Note: Descriptions are shown in the official language in which they were submitted.
.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
--2--
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
838
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.
838
-- 5
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
-- 10 --
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
3 I
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
- 12 -
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.
38
- 13 -
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
- 16 -
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
- 17 -
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
- 18 -
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~
-- 19 --
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)
12Z6838
_ 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
- 22 -
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
- 23 -
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).
1226838
- 24 -
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
- 25 -
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
1226838
- 26 -
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
- 27 -
(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
- 28 -
(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
- 30 -
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
- 33 -
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).
1226838
- 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
- 35 -
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
36 -
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
- 37 -
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
- 38 -
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~
- 39 -
TABLE 1
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
38
Jo
TUBE 2
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
- 41 -
TABLE 3
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
- 42 -
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
- 43-
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 .
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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
0
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
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a H
do U-) o or u) I CO
X r Jo 1;` 0
-I
l~Z6~38
- 44 -
TABLE 5
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
- 47 -
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
- 49 -
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
- 50 -
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~
- 51 -
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
- 53 -
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
-- 54 --
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~2683~
-- 56 --
Us
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12;~ 338
- 57 -
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.
1226838
-- 58 --
CO OX Do O
o n d'
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C ) 1 4 Z U Eli Z Jo O
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ye do I
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sty 3
- 59 -
TABLE 10
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
- 60 -
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
- 61 -
TABLE 11
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
35
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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
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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).
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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
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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.
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Carbon monoxide conversion is calculated and
plotted, as before, and the results shown in Fig. 14.
TABLE 12
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
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unit is heated, charged, and the products analyzed as
previously described. The results of the selective
solvent extraction procedure are shown in Table 13.
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TABLE 13
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.
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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.
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.
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TABLE 15
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.
35
