excess carbon dioxide gas is removed, stream 102, and may be recycled in known manner to gas absorber tower (AA). The ionic acid-base reactions take place in the crude product stabilizer (BB). Stabilizer (BB) is operated in a stirred tank mode to maximize the contact between the aqueous and organic phases. Although dependent upon the degree of stabilization required and the strength of the acid, the preferred feed ratio by volume to the stabilizer is 10 parts acid to 1 part topped crude, more preferably 5 parts acid to 1 part crude, and more preferably 1 part acid to 1 part crude. The preferred temperature is less than about 2000C (392 0F), more preferably less than about 1500C (3020F), and most preferably less than about 100 C (2120F). The pressure is kept sufficiently high to keep the majority of the carbon dioxide in the liquid phase, preferably between 6.no and 13.80 MPa (1000 and 2000 psia), and ::ore preferably between 10.35 and 13.80 MPa (1500 and 2000 psiG). Time acid and crude are maintained at stabilization conditions for sufficient time to permit the desired stabilization reactions to take place. The preferred time is between about 5 and 90 minutes, and a more preferred time is between about 15 and 45 minutes. In an alternate configuration the carbon dioxide rich gas and water are added directly to the stabilizer, eliminating the gas absorption unit. Stabilization conditions employed can remain the same. A two phase aqueous-organic product, stream (104), is withdrawn from crude product stabilizer (BB) and is sent to a liquid separator (CC) where gravity separation is used to generate a light aqueous-rich product stream (106) and a heavy organic-rich product stream (105). Liquid separator (CC) may be operated at ambient temperature. The preferred residence time is between about 10 and 60 minutes1 and more preferably between about 10 and 30 minutes. The light aqueous-rich stream is passed to a distillation unit (DD) where water, stream (107), an oxygen rich water soluble product stream (108), and a residual product stream (109) are recovered. Because of the normally high solubility of alkaline compounds in the water phase1 the majority of residual alkaline materials will be present in the residual product from this distillation unit. The heavy organic rich phase (stream (105), is passed to a second distillation unit (EE), where various distillate products such as naphtha, stream (110), light and heavy gas oil fractions, streams (111) and (112), and a residual product stream (113) are obtained. An alternate acid hydrolysis configuration is shown in Fig. 7, where an acid other than carbonic acid is used as the hydrogen donor species. The major processing units in this embodiment are a crude product stabilizer (FF), a liquid separator (GG) and two distillation towers, (HH) and (II). In the crude product stabilizer the crude product stream (12) is stabilized with an acid from stream (115) at a temperature sufficient to permit the ionic stabilization reaction to proceed to the desired extent. The preferred acid for this embodiment is sulfuric acid. A portion of the sulfuric acid could come from acid gas removal operations, unit (I) of Fig. 1. Additional required sulfuric acid would have to be added as make-up acid, stream (121). The acid and crude are intimately contacted in a stirred tank reactor operation under acid reflux conditions that is, at the boiling point of the acid. The preferred ratio of acid to topped crude is about 10 to 1, a more preferred ratio is about 5 to 1, and the most preferred ratio is about 1 to 1. The temperature of the system will be predominantly governed by the boiling point of the acid solution which is peferably between about 204 and 100 C (400 and 212 F), more preferably between about 149 and 100 C (300 and 2120F), and most preferably between about 120 and 1000C (248 and 2120F), The pressure is preferably maintained at about atmospheric pressure. The residence time in the crude product stabilizer is preferably between about 5 and 90 minutes, and more preferably between 15 and 45 minutes. The two-phase aqueous-organic product in the stream (114) is then sent to the liquid separator (GG) where gravity separation is used to generate a light aqueous-rich product in a stream (116), and a heavy organic-rich product in a stream (117). The preferred residence time is between 10 and 60 minutes and more preferably between 10 and 30 minutes. The light aqueous-rich stream (116) leaving the liquid separator is divided into two streams (118 and 120). Stream (120) is mixed with make-up acid (121) and recycled to the crude product stabilizer (FF). As noted supra, make-up acid can come from the acid-gas removal unit in the gas processing unit1 (I) of Fig. 1, and from purchased acid. The remaining aqueous liquids in the stream (118) are sent to a distillation tower (II) where water is vaporized, and transferred in stream (119) for recycling to the crude product stabilizer (FF). An organic distillate, stream (126) and a spent acid stream, stream (127) consisting of heavy organics and alkaline salts such as Na2 SO4 are also produced. The organic rich phase in stream (117) is passed to a second distillation unit (like) where various distillate products such as naphtha, stream (122), light and heavy gas oil, fractions, streams (123) and (124) and residuum, stream (125) are obtained. Fig. 8 presents a modified schematic of the integrated ionic liquefaction process using acidification in an acidification unit (FF) for the ionic liquid stabilization, and acid recovery in an acid recovery unit (AR). A third upgrading embodiment is coking as shown in Fig. 9. In this processing option the crude stream (12), from distillation tower (D) is further upgraded in a delayed or fluid bed coking operation (KK), with coke, stream (K1) being produced, and water, stream (K2) being removed. Coking is defined as a severe thermal cracking process in which one of the end products is a carbon rich solid1 i.e. coke, the other products are hydrocarbon gases, and liquids. In the present invention, the coking operation will proceed at about atmospheric pressure and at temperatures from about 4270C (8000F) to 5100C (9500F). Using 1000 parts of crude as an example1 the coking operation may produce from about 250 to 300 parts coke, from about 80 to 120 parts (C4) hydrocarbons, from about 300 to 50 parts water, and the remainder as liquids. The liquids produced will be used as recycle solvent. The gases will be sent, stream (131) to gas processing. Support equipment and facilities for ionic liquefaction can be commercially obtained and include gasification and various gas processing operations such as compression, acid gas removal, and shift conversion processing to increase hydrogen content of gas streams. The actual configuration is dependent upon the type of stabilization operation used, that is hydrogenation, acid hydrolysis, or coking. Referring to the integrated process using hydrogenation for stabilization, Fig. 1, the filter cake produced in the separation operation, unit (C), is used as a feed stream (8), to a gasifier1 unit (G). In the configuration shown in Fig. 1 the preferred gasification operation is a dry ash partial combustion process to produce a synthesis gas rich in carbon monoxide and hydrogen. The preferred 112/CO ratio of gas leaving the gasifier, stream (22), is between 0.5 and 2.0. Acid gases and light hydrocarbon gases are separated from the synthesis gas by standard operations in gas processing operations, unit (H). The ash from the gasifier, stream (21) is sent to an alkali recovery unit (J) where the majority of the alkaline compounds are separated from the ash stream by hot water extraction, and recycled, stream (4), to the preparation unit (A), and the residual ash removed, via stream (J1). The expected recovery of alkaline compounds is greater than 508, and can be greater than 75E, and even greater than 90%. The preferred temperature is between 25 and 1000C (77 and 2120F), and more preferably between 50 and 100 C (122 and 212 F). The preferred treatment rate is less than 4 kg H2O/kg ash, more preferably less than 2 kg H2O/kg ash, and most preferably less than 1.1 kg H2O/kg ash. In an alternate configuration alkaline compounds may be removed from the feed, stream (8), to the gasifier (G). The preferred processing conditions remain as before. For hydrogenation (Fig. 1) the gas processing unit (I) consists of a conventional shift converter for producing a hydrogen rich gas, stream (13), for hydrogenation at hydrogenation unit (E), and acid gas, stream (it), removal operations to remove CO2 and sulfur containing gases. In acid hydrolysis (Fig. 8) this gas processing operation consists of acid gas removal operations, where the sulfur removal operation produces sulfuric acid, stream (IA), and/or CO2, stream (1B), for use in the acid hydrolysis operation. In the coking operation the gas processing operation may consist of acid gas removal operations to produce a high calorific gas for plant fuel and for sale. The following Examples demonstrate the kinds of results that are obtainable with the feeds and conditions cited. EXAMPLE 1 A series of experiments were performed to model the integrated ionic liquefaction process from slurry mixing to hydrogenation. All experiments were performed in a single pass batch mode. The experiments were performed in a semi-batch liquefaction system. The major components of the system are the gas delivery system, reactor system and gas measurement system. The gas delivery system consists of a Whitey gas compressor. The primary components of the 3 reactor system is a 0.001 m (1.0 liter) magnedrive Autoclave manuactured by Autoclave Engineers. A knock-back condenser is used to minimize liquid loss from the reactor. System pressure is maintained using a Grove dome-loaded back-pressure regulator. The gas measurement systems consist of a Rockwell diaphragm meter for total gas volume, and a Cane Series S chromatograph for on-line analysis of water gas shift components, light hydrocarbons and Argon tracer. To produce sufficient material for all steps for the process, a series of four identical experiments were performed. Conditions for these experiments are listed in Table 1. The Texas Lignite used in these experiments was obtained from a single large parent coal sample. At the completion of each run, the reactor product was removed and filtered Samples were removed from the reactor product of the last run for selective solvent extraction analysis. The procedure is an empirical method to determine the quality of the product. Ten grams of the reactor product is extracted with three 150 ml washes of tetrahydrofuran (Time) in centrifuge bottles. Each wash is centrifuged and the liquid decanted. The solids are dried and weighed. The TI1F insoluble material correspond to unreacted carbonaceous material The THF soluble fraction is rotovaped and the THF removed. The remaining liquid is washed with toluene in the same manner as the THF. The toluene insoluble solids correspond to high molecular weight material. The toluene soluble material is rotovaped and washed with pentane in the same manner as the THF. The pentane insoluble material is lower molecular weight material and some polar material. The pentane soluble material corresponds to oils or low boiling (7600F) compounds. The MAF (moisture and ash free) and DAbIF (dry mineral matter free) yield structure was: FAF DDIF (wt t) (wt t) Yield 82.9 95.8 Toluene Ins. 29.2 29.2 Pentane Ins. 36.7 36.7 Pentane Sol 16.7 29.6 Gas 0.3 0.3 The filter cakes were collected and analyzed for moisture, ash, and sodium content. The filtrate from each run was individually distilled to remove one half of the material (by weight). The overheads of each run were collected and mixed together. The bottoms were also collected and mixed together. A summary of the material balance for these experiments is shown in Table 2. The data indicate that over 80% of the MAF coal is converted to a liquid product (81.1% calculated from cake weight, 85.68 calculated from filtrate weight). This step also indicates that 85.6t of the MAF coal can be separated by filtration for the distillation. A sample of the bottoms was upgraded and stabilized in a hydrogenation experiment. The bottoms (494 grams), Harshaw Catalyst 601-T (50 g), and Na2St9H2O (25 g) were mixed together in the reactor. The Harshaw catalyst is a Cobolt-Molybdenum catalyst. The Na2S was added to increase the activity of the catalyst. The hydrogenation experiment was performed at 4000C (725 0F) for a period of 60 minutes at temperature. A system pressure of 13.90 MPa (2012 psia) was maintained with gas flow of 5 x 10 5m3/s. A 95:5 mixture of hydrogen and argon was used as the gas. At the completion of the experimental run the reactor product was centrifuged and the decants collected. The bottoms were mixed and analyzed for moisture, ash, ash components, and by SSE. The decant was analyzed by SSE to determine pentane soluble yields. The physical properties of the topped crude were greatly changed by hydrogenation, as shown in Table 3. Solvent extraction showed that all the THF insoluble material was in the decant bottoms. Overall there was a 55% reduction in the THF insolubles, a 100% reduction in the toluene insolubles, a 65% reduction in pentane insolubles and a 56% increase in the pentane soluble. For the ionic liquefaction process this would result in 56.5 Ib. of Pentane Solubles for every 100 pound MAF coal. The gas stream from the reactor was analyzed for gas production and for the quantity of hydrogen consumed. The gas analysis showed that there was 1.7 g-mole of hydrogen consumed with corresponds to1.4 wt. z of the MAE coal fed. TABLE 1 Reactor Contents: Lignite .180 kg H20 .025 kg NaOH .018 kg m-Cresol .260 kg Tetraline .060 kg Naphthalene .020 kg l-Methylnaphthalene .010 kg gamma-picoline .010 kg Gas Type: CO/Ar (95:5) Pressure: 1500 psia (10.35 MPa) Temperature: 335 0C (6350F) Gas Flow Rate: 3 l/min (5 x 10 3 Residence Time: 60 minutes Lignite Analysis: Component Amt (wt %) H2O 28.2 Ash 6.0 C 48.8 H 3.3 N 0.7 S 0.6 0 12.4 Heating Value (Btu/lb) 8,550 TABLE 2 ANALYSIS OF DISTILLATION BOTTOMS AND HYDROGENATED PRODUCT Hydrogenated Bottoms Product Kinematic Viscosity 1902 cSt 18.5 cSt Conradson Carbon 17.10 wt z 9.58 wt % API Gravity -1.7 oAPI 7.2 oAPI THF Ins. 19.4 g 8.7 g Toluene Ins. 73.1 g 0,0 g Pentane Ins. 127.0 g 44.9 g Pentane Sol 268.7 g 418.1 g Gas -- 4.2 g TABLE 3 SUMMARY OF PRODUCT GENERATION RUNS* Grams Grams Grams Grams DMMF Organic In Organic Run No. tAF Coal Solvent Cake Filtrate Balance 189 113.5 360 26.6 453.8 101.5 190 113.8 360 15.2 456.2 100.5 191 111.3 360 20.2 472.6 104.5 192 112.6 360 23.2 443.7 98.8 TOTAL 451.2 1440 85.2 1826.3 101.1 *All values normalized to 100% mass balance to account for transfer losses. EXAMPLE II A suite of six coals were experimentally tested to determine suitability as ionic liquefaction feedstocks. The coals ranged in rank from an Australian brown coal to a high volatile B bituminous. Analyses of the suite coals are given in Table 4. Experiments were performed at the conditions noted in Table 2, except for the experiments using the Australian brown coal which were performed at lower temperature because of the high moisture content of the coal. In general all of the coals tested give satisfactory yields as shown in Table 5 although the lower rank coals seem to be more suitable ionic liquefaction feedstocks because of large oil and gas yields. TABLE 4 EXP COAL RANK ANALYSIS %H O % ASH %C %H %N %S %O Btu/lb 134 Morwell Mine, Latrobe Valley, Aus Brown 53.6 1.4 31.6 2.2 0.3 0.1 10.8 5,280 135 Morwell Mine, Latrobe Valley, Aus Brown 44.4 1.6 37.9 2.6 0.4 0.2 12.9 6,350 140 Colstrip Montana Subb 12.2 14.3 55.8 3.7 0.8 0.8 12.4 9,5000 141 Illinois No. 6 hvBb 4.2 33.8 47.3 3.5 1.0 3.8 6.4 8,490 143 Black Mesa Field Arizona hvAb 6.1 8.6 66.0 4.7 1.0 0.6 13.0 11,550 175179 Big brown Mine, TX Lignite 28.2 6.0 48.8 3.3 0.7 0.6 12.4 8,550 188 South Africa hvBb 1.4 27.1 56.8 4.0 1.1 1.3 8.3 10,030 TABLE 5 RUN FEED TEMP DMMF YIELDS, WT%* OC Total PA A O+G yield 134 Brown Coal 310 96.7 28.6 27.1 41.0 135 Brown Coal 322 94.2 24.4 27.8 42.0 140 Colstrip 335 75.5 27.8 31.9 15.8 141 Illinois No. 6 335 94.8 45.2 25.7 23.9 143 Black Mesa 335 87.1 31.9 36.1 19.1 188 South Af. 335 86.4 56.5 20.9 9.0 175- Texas 179 Lignite 335 95.3 28.5 35.4 31.5 * 80% of Na is Formate An alternative process has been developed to solvate carbonaceous material with a solvent/solute system, to filter the solid ash and undissolved organic material from the liquid, and to distill a recycle solvent fraction leaving behind a solid with reduced sulfur and ash content. The solid product formed can be used as a clean burning boiler fuel as a replacement for oil or coal. Use of this product has economic advantages over oil and eliminates many of the enviromental problems associated with the burning of coal, such as sulfur emissions and ash handling problems. This is accomplished with a solution of phenols, alkali, and water mixed with the carbonaceous material. The slurry is heated with or without the presence of gas or any gas partial pressure. The slurry is heated in the range (300 C to 360 C) for a period of from 60 minutes to 90 minutes. The pressure of the system is that exerted by the liquids present in the system at the temperatures employed, which will typically be in the range of from 2.07 to 10.3SMPa (300 to 1500 psia). After solvation of the carbonaceous material in an ionic liquefaction reactor the slurry is removed and the temperature and pressure reduced. The slurry is filtered to remove the ash and undissovled organic material that has not dissolved in the solvent/solute system. A large portion of the ash will be removed in this step including a majority of the inorganic sulfur. The filtrate is then transferred to a liquid extraction stream to remove dissolved ash components and alkali. The ash and alkali are removed by contacting the filtrate with an acid such as carbonic (H2CO3), HCl, H2SO4 or the like. Dissolved inorganic compounds will be removed because of their greater affinity for the aqueous layer as opposed to the organic filtrate. The acid will replace alkali atoms present as alkali-organic salts with hydrogen with the alkali being soluble in the aqueous layer. The liquid extraction stream will include a final water wash of the filtrate to remove any remaining inorganic constituents as well as acid components still present in the organic layer. The extracted and washed filtrate is then distilled to recover the solvent. The overheads from the distillation tower are used as recycle solvent. The distillation temperature is raised sufficiently to remove the fraction boiling under about 275 C, more preferably under about 300 C and most preferably under 325 C. The bottoms from the distillation are collected and cooled to form a solid product which has reduced sulfur and ash. Sulfur can be removed in a variety of ways. The inorganic fraction which is mainly pyrite, will be removed as an insoluble material in the filtration step after the solvation of the carbonaceous material. Organic sulfur is removed by reaction with alkali and base to form such compounds as alkali sulfides, sulfites, and sulfates. The resulting sulfur compounds are either removed during the filtration step or during the acid extraction of the filtrate depending on the solubility of the species. A sulfur scavenger may also be added to the ionic liquefaction step which is selected to react with sulfur compounds to form insoluble species. Referring to the schematic diagram in Fig. 10, a process is shown which favors the production of a product which is solid at room temperature and is useful as a fuel substitute. The feed preparation at (Al), as previously described, comminutes the carbonaceous material (stream 100) by conventional means, as previously described; and adds a water-alkali mixture, stream (300); and recycle polar solvent, stream (200) containing greater than 50% by weight of phenolic species. The comminution process may again be accomplished either dry or wet. If performed wet, then the recycle polar solvent may be used as the wetting agent. The carbonaceous feed is, as before, comminuted to 100 percent minus 74 micron (200 mesh) particle size, more preferably to 100 percent minus 147 microns (100 mesh) particle size, and most preferably to 100 percent minus 350 microns (40 mesh) particle size but in any event must be in a form which will enable the requisite solubilization for the ionic liquefaction to proceed. Using 1000 parts by weight of the stream of carbonaceous material introduced into the feed preparation at (Al) as an example, the preferred amount of polar recycle solvent for the required solubilization to proceed, introduced by stream (200) is between 1500 and 3500 parts by weight depending on the prepared form of the carbonaceous material, with 3000 parts by weight of solvent the most preferred amount. The polar recycle solvent contains preferably greater than about 50% by weight phenolic compounds, and more preferably greater than 60% by weight phenolic compounds. The preferred amount of alkaline material in stream (300) is selected to be that amount which is to produce the desired results. It has been found under the conditions disclosed herein that between about 25 parts and 400 parts by weight is effective with the more preferred amount being between about 25 and 150 parts by weight, and depending on the kinds and amounts of phenolic materials employed, the preparation of the carbonaceous material and the conditions selected, the most preferred amount is about 50 parts. The amount of water in stream (300) should preferably be sufficient to maintain the alkaline material in the ionic form in solution at the described ionic liquefaction conditions. The amount of water in stream (300) is between about 25 and 400 parts by weight, more preferably between about 50 and 250 parts by weight, and most preferably between about 100 and 200 parts by weight. Make up alkali, (stream 300A) may be added, as before. The liquefaction section (B1) is the same or similar to that shown in (B) Fig. 1. A high pressure slurry pump (not shown) can be used to pump the slurry, stream (400) and to bring the slurry to the desired system pressure. The slurry at this point may or may not have a synthesis gas added. The reactions can proceed without the added gas. If added gas is used, the gas is introduced before further processing. The composition of the synthesis gas stream can be as previously described. Preferred gas treatment rates are also as previously described. The reaction mixture is then brought to the desired temperature. The reaction mixture stream is then transferred to the liquefaction reactor (Bl) and held at the desired reaction temperature for sufficient time to permit the ionic liquefaction process steps to take place to the desired extent. The solubilization and ionic reaction steps, but not the solvent regeneration or ionic species upgrading, take place in this reactor. Temperature and pressure should be optimized for individual carbonaceous feeds. Preferred reactor residence times are from about 2 to about 120 minutes, and most preferably from about 15 to 45 minutes depending on reactor design. Tie preferred temperature range is from about 2500C to 36O0C (4820F to 6800F). System pressure should be established and maintained at such a predetermined level so as to keep sufficient water in the liquid phase to maintain ionic alkaline species in the liquid phase, rather than as salts, for a predetermined portion of the reaction. The preferred pressure range is from about 0.69MPa to 13.80MPa (100 psia to 2000 psia)1 and the most preferred pressure range is about 2.76MPa to 8.28MPa (400 psia to 1200 psia). The reaction products leaving the reactor (B1), in stream (606) of Figs. 9 and 10, are then separated into. component streams in the separation system, process (C1) of Fig. 10. The gas-slurry separators employed serve to separate the majority of the slurry product from gaseous products. The separator pressure is maintained at reduced pressure1 preferably between about 0.69 and 3.45 MPa (100 - 500 psia), and the temperature is kept at conditions sufficient to keep most of the volatile organic compounds in the liquid phase, but much of the remaining water as vapor. The preferred temperature range is 200 to 300 C (392 to 572 F); the most preferred temperature range is 200 to 250 C (392 to 482 0F). The slurry phase residence time is preferably less than 30 minutes, more preferably less than 15 minutes, and most preferably less than 5 minutes. The vapor stream (601) from the gas-slurry separator goes to a carbon dioxide separation step (El) then to sulfur recovery (K1) for separation into a flue gas stream (FG) and a sulfur stream (so). As previously described, leaf filters, candle filters, hydroclones, or comparable equipment can be used for the separation of the solids as well as separation by processes such as solvent deashing, or critical solvent extraction. The purpose of the filtration is to separate undissolved carbonaceous material1 ash and alkaline salts from the ionic liquefaction products which are liquid under the conditions employed. Due to the hygroscopic nature of the alkaline salts1 much of the water present will also be separated with the filter cake. The preferred temperature for the separation step (Cl) is preferably between about 150 and 100 C (302 and 212 F). The pressure is maintained at a sufficient level to obtain efficient separation, preferably between about 0.34 and 1.03 MPa (50 to 150 psia). In the separation step, the solids content of the liquid in the stream (609) is typically reduced to less than about 1.0 percent by weight, the mineral matter content to less than about 0.5 percent by weight1 and the alkaline species content to less than about 0.25 percent by weight. The values obtained are dependent upon factors such as the degree of comminution used in feed preparation, ionic liquefaction conditions, and filter equipment design. The separation train should preferably be operated to reduce the alkaline compound content of the stream (609), of Fig. 10, below about 0.25 percent by weight, more preferably below about 0.15 percent by weight, and most preferably below about 0.10 percent by weight. The liquid stream (609) of Fig. 10 with alkaline content less than .25 percent will have the alkaline content reduced further by washing in a solvent extractor system, such as shown in Fig. 11. The liquid stream (609) is contacted with the solution stream (610) of Fig. 10 containing added fresh water in a countercurrent method while bubbling CO2, steam (611) through the solution in a mixing vessel (M1) as shown in Fig. 11. The resulting liquids stream (615) is passed to a liquid/liquid settler (S1) for separating the liquids stream into an organic component and an aqueous component. The organic component stream (613) of Fig. 11 will have the alkaline and ash concentration reduced to less than 0.05%, more preferably to less than .02% and most preferably to less than .005%. In the preferred embodiment of the process, the liquid stream (609) and water stream (610) are contacted in mixing vessel (M1) in a feed ratio of one part filtrate (609) to 5 parts H2O, more preferably one part feed to 3 parts H2O, and most preferably one part feed to one part H2O. The preferred temperature is about 100 C (212 F), more preferably less than about 80 C (1760F)-and most preferably less than about 500C (1220F). The liquid streams and CO2 will preferably be mixed for a period of 15 minutes, and more preferably for a period of between 5 and 10 minutes. The carbon dioxide when bubbled through the aqueous solution produces carbonic acid which replaces the alkali in the organic solution with hydrogen. An organic-rich stream (613) is separated in separator (S1) and sent to a distillation unit (I1) (Fig. 10) where the organic rich stream is fractionally distilled. A light oil fraction is collected as a product stream (626). Distillate product boiling between 200 C and 325 C (392 F and 617 F) is collected as a recycle solvent stream (200) and returned to the feed preparation (Al). The high boiling fraction is collected and cooled to about 25 C. This is a product stream (627), which at room temperature, is a pseudo-plastic solid with an ash content below 1%, more preferably below 0.3%, and a sulfur content below 1.0%, more preferably below 0.5%, and most preferably below 0.3%. The solids rich filter cake (607) is washed with a water stream at solids washing unit (F1). The water stream is preferably at a temperature of 80 C. and at a ratio to the cake of 5 to 1, preferably 3 to 1, and most preferably 1 to 1. The water stream will remove liquid organics associated with the filter cake and the water-organic stream (610) and is sent to the solvent extraction system (D1). Referring further to the integrated process in Fig. 10, the washed filter cake produced in the solids washing unit (F1), is used as a feed stream (614), to a gasifier1 unit (G1). In the configuration shown in Fig. 10 the preferred gasification operation is a dry ash partial combustion process to produce a synthesis gas rich in carbon monoxide and hydrogen. The preferred H2/CO ration of gas leaving the gasifer, stream (615) is between 0.5 and 2.0. Acid gases and light hydrocarbon gases are separated from the synthesis gas by standard operations in gas processing operations, unit (H1). The ash from the gasifier, stream (616), is sent to an alkali recovery unit (J1) where a portion of the alkaline compounds are separated from the ash stream by hot water extraction, and recycled, stream (300), to the preparation unit (Al), and the residual ash removed, stream (628). The preferred recovery of alkaline compounds is greater than 50 percent by weight, more preferably greater than 75 percent by weight and most preferably greater than 90 percent by weight. The preferred temperature of the hot water used for the extraction is between 25 and 100 C (77 and 212 F), and more preferably between 50 and 100 C (122 and 2120F). The preferred treatment rate is less than 4 kg H2O/kg ash, more preferably less than 2 kg H2O/kg ash, and most preferably less than 1.1 kg ES2O/kg ash. EXAMPLE III A series of experiments were performed to model the ionic liquefaction process for the production of a solid product. The experiments were performed in a single pass batch mode modeling all process steps from slurry mixing to recovery of solid product. The liquefaction step was performed in the system described in Example I. The experiments were performed using several types of carbonaceous feedstocks. An analysis of each feedstock is presented in Table 6. The ionic liquefaction conditions for each feedstock are listed in Table 7. The product from the ionic liquefaction reactor for each experiment was removed from the reactor. This product was filtered to remove ash and unreacted carbonaceous material. The filter cake was weighed and then washed with 180 F. water to remove excess liquid product. The washed filter cake was dried and the weight recorded. Alkali and ash material were removed from the filtrate by extraction with a 10% hydrocholoric acid solution. The extractions were performed in a separatory funnel by mixing the filtrate three separate times with fresh acid solution and collecting the filtrate (organic fraction). The extracted filtrate was distilled to remove the material boiling below 3000C The bottoms from this distillation was the solid product and was collected. Any distillate collected in excess of the solvent was coal-derived distillate and recoverable as a liquid product stream. Results from this series of experiments are presented in Table 8. These data indicate that the process yields a solid product with greatly reduced ash and sulfur content and an increased heating value. TABLE 6 ANALYSES OF CARBONACEOUS FEEDSTOCKS Composition (wt%) as fed HHV Experiment Carbonaceous Number Feedstock H2O Ash C H N S O BTU/lb 241 North Dakota 31.75 8.63 43.06 2.53 0.60 0.80 12.63 7,618 Lignite 243 Colorado Wadge 7.94 8.27 64.47 4.34 1.19 0.49 13.30 11,248 Bituminous 256 Pecan Wood 4.43 0.68 47.10 5.64 0.44 0.25 41.46 8,336 262 20 Weight Bond 5.81 10.62 37.01 5.84 0.10 0.08 43.86 6,414 Minimal Rag Content 266 60 Weight paper 7.49 0.90 42.24 6.54 0.05 0.13 42.65 7,709 Kraft Wrapping 267 Rags 0.16 0.91 55.67 4.82 0.36 0.05 36.31 9,109 65% Polyester 35% Cotton TABLE 7 REACTION CONDITIONS 20 North Weight Bond 60 Rags Dakota Colo. Pecan Mininal Weight Paper 65% Polyester Carbonaceous Feed Lignite Wadge Wood Rag Content Kraft Wrapping 35% Cotton Experiment &num; 241 243 256 262 266 267 Amount (g) 150 150 100 100 100 100 Conditions: m-cresol (g) 360 360 240 240 240 240 Tetralin (g) 90 90 60 60 60 60 NaOH (g) 3.75 3.75 10 5 10 10 NaCOOH (g) 3.75 3.75 --- --- --- -- H2O 10 20 40 20 25 25 Pressure (MPa) 10.97 11.21 8.83 10.83 11.39 8.45 Temperature ( C) 340 360 340 340 340 340 Residence Time (min) 90 90 90 90 90 90 TABLE 8 EXPERIMENTAL RESULTS 20 North Weight Bond 60 Rags Dakota Colo. Pecan Minimal Weight Paper 65% Polyester Carbonaceous Feed Lignite Wadge Wood Rag Content Kraft Wrapping 35% Cotton Experiment &num; 241 243 256 262 266 267 Amount DMMF Feed (g) 85.9 124.2 92.3 83.6 91.6 97.1 wt % Solid Product 48.3 34.7 60.4 47.3 52.6 32.7 wt% S 0.22 0.23 0.02 0.00 0.20 0.00 wt% Ash 0.19 0.03 0.05 0.20 0.10 0.80 wt% Distillate 15.1 21.8 33.1 46.9 43.7 44.3 wt% Unreacted 39.4 41.3 6.0 ---- 0.0 22.7 EXAMPLE IV A suite of coals have been examined to determine their suitability as feedstocks to produce a solid product from the ionic liquefaction process. Analysis of the suite of coals appears in Table 9. The experiments were performed as described in Example 3 under the ionic liquefaction conditions presented in Table 10. In general, all of the coals tested give satisfactory yields, and produce a product with reduced sulfur and ash content. Results appear in Table 11. EXAMPLE V An experiment was performed using a Texas lignite. Analyses and reaction conditions of the experiment appear in Table 9 and 10. Tlle experiments were performed as described in Example 3 except a different extraction procedure was used. In this example the filtrate was mixed with water (1:1.5) and stirred. Carbon dioxide gas was bubbled through the agitated solution for a period of one hour. After one hour the organic fraction was removed in a separatory funnel and distilled. The results of experiments are given in Table 11. TABLE 9 ANALYSES OF COAL FEEDSTOCKS Composition (wt%) as fed HHV Experiment Carbonaceous Number Feedstock H2O Ash C H N S O BTU/lb 241 North Dakota 31.75 8.63 43.06 2.53 0.60 0.80 12.63 7,618 Lignite 243 Colorado Wadge 7.94 8.27 64.47 4.34 1.19 0.49 13.30 11,248 Bituminous 259 Rawhide Mine 25.55 6.69 49.05 3.18 0.61 0.51 14.40 8,901 Gilette, WY 261 Usibelli Mine 20.76 5.65 49.94 4.34 0.66 0.23 18.65 8,836 Nenana Field Alaska 247 Ohio No. 6 6.15 9.75 66.48 4.89 1.21 2.05 9.47 11,850 242 Illinois No. 6 7.15 13.10 61.36 4.18 1.03 3.27 9.91 11,041 245 Texas Lignite 25.91 10.58 47.98 2.67 0.79 0.67 11.40 7,819 TABLE 10 REACTION CONDITIONS North Dakota Colo. Rawhide Carbonaceous Feed: Lignite Wadge Mine Usibelli Experiment &num; 241 243 259 261 Amount (g) 150 150 150 150 Conditions: m-cresol (g) 360 360 360 360 Tetralin (g) 90 90 60 60 NaOH (g) 3.75 3.75 15 15 NaCOOH (g) 3.75 3.75 H2O (g) 10 20 10 20 Pressure (MPa) 10.97 11.21 10.76 9.31 Temperature (OC) 340 350 340 340 Residence Time (min) 90 90 90 90 TABLE 10 (continued) REACTION CONDITIONS Ohio Illinois Texas Carbonaceous Feed: No. 6 No. 6 Lignite Experiment &num; 247 242 245 Amount (g) 150 150 150 Conditions: m-cresol (g) 360 360 360 Tetralin (g) 60 60 90 NaOH (g) 3.75 3.75 3.75 NaCOOH (g) 3,75 3.75 3.75 H2O (g) 20 20 10 Pressure (MPa) 6.49 12.56 9.59 Temperature (OC) 360 360 360 Residence Time (min) 90 90 90 TABLE 11 EXPERIMENTAL RESULTS North Dakota Colo. Rawhide Carbonaceous Feed: Lignite Wadge Mine Usibelli Experiment &num; 241 243 259 261 Amount DMMF Feed (g) 85.9 124.2 94.6 105.8 wt 8 Solid Product 48.3 34.7 41.8 41.5 wt% S 0.22 0.23 0.00 0.02 wt% Ash 0.19 0.03 0.05 0.20 wtt Distillate 15.1 21.8 7.6 11.5 wt% Unreacted 39.4 41.3 49.5 47.2 TABLE 11 (continued) EXPERIMENTAL RESULTS Ohio Illinois Texas Carbonaceous Feed: No. 6 No. 6 Lignite Experiment &num; 247 242 245 Amount DMMF Feed (g) 124.5 113.8 150 wt % Solid Product 56.7 62.6 53.6 wt% S 0.71 1.31 0.02 wt% Ash 0.10 0.80 0.06 wt% Distillate 7.5 4.0 0.0 wt% Unreacted 33.0 29.9 46.4 EXAMPLE VI A hvBb coal obtained from the Ohio No. 6 Seam is preprocessed in a conventional gravity separation, screening and drying process, and is pulverized to a top size of about -200 mesh (-74 microns). A semi-batch liquefaction unit comprising a gas delivery system, a reactor system, and a gas measurement system is charged with 100g. of pulverized coal and 360g. of m-cresol. The semi-batch coal liquefaction unit is designed for continuous flow of gas, and for batch injection of solid-liquid slurries. Gas is fed to the liquefaction unit from pressurized gas bottles which are premixed with 5% argon and 95% carbon monoxide. The gas delivery system is equipped with pressure regulators, and flow controllers to maintain 1012 psia (6.98 MPa) at 0.1263 gram moles per minute gas flow rate (of CO). The reactor system consists of a 316 stainless steel, one-liter Magnedrive Autoclave manufactured by Autoclave Engineers, Erie, Pennsylvania, and an iron-constantan thermocouple connnected to an Omega Model 400A temperature indicator. The heater temperature is controlled by a Fenwall Series 5501552 temperature controller. Gas flow enters the reactor through the stirrer and exits through a knockback condensor consisting of a 3/4-inch O.D. stainless steel tube in a water jacket. The gas measurement system consists of a Rockwell Model S-200 diaphragm meter for measurement of total gas olume, a Carle Series "S" chromatograph for analysis of carbon monoxide, carbon dioxide, hydrogen and argon tracer, and a ewlett-Packard 3390 integrator to calculate and print the gas composition in mole percents. The semi-batch liquefaction reactor system is pressure tested at 1012 psia (6.98 MPa) with helium and then the premixed argon and carbon monoxide gas is introduced and the reactor is heated to 3000C (573K). After the reactor temperature and pressure are maintained for the desired reaction time, 120 minutes in this instance, the heater jacket is removed and the autoclave is cooled using forced air convection. The solid and liquid components are removed from the reactor and mixed in a high speed blender. Samples are removed from the blender and placed in 250 ml. centrifuge tubes. The samples are subjected to an empirical selective solvent extraction procedure using tetrahydrofuran (THF), toluene1 and pentane to determine total conversion, preasphaltenes, asphaltenes, and oil plus gas. The yield and product structure are defined by: XV. Yield = 100 - Grams MAF THF Insoluble Material (100), [=] wt% Grams MAF Coal XVI. Preasphaltenes (PA) = Grams MAF Toluene Insoluble Material (100), Grams MAF Coal [=] wt z XVII. Asphaltenes (A) = Grams MAF Pentane Insoluble Material (100), Grams MAF Coal [= wt % XVIII. Oil plus Gas (O + G) = Yield - PA - A, [=3 wt % The results of the selective solvent extraction procedure are shown in Table 12. Results of the gas analysis showed 93.52 percent carbon monoxide, 0.66 percent hydrogen and 0.3 percent carbon dioxide. Carbon monoxide conversion is calculated from analysis, over time, of the exit gas, and plotted with temperature and the results shown in Fig. 12, wherein: Qrepresents CO conversion, calculated from CO in the exit gas; 0represents CO conversion, calculated from CO2 in the exit gas, using water gas shift storchiometry; and represents, CO conversion, calculated from 112 in the exit gas, using water gas shift storchiometry. EXAMPLE VII The foregoing procedure is repeated using 100g. of the hvBb Ohio No. 6 coal, 360g. of m-cresol, 40g. of water, 1012 psia (6.98 MPa), 3000C (573K), of 95 percent carbon monoxide and 5 percent argon1 0.01992 gram moles per minute gas flow rate (of CO), for 120 minutes. The semi-batch liquefaction unit is charged, heated and the products analyzed as previously described. The results of the selective solvent extraction procedure are shown in Table 12 and results of the gas analysis showed 92.45 percent carbon monoxide, 0.81 percent hydrogen, and 1.15 percent carbon dioxide. Carbon monoxide conversion is calculated and plotted, as before, and the results shown in Fig. 13. EXAMPLE VIII The foregoing procedure of Example VI is repeated using 100g. of the hvBb Ohio No. 6 coal, 360g. of m-cresol, 49g. of water, 25g. of potassium hydroxide, 1012 psia (6.98 MPa), 300 C (573K), of 95 percent carbon monoxide and 5 percent argon, 0.01992 gram moles per minute gas flow rate (of CO), for 120 minutes. The semi-batch liquefaction unit is charged, heated, and the products analyzed as previously described. The results of the selective solvent extraction procedure are shown in Table 12, and results of the gas analysis showed 54.54 percent carbon monoxide, 19.44 percent hydrogen, and 21.44 percent carbon dioxide. Carbon monoxide conversion is calculated and plotted, as before, and the results shown in Fig. 14. TABLE 12 MAF Example Solvent/Solute System Conversion Organic Phase Inorganic Phase Water Solubilizing Alkali/Alkaline Agent Earth Compound Wtt VI 360g. m-cresol 0 0 40.0 VII 360g. m-cresol 0 40g. 39.0 VIII 360g. m-cresol 25g. KOH 40g. 82.5 EXAMPLE IX The foregoing procedure of Example VI is repeated using 100g. of a hvCb Colorado Wadge coal, 360g. of a synthetic recycle solvent consisting of 270g. of m-cresol, 60 g. of 1,2,3,4-tetrahydronaphthalene, 20 g. of naphthalene, and lOg. of l-methylnaphthalene, 40g. of water, 1012 -psia (6.98 MPa), 3000C (573K), 0.5 SLM of 95 percent carbon monoxide and 5 percent argon1 and 120 minutes. The semi-batch liquefaction unit is charged, heated, and the products analyzed as previously described. The results of the selective solvent extraction procedure are shown in Table 13. EXAMPLE X The foregoing procedure of Example VI is repeated using 100g. of a hvCb Colorado Wadge coal, 40g. of water, 15g. of sodium hydroxide, and the temperature, pressure1 gas composition and flow rate, and residence time of Example IX. The semi-batch liquefaction unit is charged, heated, and the products analyzed as previously described. The results of the selective solvent extraction procedure are shown in Table 13. EXAMPLE XI The foregoing procedure of Example VI was repeated, except 15g. of sodium carbonate was used replacing the 15g. of sodium hydroxide. The semi-batch liquefaction unit is heated, charged, and the products analyzed as previously described. The results of the selective solvent extraction procedure are shown in Table 13. TABLE 13 Example Solvent/Solute System Organic Phase Inorganic Phase MAF Solubilizing Alkali/Alkallne- Conversion Agent Earth Compound Water Wt z IX 360g. synthetic solvent 0 40g 22.8 X 360g. synthetic solvent 15g NaOH 40g 60.5 XI 360g. synthetic 15g Na2CO3 40g 60.1 solvent The results in Tables 12 and 13 demonstrate the significantly improved results obtained by practice of the present invention. Table 13 shows that the presence of the organic phase solubilizing agent, m-cresol1 in the absence of the inorganic phase constituent as Example VI, yields a MAF conversion of 40 wt %. In the case of Example VII with the addition of water1 the MAF conversion is 39 wt %, which is virtually unchanged from Example VI. In Example VIlI1 under operating conditions of Examples VI and VII, the synergistic effect of the alkali/alkaline-earth constituent is observed as the yield is increased to 82.52 MAF wt t. In Examples IX, X and XI, the organic phase solubilizing agent is a synthetic solvent which is considered to represent a recycle stream in a continuous liquefaction facility. The MAF wt z yields for Examples X and XI, when compared to Example IX, show the increased synergistic effect obtained by the combination of the inorganic and organic phase constituents. EXAMPLE XII The foregoing procedure of Example VI is repeated using 180g. of a hvCb Colorado Eagle No. 5 coal. The organic fraction of the solvent/solute system is 360g. of synthetic solvent consisting of 160g. of m-cresol1 160g. of tetrahydronaphthalene, 20g. of naphthalene, 10 g. of l-methylnaphthalene, and 10g. of gamma-picoline. The inorganic fraction of tie solvent/solute system is 30g. of water 18g. of NaOH, 4.5g. of Na2CO3, and 30g. Na2S 9H2O. The feed materials are reacted at 3400C, 1312 psia (8.99 MPa), 0.5 SLM of 95 percent carbon monoxide and 5 percent argon for 30 minutes. The results of the selective solvent extraction prpcedure are shown in Table 14. EXAMPLE XIII The foregoing procedure of Example XII is repeated, except that the organic fraction of the sovvent/solute system is 360g. of a synthetic solvent consisting of 260g. of m-cresol, 60g. of tetrahydronaphthalene, 20g. of naphthalene, 10g. of l-methylnaphthalene, and 10 g. of gamma-picoline. The results of the selective solvent extraction procedure are shown in Table 14. TABLE 14 Example MAF Conversion XII 66.3 XIII 73.8 Examples XII and XIII show that acceptable liquefaction yields can be obtained when the organic fraction of the solvent/solute system consists of a mixture of alkaline/alkaline-earth metal compounds. They also show the importance of phenolic compounds in the organic fraction of the solvent/solute system. In Examples XII, where m-cresol is 44.4 percent of the organic fraction1 the yield is 66.27 percent. In Example XIII, m-cresol is increased to 72.2 percent of the organic fraction, the yield is increased to 73.82 percent. The mineral contents of coals used in Examples VI through XIII are presented in Table 15. TABLE 15 Minerals (Wt%) Coal Fe203 Na2O K2O Al2O3 Sio2 CaO MgO Eagle &num;5 7.21 1.89 1.47 24.5 55.2 3.83 1.79 Ohio X6 18.30 .81 2.70 25.1 51.2 - Wadge 4.09 .62 .84 27.4 60.5 4.23 .79 The foregoing description of several embodiments of an integrated ionic liquefaction process can1 of course, be modified by adding steps or combining operations if desired to achieve different specific results from those described without departing from the spirit of this invention and the scope of the attached claims, which are limited only by the prior art application to this invention.

Claims: CLAIMS 1. A method of separating oil product from a single stage carbonaceous liquefaction process stream following ionic liquefaction in which polar solvents have been utilized for liquefaction together with alkali metal or alkaline earth compounds or mixtures thereof 1 characterized by the steps of: (1) removing non-condensible gases from said process stream; (2) removing immiscible water from said process stream; (3) removing alkaline containing solids from said process stream; (4) recovering oil products from said process stream by separating by heating said process stream to a temperature of about 177 0C to about 260 C. at atmospheric pressure for a sufficient period of time to separate said process stream into oil product fractions; and t (-5) stabilizing at least one fraction of the oil productsvso recovered. 2. A method according to claim 1, characterized by one or more of the following features: (a) at least one of the oil products recovered comprises aromatic polar solvents suitable for use in a single stage ionic liquefaction process; (b) the step of stabilization includes the catalytic hydrogenation of at least one oil product fraction, by processing said at least one oil product fraction at a temperature between about 340 C and 455 C, and a pressure of from between about 1000 psia to about 2000 psia in the presence of a hydrogenation catalyst for a sufficient period of time to permit polar solvent regeneration and ionic stabilization to proceed to a predetermined degree; ; (c) the step of stabilization includes the acid stabilization of at least one oil product fraction, said acid stabilization comprising contacting said at least one oil product fraction with an acid, preferably carbonic acid, sulfuric acid, hydrochloric acid, formic acid, acetic acid, carbamic acid and mixtures thereof, under conditions which promote hydrogen addition to the ionic organic species present in the oil product fraction1 such oil product fraction preferably being contacted with said acid at a temperature of from between about 1000C to about 205 C, at atmospheric pressure, in amounts from about 1 part to about 10 parts of acid for every part of the oil product fraction treated; ; (d) the step of stabilization includes the coking of at least one oil product fraction by subjecting said at least one oil product fraction to temperatures of from about 4270C to about 5100C for a sufficient period of time to produce gases, liquids and solids; and (e) the solids produced per 1000 parts of oil product fraction are in the range of from about 250 to about 300 parts. 3. A method of regenerating the solubilizing agent from the process stream of a single stage ionic coal liquefaction process wherein said coal liquefaction is accomplished by reaction of the coal with a polar solvent solubilizing agent and an aqueous solution containing inorganic ionic species characterized in that said solubilizing agent comprises in major amount a phenol, polycyclic phenol, substituted phenol, and mixtures thereof, preferably o-cresol, m-cresol, p-cresol, catechol, resorcinol, naphthol, and mixtures and derivatives thereof,and a minor amount of a polycyclic aromatic hydrocarbon, partially-hydrogenated polycyclic aromatic hydrocarbon, fully hydrogenated polycyclic aromatic hydrocarbon having from 1 to 4 carbon rings and mixtures thereof, preferably naphthalene, anthracene, phenanthrene, acenaphthalene, l-methylnaphthalene, 2-methylnaphthalene, tetrahydronaphthalene, gamma-picoline, isoquinoline, dihydronaphthalene, decahydronaphthalene, 9,10-dihydroanthracene, 9,10-dihydrophenanthrene, and mixtures and derivatives thereof, and carbonaceous material-derived solvents; said inorganic ionic species being dissolved in water and including a compound having a cation selected from alkali and alkaline-earth metals, and comprising the steps following said ionic liquefaction of: (1) removing non-condensible gases from said process stream; (2) removing immiscible water from said process stream; (3) removing alkaline compound containing solids from said process stream; (4) recovering oil products from said process stream; (5) separately hydrotreating said oil products; and (6) recovering regenerated solubilizing agent capable of being used directly in a single stage ionic coal liquefaction process. 4. A method of minimizing coking in coal liquefaction process streams containing between about 25 and about 400 parts of alkali metal or alkaline earth salts per 1000 parts of coal, characterized by the steps of (1) flashing the process stream by heating at temperatures sufficient to remove a substantial portion of the water, non-condensible gases and light hydrocarbons; (2) removing immiscible water remaining with the process stream after said flashing; (3) removing alkaline compound containing solids remaining in the process stream after removal of the immiscible water; (4) further separating materials that are gaseous above about 200 C by heating the process stream; (5) catalytically hydrotreating the liquid process stream to insolubilize any alkali-containing constituents; (6) separating the insoluble alkali-containing constituents; and (7) recovering the insoluble alkali-containing constituents by adding water to the separated insoluble alkali-containing constituents. 5. A method of recovering alkali-containing constituents from a coal liquefaction process stream containing more than about 25 parts by weight of alkali metal or alkaline earth salts for every 1000 parts of carbonaceous; material, characterized by the steps of (1) flashing the process stream by heating at temperatures sufficient to remove non-condensible gases, water and light hydrocarbons; (2) removing immiscible water after said flashing; (3) removing solids remaining in the process stream after removal of the immiscible water'. (4) further separating materials that are gaseous above about 190 C by heating the process stream; (5) catalytically hydrotreating the liquid process stream to insolubilize any alkali-containing constituents; (6) separating the insoluble alkali-containing constituents; and (7) recovering the insoluble alkali-containing constituents by adding water to the separated insoluble alkali-containing constituents. 6. A method of recovering solid product from a carbonaceous liquefaction process stream following ionic liquefaction in which a polar solvent or mixtures thereof have been utilized for liquefaction together with alkali metal or alkaline earth compounds or mixtures thereof, characterized by the steps of: (1) removing non-condensible gases from said process stream; (2) removing alkaline containing solids from said process stream; (3) solvent extracting said process stream with an acidic aqueous solvent for a sufficient period of time to separate said process stream into an organic richssfraction and an alkaline rich fraction; and (4) distilling the organic rich fraction under conditions which produce at least a solid product at ambient temperatures and a phenolic rich recycle solvent stream. 7. A method according to claim 6, characterized by one or more of the following features: (a) including the step of introducing carbon dioxide gas into the solvent during the solvent extraction of said process stream; (b) the aqueous acidic solvent includes carbonic acid; (c) the aqueous solvent contains an acid selected from hydrochloric acid, sulphuric acid, carbonic acid, acetic acid, hydrobromic acid, hydrofluoric acid, nitric acid, and mixtures thereof; and (d) the recycle process stream recovered comprises aromatic polar solvents suitable for use in subsequent ionic liquefaction processes 8. A method of producing a psuedo-plastic product which is solid at room temperature from an ionic coal liquefaction process wherein said coal liquefaction is accomplished by reaction of the coal with a polar solvent solubilizing agent and an aqueous solution containing inorganic ionic species characterized in that said solubilizing agent comprises in major amount a phenol, polycyclic phenol, substituted phenol and mixtures thereof, preferably o-cresol, m-cresol, p-cresol1 catechol, resorcinol, naphthol, and mixtures and derivations thereof, and a minor amount of a polycyclic aromatic hydrocarbon, partially-hydrogenated polycyclic aromatic hydrocarbon, fully hydrogenated polycyclic aromatic hydrocarbons having from 1 to 4 carbon rings, and mixtures thereof, preferably naphthalene, anthracene, phenanthrene, acenaphthalene, l-mehylnaphthalene, 2-methylnaphthalene, tetrahydronaphthalene, gamma-picoline, isoquinoline, dihydronaphE-halene, decahydronaphthalene, 9,10-dihydroanthracene, 9,10-dihydrophenanthrene and mixtures and derivatives thereof, and carbonaceous material-derived solvents; said inorganic ionic species being dissolved in water and including a compound having a cation selected from alkali and alkaline-earth metals, andcomprising the steps following said ionic liquefaction of (1) removing non-condensible gases from said process stream; (2) removing alkaline compound containing solids from said process stream; (3) removing dissolved inorganic species from said process stream and recovering an oil product containing stream; and (4) distilling said oil products stream to produce a solubilizing agent capable of being used directly in a ionic coal liquefaction process and a psuedo-plastic product useful as a fuel or coke substitute. 9. The solid product obtained from a carbonaceous liquefaction process stream following ionic liquefaction, characterized in that a polar solvent or mixtures thereof have been utilized for liquefaction together with alkali metal or alkaline earth compounds or mixtures thereof, and recovered following the treatment of said process stream to remove non-condensible gases and alkaline containing solids, said process stream being further treated by solvent extraction with an acidic aqueous solvent for a sufficient period of time to separate said process stream into an organic rich fraction and an alkaline rich fraction; and said organic rich fraction being distilled under conditions which produce a solid product at ambient temperatures and a phenolic rich recycle solvent stream. 10. The product according to claim 9, characterized by one or more of the following features: (a) carbon dioxide gas is introduced into the solvent during the solvent extraction of said process stream; (b) the aqueous acidic solvent includes carbonic acid; (c) the aqueous solvent contains an acid selected from hydrochloric acid, sulphuric acid, carbonic acid, acetic acid, hydrobromic acid, hydrofluoric acid, nitric acid, and mixtures thereof; and (d) the recycle process stream recovered comprises aromatic polar solvents suitable for use in subsequent ionic liquefaction processes. 11. A pseudo-plastic product which is solid at room temperature, obtained from an ionic coal liquefaction process, characterized in that said coal liquefaction is accomplished by reaction of the coal with a polar solvent solubilizing agent and an aqueous solution containing inorganic ionic species wherein said solubilizing agent comprises in major amount a phenol, polycyclic phenol, substituted phenol and mixtures thereof7 preferably o-cresol, m-cresol, p-cresol1 catechol, resorcinol, naphthol, and mixtures and derivatives thereof, and in minor amount comprises naphthalene, anthracene, phenanthrene, acenaphthalene, 1-methylnaphthalene, 2-methylnaphthalene, tetrahydronaphthalene, gamma-picoline, isoquinoline, dihydronaphthalene, decahydronaphthalene, 9,10-dihydroanthracene, 9, 10-dihydrophenanthrene and mixtures and derivatives thereof, and a minor amount of a polycyclic aromatic hydrocarbon, partially-hydrogenated polycyclic aromatic hydrocarbon, fully hydrogenated polycyclic aromatic hydrocarbon having from't; to 4 carbon rings and mixtures thereof, as well as carbonaceous material-derived solvents; said inorganic ionic species being dissolved in water and including a compound having a cation selected from alkali and alkaline-earth metals, where, following ionic liquefaction: non-condensible gases are removed from said process stream; alkaline compound containing solids are removed from said process stream; dissolved inorganic species are removed from said process stream; and oil products contained in said process stream are distilled to produce solubilizing agent capable of being used directly in a ionic coal liquefaction process, and a pseudo-plastic product useful as a fuel or coke substitute. 12. A method of converting carbonaceous materials to liquid products under conditions of temperature and pressure which do nut produce significant thermal bond rupture in the carbonaceous materials characterized by: contacting the carbonaceous material with a solvent/solute system consisting of: (a) an organic phase comprising a solubilizing agent containing more than about 50% by weight of an aromatic phenol, polycyclic phenol, substituted phenol and mixtures and derivatives thereof; and (b) an inorganic phase comprising an aqueous solution of a compound having a cation selected from alkali and alkaline-earth metals; said contacting being conducted at a temperature less ,than about 360 C. and a pressure of at least about 300 psia. 13. A.method according to claim 12, characterized by one or more of the following features: (a) said organic phase solubilizing agent is selected from o-cresol, m-cresol, p-cresol, catechol, resorcinol, naphthol and mixtures and derivatives thereof; (b) said organic phase further comprises one or more organic constitutents selected from polycyclic aromatic hydrocarbons, partially-hydrogenated polycyclic aromatic hydrocarbons and fully hydrogenated polycyclic aromatic hydrocarbons having from 1 to 4 carbon rings, said organic constitutent preferably being selected from naphthalene, anthracene, phenanthrene, acenaphthalene, l-methylnaphthalene, 2-methylnapthalene, tetrahydronaphthalene, gamma-picoline, isoquinoline, dihydronaphthalene, decahydronaphthalene, 9,10-dihydroanthracene, 9,10-dihydrophenanthrene and mixtures and derivatives thereof; ; (c) said organic fraction is in whole or in part derived from liquified carbonaceous material; (d) said compound in the inorganic phase is selected from alkali hydroxides, alkali carbonates, alkali bicarbonates, alkali nitrates, alkali sulfates, alkali sulfites1 alkali sulfides, alkali formates and other alkali salts1 alkaline-earth hydroxides, alkaline-earth carbonates, alkaline-earth bicarbonates, alkaline-earth nitrates, alkaline-earth sulfates, alkaline-earth sulfites, alkaline-earth sulfides, alkaline-earth-formates and other alkaline-earth salts, and mixtures thereof; (e) said compound in the inorganic phase is selected from sodium hydroxide, sodium carbonate, sodium bicarbonate, sodium sulfate, sodium sulfide, sodium nitrate, potassium hydroxide, potassium carbonate, potassium bicarbonate, potassium formate, calcium carbonate and mixtures thereof; ; (f) said compound in the inorganic phase is present in an amount from about 1 part to about 40 parts per 400 parts by weight of the solvent/solute system; (g) the aqueous phase includes water present in anamount from about 5 parts to about 60 parts per 400 by weight of the solvent/solute system; and preferably includes water present in an amount of from about 15 parts to about 40 parts per 400 parts by weight of the solvent/solute system; (h) said solubilizing agent has a boiling point from about 50 0C to about 4000C and is present in an amount from about 50 to about 100 percent by weight of the organic fraction of the solvent/solute system, preferably about 50 to about 75 percent by weight of the organic fraction of the solvent/solute system; (i) said contacting takes place at a temperature of from about 100 C to about 400 C, more preferably about 140 C to about 360 C, most preferably about 260 C to about 360 C, and at a pressure of from at least about 300 psia to about 2500 psia and for a time period sufficient to produce hydrocarbon liquids from said carbonaceous material; and (j) said pressure is from about 500 psia to about 1500 psia. 14. A method of converting carbonaceous materials to liquid products under conditions of temperature and pressure which do not produce significant thermal bond rupture in the carbonaceous materials, said method characterized by: contacting the carbonaceous material with an excess of a solvent/solute system in the range of from about 1.5 to 1 to 5 to 1 solvent/solute to carbonaceous material, said carbonaceous material is contacted with the solvent/solute system in the presence of carbon monoxide, such solvent/solute system consisting of: (a) an organic phase comprising a solubilizing agent containing more than about 50% by weight of a member selected from aromatic phenol, polycyclic phenol, susbstituted phenol and mixtures and derivatives thereof; and (b) an inorganic phase comprising an aqueous solution of a compound having a cation selected from alkali and alkaline-earth metals; said contacting being conducted at a temperature less than about 4000C and a pressure of at least about 300 psia.

Inventors: BAIN, RICHARD L., KAESZ, HERBERT D., PORTER, CLIFFORD R., RASTLE, CHARLES M.
Applicants: PENTANYL TECHNOLOGIES, INC.
Application year: 1984
Main classification: C10G1
IPC codes: C10G1