CA1226838A - Integrated ionic liquefaction process - Google Patents

Abstract

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 stabilization if necessary.

Description

.1 ~26838 INTF.GI~TEi) IONIC_TIQ~IiF'AC11~N PROCESS
This invention relates to the liquefaction of carbonaceous materials and it particular to a method of producing useful i)r-o(iuc~ts inc1u~ir1g liquid products wZlic11 can be use ciircctly in peLroleum-~ike refining processes or as d foe stow for foreteller cl1emical synthesis, or as a lo sulfur furl oil or tile like, and pseu~30-plastic product Whitehall meltinc3 points in the range of from about Luke to 200C which are solid products at amiably temperature Ann are characterized by Tory released sulfur anc3 Asia content which make them useful as fuel anc3 coke substitutes.
One of tile principal approac11cs to coal liquefaction and salvation in the past has employed reactions producing free radicals trough thermal bond rupture.
or many kinks ox coal used this typically required temperatures above await 350 C for enough free radicals to form through tl-ernlal Boone retooler ox carbon-carbon l~orlcls, car~oll-oxygen Berlioz, carbon-r1itro(3ell buoyancy Ann c~arbor1-sulfur Bills to react Whitehall other eonlpour1(1s or l1ydroycn in order to corn lower molecular weight compounds than tile cowlicks materials present in the coal. In some typical prior processes the free radicals 'corned were stateless and hydrogerlated by hy~rocJerl atoms ti1rouc~}l i1ydrogel1 transfer from solver1t ~30r10r molecules SWISS as lo 3 4-tetral1ydronaphthalene or 9 lO-dil1ydro-phenant}1rene, often used in the presence of mincer amounts of coal solubilizil1c3 ager1ts. flyer Seiko reactions the Swiss are generally s~E~1rdtecl by 30 ~listillatior1 or by solvent de;lsl1ir1g. Ion Slush systems to be effective it is important Tut tile mixtures be suitable for either effective c3istillation or dashing without excessive coxing. In processes Weller oven minor amounts of higl1ly solar solver1ts are em~loye(l either in ~226838 the liquefaction reactor or subsequently, distillation problems are encountered and when alkali compounds are present, severe coking problems can arise.
The present invention provides a method of 5 converting carbonaceous materials to liquid products under conditions of temperature and pressure which do not produce significant thermal bond rupture in the carbonaceous materials which comprises contacting the carbonaceous material with a solvent/solute system 10 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 15 solution of a compound having a cation selected from alkali and alkaline-earth metals; the contacting being conducted at a temperature less than about 400C. and a pressure of at least about 300 Asia.

3 Jo 38 These an-l oilier objects of the invention together White the features and advantages thereof will Bucknell apparerlt from the following detailed specification when read in conjunctioll with tile accoMpanyirlg drawings in which like referrals numerals refer to corresponding parts and:
Fly. 1 is a scllenlclLic flow chart of a typical integrated ionic liquefaction process useful with the present invention for coal Fig. 2 is a schematic reE)resentatioll of a liquefaction subsystem useful with the f low chart of Fig. l;
Fig. 3 is a schematic representation of a separation subsystem useful with the present invetltion;
Fix. 4 is a schelllatic representation of a distillation subsystem useful with the present invention;
Fig. 5 is a schematic representation of a hydrogenation subsystem useful with the present involutely;
jig. 6 is a schenlatic of an acid hydrolysis subsystem useful with the present invention;
Fly. 7 is a schematic of another acid hydrolysis subsystem useful with the proselyte invention;
Fig. 8 is a scilelnatic flow chart of another typical 25 ionic liqllefact:ion process using acid hydrolysis useful with the present ir-lverltiorl;
Fig. 9 is a schelrlatic flow chart of another typical ionic liquefaction process using coking useful with the present invention;
Fig. 10 is a scllerriaLic flow chart of arlotller integrated ionic liqllefactiorl process useful for the prudishly of a pseudo plastic solid fuel and coke substitute;
Fig. 11 is a schematic illustrcltion of a liquid-liquid solvent extraction system useful in the process Shirley in icky. lo and its. 12 to 14 are a series of graphic representations showing carbon monoxide conversions and temperature profile plotter against tinny for a series of 5 experimental runs accordinc3 to toe techniques of the present invention.
As used Harley, the term carbonaceous material, includes solid, semi-solid and liquid organic materials which are susceptible to the described treatment lo metalloids. Examples of solid carbonaceous materials which may be used in connection with the practice of the present invention include coal, such as anthracite, bituminous, sub bituminous and lignite coals, as well as other solid carbonaceous materials, such as wood, 15 lignin, peat, solid petroleum residuals, solid carbonaceous materials derived from coal, and the like, depending on the proc3ucts sought. examples of semi-solid and liquid carbonaceous materials include coal tars, tar Sweeney, asphalt, shale oil, heavy petroleum 20 oils, light petroleum oils, petroleum residuals, coal derived liquids and toe like.
Ionic liquefaction as used herein is intended to mean the chemical process described herein, weakly is characterized by polar solvent solubilization of the 25 polyllleric structure of carbonaceous materials susceptible to the described treatment methods, in the presence of alkali anal alkaline earth compounds in amounts which favor ionic reactiolls invoLvinc3 the solubilizecl carbonaceous material all ionic species such 30 as phenoxic3e, hydroxide, and format 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 feed stock.

In a~ditiotl to ionic reactions it is believed that ionic liquefaction as described herein may clingy the apparent molecular weight and other E-l~ysical properties of toe solubilized carbonaceous material by reducing the 5 extent of Huron bonding between carbonaceous material molecules. Because of toe reactive nature of ionic species remotely after solubilization an ionic reaction the product mixture must be further processed in order to stabilize tile desired products to be able to 10 recover liquid and solid usable products including a recycle stream rich in phenolics weakly can be used in the ionic liquefaction reactor. This processing involves removal of alkaline salts and further stabilization of organic ionic species by hydrogenation 15 acid hydrolysis solvent extraction or coking.
The tern alkaline as used yencrally and herein is synonymous with basic which inkwells without limitation alkali metal anal alkaline earth compounds.
A base can ye an aqueous SC)l.-ltiOrl Waco COrltains 011 ions or any ubstarlce wise accepts protons or any substance Wylie is an electron pair donor Typical cations are the light metals of Groups It and IT of the Periodic Table. Preferred cations are pa and K .
Typical anions include ill C03-, ICKY and ~C03 .
As applied to ionic lic~uefactiorl a polar solvent or polar solvents tneans a solubilizing agent selected from the group consisting of aronlatic alcohols, ~}lenols, polycyclic enlace atld substituted pilenols and nlixtures thereof. Typically slush solvents do riot have aft (x -hy(lroget~. Liquid mixtures of solubilizitly polar solvents used in the ionic liquefaction process of the present invention typically will contain treater than 50% by weight of Sicily polar solverlts.

3~3 More particularly, in accordance with a preferred embodiment of the invention, useful hydrocarbons are obtained prom carbonaceous neutrals by contacting the carbonaceous materials with a solvent/solute system 5 consisting of (a) an organic phase comprisirlg a solubilizing agility containing more than about fifty percent (50~) my weight of a phellolic-type solvent such as an aromatic pherlol, polycyclic phenol, substituted phenol or a mixture thereof, and (b) an inorganic phase 10 comprising an aqlleous solution of one or more alkali or alkaline-earth metals. More particularly, the solvent/solute system comprises ogle or more solubilizing agents selected prom aromatic phenolics, e.g. phenols, arid polycyclic and/or substituted phellols, typically of lo Iron 6 to 15 carbon atoms, e.g. o-cresol, m-cresol, p-cresol, catcall, resorcinol, naphthol, and mixtures and derivatives thereof. although not essential to the practice of this preferred embodimerlt of the ir-lverltion other orcJanic constituents such as aromatic alcohols, 20 polycyclic aromatic hydrocarbons, partially hydrogenated and/or fully hydrogenated polycyclic aromatic hydrocarbons, typically having from 1 to carbon rings, and more preferably from 2 to 3 carbon rings e.g.
naphthalerle, antllracene, phenanthrerle, Tetralin (tetrahydronaphthalene), gamma-picoline, isoquinoline, dihydronaphttlalelle, Decline (decailydrorlaphthalene), 9,10-dihycdroanthracene, 9,10-dihydrophenanthrene, and mixtures end derivatives thereof also may be included in the solvent/sol-lte Sesame Synthesis gas, as that term is usual herein, means a gas primarily comprised by carton monoxide and hydrogerl. Other gaseous compollel-lts present in small concentrations can include Carolyn dioxide, light hydrocarbon gases, and some impurities such as llitrogen 83~3 and still be effective in the process described herein.
In adulation, small amulets of water vapor may also be present.
It has been shown that solubiliæa~iorl of coal and other carbonaceous materials can be achieved using a variety of coal delve solverlts and outlawry organic solvents. For example, US. Patent Jo. 4,133,~-16 leaches tile adva~ltayes of using minor amourlts of finlike recycle solvents in coal liquefaction. Similar advantages are taught by Comma et at., effect of Finlike Compounds on liquefaction of Coal in the Presence of Hydrogell-Donor Solvent", fuel, Vol. 57 November 1978), pp. 681-6~5; an by Sums, et at., "Internal Rearrangement of llydroyen During floating of 15 Coals with Phenol", Fuel, Vol. 60 April 1981), pp.
335-341. It is also Noel that the use of various bases, eye. Naomi, Nikko and Nal3CO are useful in carbonaceous liquefactiorl processes. See for example, Donovall et at., "Oil Yields from Cellulose 20 Liquefaction", Fuel, Vol. 60 October 1981), EN 899-902 and Roy et at., "Study of Treatments of Sub bituminous Coals by Naomi Solutions", Eel Vol. I (December 1981), PEW- 1127-113().
Ross and Blessing have described low coal may be solubilized and h~droyen added to the organic product by aqueous base "T-lydroconversion of a Bitulninous Coal with Cole", Fuel, Vol. 57 (June 197~), p. 379. They slave I I

also described flow alcohols h3virlg an ':~ -hydroc3en atom are effective hydroyell dollar solvents wren catalyzed by alkaline compounds (US. Intent NO. 4,29B,450). They state, }however that alcohols not having an -hydrogen are not effective solvents.
The unexpectedly high solubilization and liquefaction achieved in the present inventiotl through the synergistic effect ill a liquefactiorl reactor of a solvent/solute system combirling a pllenolic solvent 10 water and added amulets of an alkali or alkaline-earth metal compound, with or without the additiol-l of synthesis gas depellcling on tile reactants selected has made possible the investigation and discovery of other promising opportunities for enhancing the liquefaction 15 and the hydrogen to carbon ratio of the resultant products which will be more fully described hereirlafter.
Appeal et at. in their paper entitled On the Mechallism of Lignite Liquefactiorl with Carbon Monoxide and water m. and In. Vow 47 (1967) p. 1703 20 describe IIOW usirlg frostily Powdered low-rank coal and a selected solvent will produce a 72% yield of a benzelle-solllble oil when usirlg operating pressures near 5~00 Sue and telllperatures in excess of 365C. pull et at., also describes the use of a solvent comprisinc3 25 alplla-naphthol (a phenol) phenanthrelle ( d polycyclic aromatic hydrocarbon) and water in the presence of naturally occurrir-lg amolJnts of alkali or alkaline-earth metal compounds at similar ol)er~tillg collcli~ic)Zls. In addition testinc3 is described involvirlg tile additiorl of K2C03 in water as a solvent with the conclusion that the addition of K2C03 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 corlditions employed.

31~

Farcasiu et at., US. Patent No. 4,1~3,646 (1979) leach that an improved liquefaction process can be obtairled using a Taoist liquefaction process wherein a donor solvent in co~nbinatioll with minor amounts of 5 finlike compounds is reacted with coal and Hydrogen at 600 - ~50F. the unrequited coal was recovered by filtration, finlike compounds were recovered by distillation, or extraction, anal the resulting substantially finlike free distillation residue was 10 subjected to various upgrading treatments such as delayed coking and hydrotreating.
Farcasiu et at, did not obtain the synergistic effect realized by the present invention using solvents containing greater than 50~ by weight of finlike 15 compounds with added alkaline compounds in a single stage liquefaction reaction. Farcasiu et at, similarly also did not recognize that pherlolic compound recovery from the product stream must be preceded by alkaline compound removal. Likewise, Farcasiu et at does not 20 show how alkaline COnlpOlilld separation may be controlled by varying the water level during the ad-led alkaline compound removal step. Further, there is no disclosure in Farcasiu et at of how the process could be advantageously improved by addirlg synthesis gas to the 25 first stage reaction, either as a reaCtallt in the format ion- chemistry described hereinafter, or to produce a hydrogen enriched stream for upyradirl~
operations. reside the foregoing, Farcasiu et at does not disclose that residues contaillitlg lore than about 50% by Walt phenolics may be uE)~raded by hydrotreatirlg or coking, but instead indicates that toe residue should be substantially phenol free.
The foregoing and other art, in sunlmary, has not recognized that a liquefaction process for coal or other carbonaceous material can be substantially improved by I

the use of an organic solvent containing greater than 50% by Walt of pllel-lolic c~mpourlds in amounts between about lo to 5.0 times the weight of carbonaceous materials, in combillation with between about 25 to about 5 400 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 10 solvent/solute pair are reacted together at temperatures less Thor about 3G0 C and pressures between about 300 Asia (2.0~ Ma) to about 2500 Asia (17.2 Ida Further improvement can be obtained by tile presence of synthesis gas havirlg a TAO ratio between about 0.5 and 2.0 in 15 amounts between about 0.16 to about 1.25 m3/Kg of carbonaceous material when combined with stabilization of the reaction product before further upgrading, after removal of the alkali and water.
Gore particularly, in accordarlce with a preferred 20 embodimerlt of tile inventioll the solvent/solute systems useful in the practice of the invention are solubilizing mediums comprising organic and inorgarlic fractions or constituents which may syllable a portion of the carbonaceous material and/or may otherwise enhance liquefaction of the carbonaceous material. As noted swooper, tile organic fractions of the solvent/solute systems comprise one or more solubilizing agents selected from tile group consistillg of 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, catcall, resorcinol, naphtllol, and mixtures and derivatives thereof. Although not essential to the practice of this embodiment of tile inventiorl the solvent/solute systems in many instances will include other organic constituents. Suitable organic constituents inkwell aromatic alcohols polycyclic aromatic hydrocarbons partially hydrogenated and/or fully hydrogellated polycyclic aromatic hydrocarbons typically having from 1 to 4 carbon rinks and more preferably from 2 to 3 Carolyn rinks e.g. napllthalene anthracene, phenanthrene, acenallthene, l-methylnaphthalene 2-methylnapht}lalelle Tetralin (tetrahydronap~thalene), gamma-picoline isoquinoline, 10 dillydronapl~thalene, Decline t~ecalhydronapl~tl~alene), 9,10-clihyc3roantllracene 9 10-dihydrophenanthrene and mixtures and derivatives thereof.
During initial phases of operation, some of the above mentioned solubilizing agents anywhere other organic 1 constituents will be present: then in subsequent operatiorl the organic constituents will be carbonaceous material-derived pherlols of the type and polycyclic aromatic hy(3rocar~)ons of the type or derivatives related to toe type c3escribed herein before.
20 Particularly useful organic please solubilizing agents and/or other organic fraction constituents have a boiling point above 50C more preferably of from about 100C to about 460C arc most preferably of from about 150C to about 400~C`. In the practice of 25 this latter embodimerlt the solubilizirlg agent is typically from about 50 to 100 weight percent of the organic fraction of the solvent/solute system.
Suitable inorganic fraction constitllents of tile solvent/solute system ionic water, all alkali and/or 30 alkaline-eartll metal complies arid their derivatives.
The water corltent Cain be from about 5 parts to about 60 parts per 400 parts by weight of the solvent/solute system more usually about 15 parts to about I parts per 400 parts by weight ox the solvent/solute system.

1~2~;838 Suitable examples of alkali an alkaline-earth metal compounds include hydroxides carbonates bicarbonates, nitrates, sulfates, sulfites, sulfides formats and other salts nnixtures thereof and the like, although other compo-lnc3s may be employed for the pyres Specific examples include Noah Nikko Na~lC03 Nazi' Noah, KIWI K2C03 COOK Cook mixtures thereof and the like. Presently preferred species are Naomi KIWI and awoke in from about 1 10 part to about 40 parts per 400 parts by weight of the solvent/solute system more usually 1 to about 15 parts per ~00 parts by eight of the solvent/solute system.
It is understood that the amount of alkali or alkaline-earth metal complied present for purposes of 15 the present inventiorl is an added amount i.e. an amount in excess of the amount which would be present from the various naturally occurring alkali or alkaline-earth metal companies. Flowerier it is understood that the alkali or alkaline-c,rth metal compound content will be 20 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 tile solvent/solute system required in the reaction mixture us slurry is dependellt upon tile amount and nature of tile carborlaceous material to be treated. Generally it is preferred to employ up to about 500 parts of tic solvellt/solute systelll to 100 30 parts of carbonaceous material more preferably at least about 350 parts of solvent/solllte to 100 parts of carbonaceous material and nicety preferably at least 150 parts of solvent/solute systelll to 100 parts of carbonaceous material.

cording to tilts embodiment, carbonaceous material is solubilized in tile solvent system - alkali mec3ium to form a reaction mixture or slurry. Frequently, the reaction collditiorls are water gas shift reaction 5 conditions, as hereillbefore described eye reaction mixture is heated to a sufficient temperature, typically below about 400 C., and pressure to obtain enhanced syllabling of tile carbonaceous material for production Andre conversion 10 of hydrocarbon liquids, as herein before defined, from the carbonaceous material. Err most purposes, it is contemplated that sufficient temperature levels for the solvent/solute system are from about 100C to a temperature below about 400C under the reaction 15 conditions employed, more Preferably from about 140C
to about 380C, and most preferably from about 260 C
to about 360C, at a pressure of at least about 300 Asia (2.06 Ma), more preferably from about S00 Asia (3.4~ Moe) to about 25()0 Asia (17.2 Ma), and most 20 preferably from about 5~0 Asia (3.45 Mesa) to about 1500 Asia (10.35 Ma). It has been phonic that under the foregoing reaction condition, relatively short periods of time result ill the production of tile desired product. Alto sufficient times are dependent upon 25 the nature of tore carbonaceous material, the reaction conditions employed, and tile like, for the most purposes, it is contemplated that reaction times of at least about 1 minute, more preferably fr(-lllclbout 10 minutes to about 120 moonlights, and most preferably from I about 15 minutes to about 60 minutes are sufficierlt to result in enhanced syllabling and tile reduction and/or conversion of hydrocarbon liquids.
As will be appreciate by those skilled in the art, the solvent/solute systems containing coal or other carbonaceous material solubilized according to the present inventior-l, may be foreteller treated as described herein .

~26838 - I -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 liquefactiorl, the predominant chemistry can be termed as solvation-ionic chemistry, involving solubilization of tile coal polymer by polar finlike solvents, followed in situ by attack upon the coal structure by ionic species such as 10 phenoxide and format 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 grouts present in the coal. The esters, kittens, and ethers present, are all sites for 15 n~cleophilic substitution. Hydroxide anal carboxylic acid containing functional groups are considered to be substantially unreactive to nucleophilic substitution.
The~phenoxide ion mechanism for an ether group is:

.

i83~

Initiation ___ OWE
SHEA ' + Elm Reaction II.

-c~3 R-O-R OR' - O

Hydroxide Regeneration 15 III.
R' - I H20 .- > R' - OH -t OH-Solvent Regeneration I. _ IV.

0~3 clue + Ho + R - En l~Z~838 The parallel format ion mechanism V to VII is V. CO + I 2 5 TV 02 R Roll + C02 VII. Roll + 1~20 if + Oil-This mechanism is believed to be enhanced at the mild 10 conditions employ because of the intimate contact made possible by tile finlike solvent sQlubilization where there is competition for available ionic species in the presence of the water gas shift reaction VIII to X.

15 VIII. KIWI + Lowe 3 ll2 IX. COY + CO + Lowe -- ICKY + KIWI-X. ' I I-_ _ COY + C02 + ~12 The summation of VIII, IX and X is the net water gas shift reaction XI:

XI. CO + owe C2 + 1l2 Ionic liquefaction, therefore, offers many process advantages over conventiollal thermal rllpture free radical liquefaction. 'I'll teln~)eratllre Rome for ionic liquefaction is typically ull(ler 360 C. it these temperatures tllerlllal pond rupture is riot the primary mechanism because it proceeds slowly. The lower temperature employed has the effect therefore of reducing the amount of methane, ethylene, ethereal, and acetylene produce from 20% to less titan 1% of the MA
carbonaceous material.

~'~26838 System pressure for ionic liquefaction can be obtainer from toe 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 Tao solvents are in the liquid phase. this is typically 3.45 - 10.35 Ma (S00 - 1500 Asia). Tile ionic li~uèfdction mechanism described Erwin to produce a liquid can in addition use synthesis gas which is composed of carton monoxide and 10 hydrogen instead of requiring expensive pure hydrogen.
This synthesis gas is the fuel for the production of pure hydrogen therefore Tao predilection of pure hydrogen is not needed to obtain the desired improvement in hydrogen to carbon ratios possible in the liquid 1 products of ionic liquefaction In free radical liquefaction a typical mechanism is:
Thermal Rupture ZOO coal OK-hydrogen Donation ____ .__ XIII. C10 ~112 t I ROY + C10 lo Tetralin Dillydronaphtllalene (~etrahydl-onap}l~]la~ene) Solverlt Rcgel)eraLion _ _ . _ _ _ XIV. C10 Lowe I - - - I 11 In order to promote he Solverlt Re-~eneratioll reaction (equatioll XlV) free ridicule liquefaction is typically performed Usual lli-jll purity 11ydlo(3ell at pressures bottle 13.~3 - 17.2 aye (2000 - 2500 Asia).
The ionic liquifactioll described Lorraine at the conditions employed will prowls polar products having low molecular wits. isle free-radical mecilanism will produce mainly nonpolar materials. The free-ra(3ical 122683~3 type mechanism can, in additioll, lead to polymerization reactions Welch produce undesirable high molecular weight materials. essay type of reactions are not favored and therefore avoided in the ionic liquefaction described herein.
Tile Presence of highly polar pherlolic compounds, alkali compounds and the ionic forms thereof, along with the oxygenated cor~lpounds derived from the carbonaceous material, durit-lg the ionic liquefaction 10 process described herein, can, however, lead to potential processing problems downstream from the liquefaction reactor. Sole oxygenated compoullds, including many phenols, are often thermally unstable, especially in the presence of coking promoters such as 15 ionic alkaline species. These compounds are concentrated in the bottoms durirlg distillation operations. It is, therefore, important to minimize the amount of alkaline compounds present in the distillate feed stream. A nlajor proportion of the alkaline 20 companies oily, therefore, be removed before distillatiori;is Boone.
In addition, the proposed phenoxide ion elicitor shows a solvent incorporation step as an inhererlt part of the reaction mechanism. Ullder the con~:litiolls 25 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 recJenerate solvent and obtain additional product, a p(lrtiotl of the lic3llefactior 30 product liquids shallowly E~Leferably undergo further reoccur in the wrists of i)ydroc3etl under conditions which will break the ether bond between the phenoxide ion and the coal derived orgatlic species. The reaction nay be performed at conditions Weakly are severe enough 35 to break the ether bond, but are not severe enough to 122~83~

saturate the aromatic rework, or remove the oxygen atom as water. The hydrogenation acid treatment or solvent extraction steps will also serve to stabilize toe product by re~ucirlg the concelltration of the most urlstable ionic 5 species an reduce tile ash contralto of tile product.
Because the ionic liquefaction products produced by the liquefaction processes described are typically high in reactive oxygenatec3 species there can be a tendency for the products to oxidize and/or polymerize with time 10 or with thermal treatment. In title case where additional solvent recovery arid product upgrading are necessary specific sequential precisely steps are then needed. A
process incorporatillg these necessary steps is shown in Figure 1.
Referring to the schematic diagram in Fig. 1 the feed proration at (A) comments tile carbonaceous material, stream (1) Lye conventional means Sicily as ham~ermills or ball mills or comparable equipment and adds a water-alkali mixture stream (4); and recycle 20 polar solvent streams (2) and (3) containing greater than 50~ by;weigtlt of finlike species. Tao comminution process may be accomplished either dry or wet. If performed wet theft the recycle polar solvent may be used as the welting agility if proper preclutiolls are 25 taken. The carbollaceous feed is preferably commented to lo percent minus 74 micron (200 mesh) particle size more preferably to 100 percent minus 147 microns (100 mesh) particle size aureole most preferably to lo percent minus 350 microns (40 mesh) particle size but irk any 30 event must be in a form which will enable tile requisite solubilization for tile ilk liquefaction to proceed Using 1000 parts by weight of stream (1) carbonaceous material as an e~anlple the preferred amount of polar recycle solvellt for the recolored solubilization to prosaic recycle streams (2) plus (3) _ 20 _ is between 1500 anal 3500 parts by weight depending on the prepared form of the carbonaceous material, with 2000 parts by weight of solvent toe most preferred amount. The polar recycle solvent contains preferably 5yreater than about 50% by weigh of pherlolic compounds, and more preferably greater than 60~ by weight finlike compounds.
The preferred amount of alkaline material in water-alkali mixture stream (4) is selected to be enough 10 to produce the desired results. It has been found under the conditions disclosed herein that between about 25 parts and 400 parts by Walt per 1000 parts by weight of carbonaceous material is effective with the more preferred amount being between about 25 and 150 parts by 15 weight, and depending on the kinds and amounts of finlike compoullds employed, the preparation of the carbonaceous material and the conditions selected, the most preferred amount is about 50 parts. The anoint of water in stream (4) should preferably be sufficient to 20 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 regeneratiotl and ionic species stabilization 25 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 30 be provided to stream (4) via a make-up alkali stream (PA).
he 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 (S) from the 35 preparation section (A) of Fig. 1 and to bring the Sue slurry to tile dozier system pressure. The slurry then passes ~l~rou(JII to sl~lrry-llot precut float exchanger (L) to provide initial slurry littoral Lllrougll waste heat recovery from tile product slurry stream in the Separation Section (~) of Fly. 1. Toll slurry is theft mixes with hicJ11 rouser synthesis gas stream (6), from the gasification as processillg section, secLiol-s (G) arid I of Fig. 1.
If syllthesis yes is used, the composition of the 10 synthesis gas slur n Swahili primarily be composed of carbon monoxide and llydlogerl, altllou~ll small amounts of impurities Sicily us neutron Ann carbon dioxide can be preserlt and still achieve tile most referred results.
Lowe preferred 1l2/Co ratio unwire the conscience 15 described is beLweerl ablate 0.5 and 2.0; a more preferred ratio is between about 0.5 Ann 1.5; an-3 the most referred ratio is ennui about 0.75 arid 1.25. Other concentrations untrue different reaction conditions Jay Al sole effective.
Proofer (was treatnlellt rates are between about 1.25 m keg carbonaceous m~Lerial and 0.16 McKee I 000 SCF/ton to 5 000 So oil) tile most proofer gas treatment rates are between about 0.47 and ~.16 m3/kg (15,000 to 5,0()0 clown).
Referring to jig. 2 tile reactioll mixture, stream (26), is then royalty to the desire t~ml~erature in a gas fire tube huller I operated under turL~ulerlt flow conditions to optima float Lrarlsfer. irises derive low or medium ca~orific value was from JaS processing sections (If) and (1), lines elf aureole IF, is used to fire tile prelature Lowe reduction Myra stream (27), is then transferred to the liquefactioll reelection (N) and held at the desired reaction tenlpc?rature long enough for the initial ionic liquefdctioll process steps to take 1226~338 place to tile desired exterlt. The solubilization and ionic reaction stalks but not toe solvent regeneration or ionic species upgrading, take place in this reactor.
The preferred reactor resign is a high length to diameter ratio reactor Whitehall internal waffles to provide sufficient mixing of the polar solvent, alkali, and carbonaceous materials for the reaction mechanisms to proceed efficiently DeE~endir1g Spoil the desired product distribution, Audi desired reactant residence time lo distribution, outlawry reactor designs can, 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 mortise, and most preferably from about , 15 to 45 minutes depen(Tinc3 on reactor design. The preferred temllerature range is from about 250C to 350 C. (482 F to Go F`). It is important that a syStelll pressure is selected and maintained at such a precleterrnil1ed level which will keep sufficient water in the Lockwood phase to minutely ionic alkaline species in the liquid phase, rather than as salts, during a substantial pc)rtio1l of the reaction time. The preferred pressure range is between await 3.45~ a to l7.24ME~a (500 Asia to 2500 Asia), and toe most preferred pressure range is about aye to 8.28MPa (800 Asia to 1200 Asia) for achievil-1cJ tile foregoing lookout phase.
The reaction products exiting reactor (N), stream (7) of Yips. 1 and 2, are theft separate into component streams in the separatioll sisterly, section (C) of Fig.
l. A more detailed schematic of tilts system is shown my 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 toe slurry product from gaseous products. Tulle separator pressure is maintained at a lesser or reduced pressure than the lZ26~338 liquefaction reactor, preferably between about 0.69 and

3.45 Ma (100 - 500 Asia), anc3 the temperature is kept at conditions sufficient to keep most of the volatile organic compounds in the liquid phase, but much of the 5 remainir-lg water as vapor. The preferred temperature range is 200 to 300 C (392 to 572F) the most preferred temperature range is 200 to 250C (392 to 482F). The slurry phase residence title 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 canonized, along with residual organic compounds, such as polar solvents and (C4+) hydrocarbons. Typical concentrations of water vapor in the gas Pilate enterirlc3 the corldellser will be between about 25 to 45 mole percent and leaving the water condenser the water vapor in the gas phase will have been reduce to between about 0.01 and 0.05 mole E~ercellt. the oiler com~)onerlts of Tao gas stream are typically,~arbon monoxic3e, carbon dioxide, Of through C3 hydrocarbons, COST ll2S, Nl13 and possibly snowily amounts of other gases, such as nitrogen. The preferred H2/C0 ratio of tilts gas stream should ~erlerally be greater than about 2 to 1, and most referable greater than about 9 to 1. The water rich stream, stream (31), from the condenser (P) is preferably withdrawn through, stream (33), from the se~)aratiorl system, and may be used for plant cooling water or tile like. Louvre, water may be addec3 back to tile slurry from the gas-slurry separator (stream 32), if Ned, to improve tile slurry filtration characteristics. the slurry stream (24), goes to the feed slurry-llot product heat exchanger (K), then through a pressure letdown valve and through stream (30), to the filtration apparatus (Q).

Leaf filters, candle filters, centrifuges, hydroclones, or comparable equipment can be used for the filtration. Solids may also be separated by processes such as solvent (essaying, or critical solvent 5 extraction. toe purpose of tile filtration step is to separate unrequited carbonaceous material and alkaline salts from the ionic liquefaction precut likelihoods. Due to the hydroscopic nature of the alkaline salts, much of the water preserlt will also be separated with the filter lo cake.
The preferred temperature for filtration is preferably bitterly about 150 and 100C (302 and 212 F). the pressure is maintained at a sufficient level to obtain efficient filtration, preferably between about 0.34 and 1.03 spa (50 Asia to 150 Asia). In the filtration step, the solids content of the liquid is typically reduced to less thrill about lo percent by weight, the mineral matter content to less than about pursuant by White, Ann the alkaline species contralto to less thrill about 0.25 percent by weight. The values obtained are deperl~ent upon factors such as the degree of cornminutiorl used in feed preparation, the ionic liquefaction conditions selected, and the design of the filter equipment design. The SeparatiOrl train should preferably be operate to reduce the alkaline compound content of the filtrate stream (10) of Ego. 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. ~lkalirle cornpourld concentratiorls above tilts level are not desirable swirls that can Lowe Lo Cockney problems, and to unwanted organic ionic corrlpoulld 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 tlydrogenation (E), although the ~L226838 upper limit is deponent upon the exact nature of tile product. Roy solids riot) filter cake, stream 8 is used as feed to a yc3sification system (G) for production of synthesis gas and recovery of alkaline compounds (J).
The gas stream, stream ('JOB) of Fig. 1, is sent to gas processing (I) for upgrading before use as hydrogen rich gas in subsequent llydrogenatiorl (E) operations.
The liquids from the separation system are sent, stream (10) of Yip. 1, to a distillation tower (D) where the crude ionic liquefaction product is fractionally distilled. The products from the distillation are water, stream (If) of Fig. 1, a recycle solvent stream, stream (2) of Fig. 1, and a crude product, 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 finlike materials is recovered. hire a concentrated ionic compound stream is produced.
The production of the concentrated crude stream serves to reduce the size of downstream stabilization operation. This operatiotl is typically accomplished by 25 taking a 200C (3'32F) cut of the incoming feed.
The distillation column is typically operated at or near atmospheric pressure preferably at about 0.10-0.15 Ma (14.5-21.75 Asia). Tile condenser is operated as a partial condellser to separate the overhead product into an organic rich liquid flus and a water Rockwell 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, althouytl the exact value will vary depending on the bottoms viscosity desired.
he 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 38C (10 and 100 centistokes at 100F). Tile preferred density of this stream is between about 1000 and 1200 kg/m (1.0 and 1.2 grams/ml).
Tile overheads organic liquid preferred viscosity is between about 5 x I 6 and 2 x 10 5 m2/s at 38 C
(5 to I centistokes at 100F). The preferred density of this stream may be between about 950 and 1100 kg/m (0.95 and 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 38C (5~0 and 2000 centistokes at 100 F). Tile preferred density of the crude is between 1100 and 1200 kg/m3 (1.1 and 1.2 grarns/ml). The crude will normally contain at least about 40% finlike complies Three processes may be used in the ionic liquefaction process describe to upgrade the resultant crude. Tile first embodiment shown in Fig. 1, is hydrogenation. Iteratively tile crude may be stabilized by assay hydrolysis or by coking. A schematic of a typical hydrogenation anti distillation train is shown in Fig. 5 a schelllatic of a typical dCld 30 hydrolysis train is S}lf)Wrl in jig. 6 Ann a schematic of a typical process incorporating coking is shown in Fig.
9.
Referring to the process schematic shown in Fix. 5 the hydrogenation train may consist of four Niger processing units: (1) crude preheater (R) and (S ) byway (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 fire tube heater (S'). The preheated crude stream is then passed to the 10 hydrogenation reactor (T) where it is reacted with a hydrogen-containing gas which is introduced into the hydrogerlation reactor (T) as stream 13. The preferred operating condition for toe 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, end 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 rinks or the removal of organic oxygen as water. Preferred operating temperature is between about 343 and 454C (650 and B50~, and most preferably between about 343 and 400C (G50 and 752~). The preferred pressure range is bitterly about 6.9 and 13.8 Ma (1000 and 2000 Asia). Tile preferred catalyst types are standard hydrogenation catalysts SEIKO as Comma and Nemo, typically about 1/32, lug or I irlch extradite form. Preferred catalyst loadings are about 0.01 to 1.0 kg cat/kg oilier, more prowar Lennox are about 0.01 to 0.5 kg cat/kg oilier, the lost preferred loadings are about 0.05 to 0.15 kg cat/kg oiler Preferred hydrogen treatment rates are about: 178 to 1424 Mom oil 6i838 (1000 to 8000 Squabble oil); more preferred hydrogen treatmerlt rates are about 178 to 712 m3H2/m3 oil (lQ00 to 4000 SCF`l~2/BBl oil. Preferred hydrogen consumption is less than about 3.90kg ll2/m3 oil (1.36 lb 1l2/BBl oil), and more preferably less than about 1.94 kg 112/nl3 oil (O.G8 lb blue oil). It is desirable to hold the crude at the temperature selected for sufficient time to permit the solvent regeneration and ionic stabilization 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 Witty other exit gases, stream (39), from the hydrogenation train, and passed via stream (1~3) 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. 'I've leach ate is recycled Jo 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 (Fig. 5).
Refined ionic likelihoods in stream (17) are separated from the treatment gases in a gas liquid separator (U).
The gases in stream (39), are sell 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 I wit percent oils, more preferably than go wt. percent oils, and most lZZ6~38 _ 29 -preferably greater thrill 95 wt. percent oils, as defined by pontoon volubility and Assyria D1160-77 distillation results. Tile preferred viscosity of the refined oil is between about 5 x 10 6 and about l x 10 4 m2/s at 38C (5 to 100 centistokes at 100F), and more preferably bottle about 5 x 10 6 and 2 x 10 5 m2/s (5 to 20 scientists at 100~). The preferred density of the refined oil is between about 1000 and 1100 kg/m3 (1.0 to 1.1 grams/nll). 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 kettle hereinafter.
The second upgrading embodiment is acid hydrolysis, was shown in Fig. 6. In this process alternative, hydrogen is added to the tonic organic species in the crude through tonic reactions with acid. In one method, Caribbean acid produced in the acid hydrolysis train is used as the hydrogen source, although smell amounts of 20 other acids alone or in combinatiotls can be use, such as sulfuric Reid, hydrochloric acid, formic acid, acetic acid, and earbamie acid, which will enhance the rate and extent of ionic hydrogenation.
Referring to Fig. 6, the acid hydrolysis train 25 consists of five units, (1) a gas absorber tower (PA) to produce carbonic acid, (2) a crude stabilizer (BY), (3) a liquid phase separator (CC), (4) a high-oxygen product distillation tower (DUD), and (5) low-oxygen product distillation tower (EYE).
In the 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 coxed or carbon dioxide rich gas from the ionic liquefaction reactor or 35 other source is fed in the bottom of the gas absorber lZ26~38 tower (A) 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 5 between 1650 mole/m3 and 1000 mole/m3, which is passed to crude product stabilizer (BY). The absorber is preferably operated in the temperature range 16 to 66C (60 to off and more preferably in the range 16 to 32C (60 to 90F). The preferred operating 10 pressure is in the range of about 0.34 to 10.35 MPa(50 to 1500 Asia), and more preferably in the range 6.90 to 10.35 Ma (1000 to 1500 Asia). The tower may be operated as a bubble cap column, although other designs such as packed columns, venturi scrubbers, or spray 15 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 (BY). Stabilizer (BY) is 20 operated in a stirred tank mode to maximize the contact between the aqueous and organic phases. Although dependent upon tile degree of stabilization required and the strength of the acid, the preferred feed ratio by volume to the stabilizer is lo 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 200 C
(392 F), more preferably less than about 150 C
(302 F), and most preferably less than about 100 C
(212 F). The pressure is kept sufficiently high to keep the majority of the carbon dioxide in the liquid phase, preferably between 6.90 and 13.80 Ma (1000 and 2000 Asia), and more preferably between 10.35 and 13.80 Ma (1500 and 2000 Asia). The acid and crude are maintained at stabilization conditions for sufficient _ . _ 122~;838 _ 31_ 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 an water are added directly to the stabilizer, eliminating the yes adsorption unit. Stabilization conditions employed can remain the same.
A two phase aqueous-organic product, stream (104), is withdrawn from crude product stabilizer (BY) 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 15 temperature. The preferred residence time is between about 10 and 60 mirlutes, and more preferably between about 10 and 30 minutes.
The light aqueous-rich stream is passed to a distillation unit (DUD) where water, stream (107), an 20 oxygen Rockwell water soluble product stream (10~3), and a residual product stream (109) are recovered. Because of the normally high volubility of alkaline campaniles in the water phase, the majority of residual alkaline materials will be present in the residual product from 25 this distillation Unlit.
Isle heavy organic rich phase (stream (105), is passed to a secorld distillation unit (HE), where various distillate products such as naplltha, stream (110), light and heavy gas oil fractions, streams (111) and (112), 30 and a residual product Starr (113) are obtained.
An alternate acid hydrolysis configuratiorl is shown in jig. 7, where an acid other than carbonic acid is used as the hydrogen dolor species. The major processing units in this embodiment are a crude product I

stabilizer (OF), a liquid separator (GO) and two distillation towers, (flit) and (II).
In the crude product stabilizer the crude product stream (12) is stabilized with an acid from stream (115) 5 at a temperature sufficient to permit the ionic stabilization reactiorl 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 reflex 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 preferably 20 between about 204 arid 100C (400 and 212F), more preferably between about 149 and 100C (300 and 212 F), and most preferably between about 120 and 100C (24~ and 212F). The pressure is preferably maintained at about atmospheric pressure. The residence time in the crawled product stabilizer is preferably between about 5 and 90 minutes, and more preferably between 15 and 45 minutes.
The two-phase aqueous-oryanic product in the stream (114) is then sent to the liquid separator (GO) where 30 gravity separation is usual to generate a light aqueous-rich prodllct 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.

.

12Z~3~3 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 (OF). As noted swooper, make-up acid can come from the acid-gas removal unit in the gas processing unit, (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 OFF). An organic distillate, stream (126) and a spent acid stream, stream (127) consisting of heavy organic and alkaline salts such as Nazi are also produced-The organic rich phase in stream (117) is passed to lo a second distillation unit lo 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 OF for the tonic 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 (OK), with coke, stream (Al) being produced, end water, stream (K2) being removed.
Coking is defined as a severe therlTIal cracking process in which one of tile erlcd products is a carbon rich solid, 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 427C (800F) to 510C
(950F).

- I -Using 1000 parts of crude as an example, 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 51iquids. The liquids produce 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, 15 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 aphid stream I to a gasifies, unit (G). In the 20 configuration Shirley in jig. 1 the preferred gasification operation it a dry ash partial combustion process to produce a synthesis gas rich in carbon monoxide and hydrogen. The preferred ~2/C0 ratio of gas leaving the gasifies, stream I is between 0.5 and 2Ø Acid 25 gases and light: hydrocarbon gases are separated from the synthesis gas by standard operations in gas processing operations, unit I
The ash from the guesser, stream (21) is serlt to an alkali recovery unit (J) where the majority of the 30 alkaline compounds are separated from tile ash stream by hot water extraction, all recycled, stream (4), to the preparation unit (A), and the residual ash removed, via stream (Jo). The expected recovery of alkaline compounds is greater than 50%, and can be greater than 35 75%, and even greater than 90%. The preferred ~22f~838 temperature is between 25 and 100 C (77 and 212 F), and more preferably between So and 100C (122 and luff). Lowe preferred treatment rate is less than 4 kg ~20/kg ash, more preferably less than 2 kg H20/kg 5 ash, and most preferably less than 1.1 kg El20/kg ash.
In an alternate configuration alkaline compounds may be removed from the feed, stream (8), to the gasifies (G). The preferred processing conditions remain as before.
For hydrogenation (jig. 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 C02 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 C02, stream (lo), for use in the acid hydrolysis operation.
In tlle,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.

sty EX~IPIE 1 _ .
A series of experiments were performed to model the integrated ionic liquefactioIl 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 cotnponents of the system are the gas delivery system, reactor system and gas measuremeIlt system. The gas delivery system consists of lo White gas compressor. The primary components of the reactor system is a 0.001 my (1.0 liter) magnedrive Autoclave manufactured 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 Carte Series S
chromatography for on-line analysis of water gas shift components, light hydrocarbons and Argon tracer.
To produce sufficient material for all steps for the process, Azores 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 2sthe completion of essay 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. Tern grooms of the reactor proc~Llct it extracted with three 150 ml washes of tetratIydrof-lrarJ (Ralph) in centrifuge bottles. Each wash is centrifuged anti the liquid decanted. The solids are dried and weighed. The TflF
insoluble material correspond to unrequited carbonaceous 3smaterial. The TOUGH soluble fraction is rotovaped and the ~226838 THY removed. The remaining liquid is washed with Tulane in the same manner as the TO The Tulane insoluble solids correspond to high molecular weight material. The Tulane soluble material is rotovaped and 5 washed with pontoon in the same manner as the THY. The pontoon insoluble material is lower molecular weight material and some polar material. The pontoon soluble material corresponds to oils or low boiling (760F) compounds. The MA (moisture and ash free) and DM~IF
(dry mineral matter free) yield structure was:
MA DMMF
(wit %) (wit %) Yield 82.9 95.8 15 Tulane Ins. 29.2 29.2 Pontoon Ins. 3G.7 36.7 Pontoon 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 r[lixed together. A summary of the material balance for these experiments is shown in Table 2. The data indicate that over 80% of the MA coal is converted to a liquid product (81.1% calculated from cake weight, 85.6% calculated from filtrate weight).
This step also indicates that 85.6~ of the MA 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), Horatio Catalyst 601-T (50 g), and Nazi (25 g) were mixed together in the reactor. The Horatio Tao catalyst is a Cobolt-Molybdenum catalyst. The Nays was added to increase the activity of the catalyst. The hydrogenation experiment was performed at 400 C
(725F) for a period of 60 minutes at temperature. A
system pressure of 13.90 Ma (2012 Asia) 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 SUE. The decant was analyzed by SUE to determine pontoon 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 TIFF insoluble material was in the decant bottoms. Overall there was a 55% reduction in the TOUGH insoluble, a 100% reduction in the Tulane insoluble, a Go% reduction in pontoon insoluble and a 56% increase in the pontoon soluble.
For the ionic liquefaction process this would result in 56.5 lb. of Entwine Solubles for every 100 pound MA
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 to 1.4 wt. % of the MA coal fed.

. . .

l~Z~i~3~

Reactor Contents:
Lignite .180 kg ~2 .025 kg Naomi .018 kg m-Cresol .260 kg Tetraline .060 kg Naphthalene .020 kg l-Methylnaphthalene .010 kg gamma-picoline .010 kg Gas Type: Corey (95:5) Pressure: 1500 Asia (10.35 Ma) 15 Temperature 335C (635F) Gas Flow Rate: 3 loin (5 x 10 5 m us) Residence Time: 60 minutes 20 Lignite Analysis Component At (wit %?
Lowe 28.2 Ash 6.0 C 48.8 Al 3.3 0.7 S 0.6 12.
Heating Value Bulb 8,550 Jo ANALYSIS OF l)IsrrII.L~TION BOTTOMS
AND IIYDROGEN~TED PRODUCT
- . .
hydrogenated ~ottomsProduct Kinematic Viscosity cyst 18.5 cyst Conrad son Carbon 17.10 wit % 9.58 wit APT Gravity -1.7 APE APE
THY Ins. 19.4 g8.7 g Tulane Ins. 73.1 g0.0 g Pontoon Ins. 127.0 g44.9 g Pontoon Sol 268.7 g418.1 g Gas -- 4.2 g , . . . _ . . , _ .

~Z~38 SIJMM~RY OF Prosier GENERATION RUNS*

Grams Grams Grams Grays l)MMF Organic In organic Run No. MY 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 15 *All values normalized to 100% mass balance to account for transfer losses.

, Jo .. _ .. _ .. . .... _ _ _ .. .

1~26;838 EXAMPLE II
A suite of six coals were experimentally tested to determine suitability as ionic liquefaction feed stocks.
The coals ranged in rank from an Australian brown coal 5 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 10 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 feed stocks because of large oil and gas yields.

i;

~22~838 Q o o o o o o o o I
CO Us o r5~ o Us o m m ED I , rho I o C o .

-I I r CO
Us OX o o o o o _ 0 o o I
Z
do O O O O Jo O
Al Al .~.
o ' Rex) O I

Jo I I` rho I;
It an I r , O
. . . .
0 ED r.
a o En . . r.
I Rex us X 3: ?
o o on I
m m c I
Us lo I rrJ
o ' O 1) 3 G) 9 U I NO to I rJ~)~ C I
O 'I or) to) I 'I I O Owe O I Al h r4 Jo O W
a H

do U-) o or u) I CO
X r Jo 1;` 0 -I

l~Z6~38 R FEED TEMPT D~MF YIELDS, WIT%*
C Total PA AGO
Yield s 134 Brown Coal 310 96.728.6 27.1 41.0 135 Brown Coal 322 94.224.4 27.8 42.0 140 Colstrip 335 75.5 27.831.9 15.8 10 141 Illinois No. 6 335 94.8 45.225.7 23.9 143 awoke Mesa 335 87.131.9 36.1 19.1 188 South Al. 335 86.456.5 20.9 9.0 175- Texas 15 179 Lignite 335 95.3 28.535.4 31.5 ' * 80~ of Nails Format lZ26838 n alternative process has been developed to salivate 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 environmental 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 (300C to 360C) 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 pow (300 to 1500 Asia).
After salvation 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 undissolved 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 (H2C03), Hal, H2S04 or the like. Dissolved inorganic compounds will be removed because of their greater affinity for the aqueous layer as opposed to the organic filtrate.

~zz~

_ 46 _ 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 J 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 275C, more preferably under about 300C and most preferably under 325C. The bottoms from the distillation are collected and cooled to form a solid product which has 15 reduced sulfur and ash.
Sulfur can be removed in a variety of ways. The inorganic fraction which is mainly pyrites will be removed as an insoluble material in the filtration step after the salvation of the carbonaceous material.
20 Organic sulfur is removed by reaction with alkali and base to form such compounds as alkali sulfides, sulfites, and sulfates. Tao resulting sulfur compounds are either removed during the filtration step or during the acid extraction of the filtrate depending on the 25 volubility 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 30 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, comlninutes the carbonaceous material (stream 100) by conventional means, as 35 previously described; and adds a water-alkali mixture, 1~2~38 stream (300); and recycle polar solvent, stream (200) containing greater than 50~ by weight of finlike 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, commented 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 finlike compounds, and more preferably greater than 60~ by weight finlike compounds.
The preferred amount of alkaline material in stream US (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 amulets of finlike 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 sty _ 48-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 lo and 200 parts by weight. Make up alkali, (stream AYE) may be added, as before.
The liquefaction section (By) is the same or similar 10 to that shown in (B) Fig. l. 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 15 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 20 described.
The reaction mixture is then brought to the desired temperature. Roy reaction mixture stream is then transferred to the liquefaction reactor (By) and held at the desired reaction temperature for sufficient time to 25 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 30 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. The preferred temperature range is from about 250C to 35 360C (482 F to 680 F). System pressure should be ~2Z6~38 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 5 reaction. The preferred pressure range is from about 0.69MPa to 13.80MPa (100 Asia to 2000 Asia), and the most preferred pressure range is about 2.76MPa to 8.28MPa (400 Asia to 1200 Asia).
The reaction products leaving the reactor (By), in stream (606) of Fits. 9 and 10, are then separated into .
component streams in the separation system, process (Of) 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 pressure, preferably between about 0.69 and 3.45 Ma (100 - 500 Asia), and the temperature is kept at conditions sufficient to keep most of the volatile organic compounds in the liquid phase, but much of the running water as vapor. The preferred temperature range is 200 to 300 C (392 to 572 F); the most preferred~emperature range is 200 to 250C (392 to 482F). 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 (Al) for separation into a flue gas stream Go and a sulfur stream (So).
As previously described, leaf filters, candle filters, hydroclones, or comparable equipment can be used for the separation of tile solids as well as separation by processes such as solvent dashing, or critical solvent extraction. The purpose of the filtration is to separate undissolved carbonaceous ~2~38 material, ash and alkaline salts from the ionic liquefaction products which are liquid under the conditions employed. Due to the hydroscopic nature of the alkaline salts, much of tile water present will also 5 be separated with the filter cake.
The preferred temperature for the separation step (Of) is preferably between about 150 and 100 C !302 and 212F). The pressure is maintained at a sufficient level to obtain efficient separation, 10 preferably between about 0.34 and 1.03 Ma (50 to 150 Asia). 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 weight, 15 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 equipmetlt design. The separation train should 20 preferably be operated to reduce the alkaline compound content ox the stream (609), of jig. 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) One Fig. 10 colltaining added phrasal water in a countercurrent method while bobolink C02, steam (611) througtl the solution in a mixing vessel (Ml) as shown in Fig. 11. The resulting liquids stream (615) is passed to a liquid/liquicl settler (So) for separating the liquids stream into an organic component and an aqueous component. The organic component stream (613) ~;~Z~3~

of jig. 11 will have the alkaline arid 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 5(609) and water stream (610) are contacted in mixing vessel (Ml) in a feed ratio of one part filtrate (609) to 5 parts Eye, more preferably one part feed to 3 parts Ho and most preferably one part feed to one part ~32 The preferred temperature is about 100 C
(212F), more preferably less than about ~30C
(176 Phoned most preferably less than about 50 C
(122 F). The liquid streams and COY will preferably be mixed for a period of 15 minutes, and more preferably for a period of between 5 and 10 minutes. The carbon dockside 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 (So) and sent to a distillation unit (If) (Fig. 10) where the organic rich 20 stream is fractionally distilled. A light oil fraction is collected as a product stream (626). Distillate product boiling between 200C and 325C (392F and 617F) is collected as a recycle solvent stream (200) and returned to the feed preparation (Al). The high 25 boiling fraction is collected and cooled to about 25C. 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~, 30 and most preferably below 0.3~. the solids rich filter cake (607) is washed with a water stream at solids washing unit (Eel). The water stream is preferably at a temperature of ~0C. and at a ratio to the cake of 5 to 1, preferably 3 to 1, and most preferably 1 to 1.
35 The water stream will remove liquid organic associated ~.~2~38 with the filter cake and the water-organic stream (610) and is sent to the solvent extraction system (Do).
Referring further to the integrated process in Fig.
10, the washed filter cake produced in the solids washing unit (Fly), is used as a feed stream (614), to a gasifies, unit (Go). 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 KIWI ration of gas leaving the gasifer, stream (615) is between 0.5 and 2Ø Acid gases and light hydrocarbon gases are separated from the synthesis gas by standard operations in gas processing operations, unit (Hi).
The ash from the gasifies, stream (~16), is sent to 15 an alkali recovery unit (Jo) 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 20 compounds is greater than So 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 Luke (77 and 212F), 25 and more preferably between 50 and 100C (122 and luff). The preferred treatment rate is less titan 4 kg H2O/kg ash, more preferably less than 2 kg H2O/kg ash, and most preferably less than lo kg H2/kg ash.

122~838 EXPEL 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 feed stocks. An analysis of each feed stock is presented in Table 6. The ionic liquefaction conditions for each feed stock 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 15 unrequited 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 20 filtrate by extraction with a 10% hydrochloric acid solution. The extractions were performed in a separator funnel by mixing the filtrate three separate times with fresh acid solution and collecting the filtrate (organic fraction). The extracted filtrate was 25 distilled to remove the material boiling below 300~C.
The bottoms from this distillation was the solid product and was collected. Any distillate collected in excess of the solvent was coal-cierived distillate and recoverable as a liquid proc3uct stream Results from 30 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.

~226838 0 0 ED a a _- o o It En 0 O ED Lo a) ED
o o a Lo 0 Lo 0 o Jo o Us o o o o o o o o on go O r o o o o o I
O
En us I d' Lo ED CO Lo 0 Lo 3 Jo 1` o I
::~ _ O o O I: I n U O U Lo I; ..
o o It CO O
En Hi I: , u u 2 I co ox o o o o o us Lo d' I) O Ox d' 0 Jo I ON Lo t` O

Us 0 O 0 pa cay o ox us 0 CAL Us a) x Jo o , J 0 a o U Jo o o 0 o a 3 C1~ -I V

0 0 arc rho 6 a) V 11~ U
I V O 0 3 -1 3 w I"
Lo 1 U I. 0 I) do do 0 0 Al O ) O I O O h 0 Lo Lo u zip u m u ,1 I
Q) a -I ED I
Lo I I to N

by -- so --o :~, V

do It O O O O I Us O O
on ED O I ' ED

P.

o Jo I o I) Lo ED O or o o It I o I VC
O
UP: V
Ox O
OX
ho ~-_~ 0 0 I c I

h O U O ED o o I 0 o Eye O us o o o I o o P. 3 ! or I> a "
O Us --lea '`'`
O us ED O ' ' I I

V
V , JO O Us Lo X O O 1`1` O
O I Us ED O ' ' ., _ I
-- O
X C) h O ^ U) I:: C W
u --I O v 6 ~1 o r it V
O o Us I O to Owe 0 a G us I o I I I h Jo CJ I h Z:Z T: I Jo ~2683~

Us O
JO
O O
-It` ox 0 ICKY ED I) Nat) O O d' N

or O 'I DUD O Ox`
I ' N ' O
' 3 I N Us O O O

C:

D o N CO O

En o O O O
I:; a ':', O Us O ' ' W Pi N ED O O JO
X

O I) No I pa Ed' ' N O
O I O O

a) I JO Al O I Nay Z co O O a a) o Us C U

3 3 3 3 Jo 12;~ 338 EXAMPLE IV
A suite of coals have been examined to determine their suitability as feed stocks to produce a solid product from the ionic liquefaction process. Analysis 5 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 10 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.
15 Analyses and reaction conditions of the experiment appear in Table 9 and 10. The 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.
20 Carbon dioxide gas was bubbled through the agitated solution fur a period of one hour. After one hour the organic fraction was removed in a separator funnel and distilled. The results of experiments are given in Table 11.

CO OX Do O
o n d' N Ox 0 a m I okay o own o JO Do .
O
O In I 1`
0 Lo O ED
.
Us o o o o o 0 I Do DO N O I
Z O O O O
a W
aye I 1`
us us In ox :~: . . . . .
U it' O _ X
ox Us I_ O Jo r Ox O U Do or I
.,1 1 O Jo .
m u us O or) I` Us In o ox n o us O I gin I o us In In In on it: n zip o I; to Jo n o to) Nil to a) to Us pa arc pa Jo C Al I O -r1 O O D Z C
C) X Al Al ill to U O (I tnr1 o pa c a) a I o r1 awry pa I Z O
OUT I Al I Err I Us Roy c o aye u70 I
3 Al r I C ~11 r1 r-l X
~1(1) O r1O r1 .r1 Us (U Al I I a) C ) 1 4 Z U Eli Z Jo O

rl a) QlE3 I l I N In ye do I
Z N to to to to sty 3 REACTION CONDITIONS

North Dakota Color Rawhide Carbonaceous Feed: LigniteWadge Mine Usably .. .. _ ..
Experiment # 241 243 259 261 Amount (g) 150 150 150 150 Conditions:

15 m-cresol (g) 360 360 360 360 Tetralin (g) 90 90 60 60 Noah (g) 3.753.7515 15 ;
Nikko (g) 3.75 3.75 ___ ___ HO (g) 10 20 10 20`

25 Pressure (Ma) 10.9711.2110.769.31 Temperature ( C) 340 350 340 340 Residence Time (mint 90 90 90 I

I

TABLE 10 (continued) . _ .

REACTION CONDITIONS

Ohio Illinois Texas Carbonaceous Feed: No. 6 No. 6 Lignite . . . _ . . _ . _ 10 Experiment # 247 242 245 Amount (g) 150 150 150 Conditions:

m-cresol (g) 360 360 360 Tetr~lin (g) 60 60 90 2 Noah (g) 3.75 3.75 3.75 Nikko (g) 3,75 3,75 3,75 H20 (g) 20 20 10 Pressure (Ma) 6.49 12.56 9.59 Temperature ( C) 360 360 360 Residence Time (mix) 90 90 90 ;l;~;Zti838 EXPERIMENTAL RESULTS

North Dakota Color Rawhide Carbonaceous Feed: Lignite Wedge Mine Usably . _ .
Experiment # 241 243 259 261 Amount DMMF Feed (g) 85.9 124.2 94.6 105.8 wit Solid Product 48.3 34.7 41.8 41.5 wit% S 0.22 0.23 0.00 0.02 White Ash 0.19 0.03 0.05 0.20 wit% Distillate 15.1 21.8 7.6 11.5 White Unrequited 39.4 41.3 49.5 47.2 l;~Z~i838 TABLE 11 (continued) EXPERIMENTAL RESULTS

Ohio Illinois Texas Carbonaceous Feed: No. 6 No. 6 Lignite - '-Experiment # . 247 242 245 Amount DMMF Feed (g) 124.5 113.8 150 15 wit % Solid Product 56.7 62.6 53.6 wit% S 0.71 1.31 0.02 White Ash 0.10 0.80 0.06 White Distillate 7.5 4.0 .

White Unrequited 33.0 29.9 46.4 EXAMPLE VI
A hvBb coal obtained from the Ohio No. 6 Seam it 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 3609. 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 So argon and 95% carbon monoxide. The gas delivery system is equipped with pressure regulators, and flow controllers to maintain 1012 Asia (6.98 Ma) 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 connected to an Omega Model AYE
temperature'lndicator. The heater temperature is controlled by a Phenol Series 5501552 temperature controller. Gas flow enters the reactor through the stirrer and exits through a knock back condenser consisting of a 3/4-inch OLD. 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 volume, a Carte Series "S" chromatography for analysis of carbon monoxide, carbon dioxide, hydrogen and argon tracer, and a llewlett-Packard 3390 integrator to calculate and print the gas composition in mole percents. The semi-batch liquefaction reactor system is pressure tested at 1012 Asia (6.98 Ma) with helium and then the premixed argon and carbon monoxide gas is 35 introduced and the reactor is heated to 300 C (573K).
* Trade Mark -` ~Z26838 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 (THY), Tulane, and pontoon to determine total conversion, preasphaltenes, asphaltenes, and oil plus gas.
The yield and product structure are defined by:
XV. Yield =
100 - Grams MA THY Insoluble Material (100), [=] wit%
Grams MA Coal XVI.~ Preasphaltenes (PA) =
Grams MA Tulane Insoluble Material (100), . _ Grams MA Coal [=] wit %

XVII. Asphaltenes (A) =
Grams MA Pontoon Insoluble Material (100), Grams MA Coal t=] wit %

XVIII. oil plus Gas (0 G) = Yield - PA - A, [=] wit %

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:
O represents _ conversion, calculated from CO in the exit gas O represents _ conversion, calculated from COY
in the exit gas, using water gas shift storchiometry, and represents, CO conversion, calculated from Ho 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, 3609. of m-cresol, 409. of water, 1012 Asia (6.98 Ma), 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 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 1009. of the hvBb Ohio No. 6 coal, 360g. of m-cresol, 40g. of water, 259. of potassium hydroxide, 1012 Asia (6.98 Ma), 300C (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.

~LZ26~338 ..

Carbon monoxide conversion is calculated and plotted, as before, and the results shown in Fig. 14.

MA
Example Solvent/Solute System Conversion organic Phase Inorganic Phase Water Solubilizing Alkali/Alkaline-Agent Earth Compound Wit%
VI 360g. m-cresol 0 0 40.0 VII 360g. m-cresol 0 40g. 39.0 10 VIII 360g. m-cresol 25g. KOCH 40g. 82.5 EXAMPLE IX
The foregoing procedure of Example VI is repeated using 100g. of a hvCb Colorado Wedge 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 10g. of l-methylnaphthalene, 40g. of water, 1012 Asia (6.98 Ma), 300C (573K), 0.5 SUM of 95 percent carbon monoxide and 5 percent argon, 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 Wedge coal, 40g. of water, 15g. of sodium hydroxide, and the temperature, pressure, 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.

Jo lZ2~838 Example Solvent/Solute System Organic Phase Inorganic Phase MA
Solubilizing Alkali/Alkaline- Conversion Agent Earth Compound Water Wit %
IX 360g. synthetic solvent 0 40g 22.8 X 360g. synthetic solvent 15g Noah 40g 60.5 XI 360g. synthetic 15g Nikko 40g 60.1 solvent The results in Tables 12 and 13 demonstrate the significantly improved results obtained by practice of 15 the present invention. Table 13 shows that the presence of the organic phase solubilizing agent, m-cresol, in the absence of the inorganic phase constituent as Example VI, yields a MA conversion of 40 wit %. In the case of Example VII with the addition of water, the MA
20 conversion-is 39 wit %, which is virtually unchanged from Example VI. on Example VIII, 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 MA wit %. 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 MA wit % 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.

~Z26838 of synthetic solvent consisting of 160g. of m-cresol, 160g. of tetrahydronaphthalene, 20g. of naphthalene, 10 g. of l-methylnaphthalene, and 10g. of gamma-picoline.
The inorganic fraction of the solvent/solute system is 5 30g. of water 18g. of Noah, 4.5g. of Nikko, and 30g. Nays OWE. The feed materials are reacted at 340C, 1312 Asia (8.99 Ma), 0.5 SUM of 95 percent carbon monoxide and 5 percent argon for 30 minutes. The results of the selective solvent extraction procedure 10 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 3609. of a synthetic solvent consisting of 2609. of m-cresol, 60g. of tetrahydronaphthalene, 20g.
of naphthalene, 109. of l-methylnaphthalene, and 10 g.
of gamma-picoline. The results of the selective solvent extraction procedure are shown in Table 14.

Example MA Conversion XII I; 66.3 ZOO 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 finlike compounds in the organic fraction of the solvent/solute system. In Examples XII, where m-cresol is 44.4 percent of the I organic fraction, 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.

.. . ....

Jo Z2t~838 Minerals (Wit%) Coal Foe NATO Kiwi Aye Sue Coo Moo _ Eagle #57.21 1.891.47 24.5 55.2 3.83 1.79 Ohio #618.30 .812.70 25.1 51.2 - -Wedge 4.09 .62 .84 27.4 60.5 4.23 .79 The foregoing description of several embodiments of an integrated ionic liquefaction process can, of course, be modified by adding steps or combining operations if desired to achieve different specific results from those 20 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 (18)

The embodiments of the invention in which an exclusive property or privilege is claimed are defined as follows:

1. 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 which comprises:
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.

2. A method according to claim 1, wherein said organic phase solubilizing agent is selected from o-cresol, m-cresol, p-cresol, catechol, resorcinol, naphthol and mixtures and derivatives thereof.

3. A method according to claim 1, wherein 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.

4. A method according to claim 3, wherein said organic constitutent is selected from naphthalene, anthracene, phenanthrene, acenaphthalene, 1-methylnaphthalene, 2-methylnapthalene, tetrahydronaphthalene, gama-picoline, isoquinoline, dihydronaphthalene, decahydronaphthalene, 9,10-dihydroanthracene, 9,10-dihydrophenanthrene and mixtures and derivatives thereof.

5. A method according to any of claim 1, wherein said organic fraction is in whole or in part derived from liquified carbonaceous material.

6. A method according to claim 1, wherein said compound in the inorganic phase is selected from alkali hydroxides, alkali carbonates, alkali bicarbonates, alkali nitrates, alkali sulfates, alkali sulfites, alkali sulfides, alkali formates and other alkali salts, alkaline-earth hydroxides, alkaline-earth carbonates, alkaline-earth bicarbonates, alkaline-earth sulfites, alkaline-earth sulfides, alkaline-earth-formates and other alkaline-earth salts, and mixtures thereof.

7. A method according to claim 1, wherein said compound in the inorganic phase is selected fom 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.

8. A method according to claim 6, wherein 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.

9. A method according to claim 1, wherein the aqueous phase includes water present in an amount from about 5 parts to about 60 parts per 400 by weight of the solvent/solute system.

10. A method according to claim 9, werein the aqueous phase 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.

11. A method according to claim 1, wherein said solubilizing agent has a boiling point from about 50°C
to about 400°C and is present in an amount from about 50 to about 100 wt. percent of the organic fraction of the solvent/solute system.

12. A method according to claim 11, wherein said solubilizing agent is present in an amount from about 50 to about 75 percent by weight of the organic fraction of the solvent/solute system.

13. A method according to claim 9, wherein said contacting takes place at a temperature of from about 100°C to about 400°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.

14. A method according to claim 13, wherein said temperature is from about 140°C to about 360°C and said pressure is at least about 300 psia.

15. A method according to claim 14, wherein said temperature is from about 260°C to about 360°C.

16. A method according to claim 14, wherein said pressure is from about 500 psia to about 1500 psia.

17. 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, the method consisting essentially of:
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, 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, 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 400°C and a pressure of at least about 300 Asia.

18. The method of claim 17, wherein the carbonaceous material is contacted with the solvent/solute system in the presence of carbon monoxide.