Note: Descriptions are shown in the official language in which they were submitted.
1 1'7 ~1~Z
6117B April 6, 1981
IMM08ILIZATION OF VANADIA DEPOSITED
ON CATALYTIC MATERIALS DURING
CARBO-METALLIC OIL CONVERSION
~escription
Technical Field
This invention relates to processes for converting
heavy hydrocarbon oils into lighter fractions, and especial-
ly to processes for converting heavy hydrocarbons containing
high concentrations of coke precursors and heavy metals into
gasoline and other hydrocarbon fuels.
~ackground Art
The introduction of catalytic cracking to the petroleum
industry in the 1930's constituted a major advance over pre-
L5 vious techniques with the object of increasing the yield ofgaseoline and its quality. Early fixed bed, moving bed, and
fluid bed catalytic cracking FCC processes employed vacuum
gas oils (VGO) from crude sources that were considered sweet
and light. The terminology of sweet refers to low sulfur
content and light refers to the amount of material boiling
below approximately 1000-1025F.
The catalysts employed in early homogenous fluid dense
beds were of an amorphous siliceous material, prepared syn-
thetically or from naturally occurring materials activated
by acid leaching. Tremendous strides were made in the
1950's in FCC technology in the areas of metallurgy, pro-
cessing equipment, regeneration and new more-active and more
stable amorphous catalysts. ~owever, increasing demand with
respect to quantity of gasoline and increased octane number
requirements to satisfy the new high horsepower-high com-
pression engines being promoted by the auto industry, put
extreme pressure on the petroleum industry to increase FCC
capacity and severity of operation.
A major breakthrough in FCC catalysts came ln the early
1960's with the introduction of molec-llar sieves or zeo-
lites. These materials were incorporated into the maerix of
li74192
amorphouY and/or amorphous/kaolin material~ constituting
the FCC catalyst~ of ~hat time. These new zeolitic cata-
lyst~, containing a crystalline aluminosilic~te zeolite in
an amorphous or amorphous/kaolin matrix of silica, alumina,
silica-alumina, kaolin, clay or the like were at least lO00-
lO,000 times more active for cracking hydrocarbons than the
earlier amorphous or amorphous/kaolin containing silica-
alumina catalysts. This introduction of zeolitic cracking
catalysts revolutionized the fluid catalytic cracking pro-
cess. Innovations were developed to handle these high ac-
tivities, such as riser cracking, shortened contact times,
new regeneration proce~ses, new improved zeolitic catalyst
developments, and the like.
The new catalyst developments revolved around the de-
velopment of various zeolites such as synthetic types X andY and naturally occurring faujasites; increased thermal-
steam (hydrothermal) stability of zeolites through the in-
clu~ion of rare earth ions or ammonium ions via ion-exchange
technique~; and the development of more attrition resistant
matrices for supporting the zeolites. These zeolitic cata-
lyst developments gave the petroleum industry the capability
of greatly increasing throughput of feedstock with increased
conversion and selectivity while employing the same units
without expansion and without requiring new unit construc-
tion.
After the introduction of zeolite-containing catalysts
the petroleum industry began to suffer from a lack of crude
availability as to quantity and quality accompanied by in-
creasing demand for gasoline with increasing octane values.
The world crude supply picture changed dramatically in the
late 1960's and early 1970's. From a surplus of light,
sweet crudes the supply situation changed to a tighter sup-
ply with an ever-increasing amount of heavier crudes with
higher sulfur contents. These heavier and higher sulfur
crudes presented processing problems to the petroleum re-
finer in that these heavier crudes invariably also contained
much higher metals and Conradson carbon values, with accom-
1174192
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panying significantly increased asphaltic content.
Fractionation of the total crude to yield cat cracker
charge stockq also required much better control to ensure
that metals and Conradson carbon values were not carried
overhead to contaminate the FCC charge stock.
The effects of heavy metals and Conradson carbon on a
zeolite-containing FCC catalyst have been described in the
literature as to their highly unfavorable effect in lowering
catalyst activity and selectivity for gasoline production
and their harmful effect on catalyst life.
These heavier crude oils also contained more of the
heavier fractions and yielded a lower volume of the high
quality FCC charge stocks which normally boil below about
1025F and are usually processed so as to contain total me-
tal levels below 1 ppm, preferably below 0.1 ppm, and Con-
radson carbon values substantially below 1Ø
With the increasing supply of heavier crudes, Which
yield les~ gasoline, and the increasing demand for liquid
transportation fuels, the petroleum industry began a search
for processes to utilize these heavier crudes in producing
gasoline. ~any of these processes have been described in
the literature and include Gulf's Gulfining and Union Oii's
Unifining processes for treating residuum, UOP's Aurabon
process, Hydrocarbon Research's ~-Oil process, Exxon's Flexi-
coking process to produce thermal gasoline and coke, H-Oil's
Dynacracking and Phillip's Heavy Oil Cracking (HOC) proces-
ses. These processes utilize thermal cracking or hydro-
treating followed by FCC or hydrocracking operations to han-
dle the higher content of metal contaminants (Ni-V-Fe-Cu-Na)
and high Conradson carbon values of 5-15. Some of the draw-
backs of these types of processing are as follows: coking
yields thermally cracked gasoline which has a much lower oc-
tane value than cat cracked gasoline, and is unstable due to
the production of gum from diolefins, and requires further
hydrotreating and reforming to produce a high octane ?ro-
~uct; gas oil quality is degraded due to thermal reactions
which produce a product containing refractory ?olvnuclear
ii741~2
aromatics and high Conradson carbon levels which are highly
unsuitable for catalytic cracking: and hydrotreating re-
quires expen~ive high pressure hydrogen, multi-reactor sys-
tems made of special alloys, costly operations, and a sepa-
rate costly facility for the production of hydrogen.
To better understand the reasons why the industry hasprogressed along the processing schemes described, one must
understand the known effects of contaminant metals tNi-V-Fe
Cu-Na) and Conradson carbon on the zeolite-containing crack-
ing catalysts and the operating parameters of an FCC unit.Metal content and Conradson carbon are two very effective
restraints on the operation of an FCC unit and may even im-
pose undesirable restraints on a Reduced Crude ConverSiOn
(RCC) unit from the standpoint of obtaining maximum conver-
sion, selectivity and catalyst life. Relatively low levelsof these contaminants are highly detrimental to an FCC unit.
As metals and Conradson carbon levels are increased still
further, the operating capacity and efficiency of an RCC
unit may be adversely affected or made uneconomical. These
adverse effects occur even through there is enough hydrogen
in the feed to produce an ideal gasoline consisting of only
toluene and isomeric pentenes (assuming a catalyst with
such ideal selectivity could be devised).
The effect of increased Conradson carbon is to increase
that portion of the feedstock converted to coke deposited on
the catalyst. In typical VGO operations e~ploying a zeolite-
containing catalyst in an FCC unit, the amount Or coke de-
posited on the catalyst averages around about 4-S wt~ of the
feed. This coke production has been atrributed to four dif-
ferent coking mechanisms, namely, contaminant coke from ad-
verse reactions caused by metal deposits, catalytic coke
caused by acid site cracking, entrained hydrocarbons resul-
ting from pore structure adsorption and/or poor stripping,
and Conradson carbon resulting from pyrolytic distillation
of hydrocarbons in the conversion zone. There has been pos-
tulated two other sources of coke present in reduced cruces
in addition to the four present in VGO. ~hey are: (1) ad-
1 1'~"~192
sorbed and absorbed high boiling hydrocarbons which do notvaporize and cannot be removed by normally efficient strip-
ping, and (2~ high lecular weight nitrogen-containing hy-
drocarbon compounds adsorbed on the catalyst's acid sites.
Both of these two new types of coke producing pheDomena add
greatly to the complexity of resid processing. Therefore,
in the processing of higher boiling fractions, e.g., re-
duced crudes, residual fractions, topped crude, and the
like, the coke production based on feed is the sum-ation of
the four types present in VGO processing (the Conradson car-
bon value generally being much higher than for VGO), plus
coke from the higher boiling unstrippable hydrocarbons and
coke associated with the high boiling nitrogen-containing
molecules which are adsorbed on the catalyst. Coke produc-
tion on clean catalyst, when processing reduced crudes, maybe estimated as approximately 4 wt3 of the feed plus the
Conradson carbon value of the heavy feedstock.
The coked catalyst is brought back to equilibrium ac-
tivity by burning off the deactivating coke in a regenera-
tion zone in the presence of air, and the regenerated cata-
lyst is recycled back to the reaction zone. The heat gen-
erated during regeneration is removed by the catalyst a~d
carried to the reaction zone for vaporization of the feed
and to provide heat for the endothermic cracking reaction.
The temperature in the regenerator is normally limited be-
cause of metallurgical limitations and the hydrothermal sta-
bility of the catalyst.
The hydrothermal stability of the zeolite-containing
catalyst is determined by the temperature and steam par-
tial pressure at which the zeolite begins to rapidly loseits crystalline structure to yield a low-activity amorphous
material. The presence of steam is highly critical and is
generated by the burning of adsorbed and absorbed (sorbed)
carbonaceous material which has a significant hydrogen con-
tent (hydrogen to carbon atomic ratios generally greaterthan about 0.5). This carbonaceous material is ?rlnci?ally
the high-boiling sorbed hydrocarbons with boillng ?oints as
~17'~1~Z
high as 1500-17000F or above that have a modest hydrogen
content and the high boiling nitrogen containing hydro-
carbons, as well as related porphyrins and asphaltenes.
The high molecular weight nitrogen compounds usually boil
above 1025F and may be either basic or acidic in nature.
The basic nitrogen compounds may neutralize acid sites while
those that are more acidic may be attracted to metal sites
on the catalyst. The porphyrins and asphaltenes also generally
boil above 1025F and may contain elements other than carbon
and hydrogen. As used in this specification, the term
"heavy hydrscarbons" includes all carbon and hydrogen compounds
that do not boil below about 1025F, regardless of the presence
of other elements in the compound.
The heavy metals in the feed are generally present as
porphyrins and/or asphaltenes. However, certain of these
metals, paxticularly iron and copper, may be present as the
free metal or as inorganic compounds resulting from either
corrosion of process equipment or contaminants from other
refining processes.
As the ~onradson carbon value of the feedstock
increases, coke production increases and this increased load
will raise the reseneration temperature; thus the unit may
be limited as to the amount of feed that can be processed
because of its Conradson carbon contents. Earlier VGO units
operated with the regenerator at 1150-1250 F. A new development
in reduced crude processing, namely, Ashland Oil's "Reduced
Crude Conversion Process", as described in pending Canadian
applications 364,647; 364,655, 364,665, and 364,666,
all filed November 14, 1980, can operate at regenerator
temperatures in the range of 1350-14000F. But even
these higher regenerator temperatures place a limit on the
Conradson carbon value of the feed at approximately 8,
which represents about 12-13 wt% coke on the catalyst based
~n the weight of feed. This level is controlling unless
considerable water is introduced to further control
temperature, which addition is also practiced in Ashland's
RCC processes.
i~
11'7~ Z
The metal-containing fractions of reduced crudes con-
tain Ni-V-Fe-Cu in the form of porphyrins and asphaltenes.
These metal-containing hydrocarbons are deposited on the
catalyst during processing and are cracked in the riser to
deposit the metal or are carried over by the coked catalyst
as the metallo-porphyrin or asphaltene and converted to the
metal oxide during regeneration. The adverse effects of
these metals as taught in the literature are to cause non-
selective or degradative cra~ing and dehydrogenatic,n to
produce increased amounts of coke and light gases such as
hydrogen, methane and ethane. These mechanisms adversely
affect selectivity, resulting in poor yields and cuality of
gasoline and light cycle oil. The increased production of
light gases, while impairing the yield and selectivity of
the process, also puts an increased demand on gas compressor
capacity. The increase in coke production, in addition to
its negative impact on yield, also adversely affects cata-
lyst activity-selectivity, greatly increases regenerator
air demand and compressor capacity, and may result in uncon-
trollable and/or dangerous regenerator temperatures.
These problems of the prior art have been greatly mini-
mized by the development at Ashland Oil, Inc., of its Re-
duced Crude Conversion (RCC) Processes described in the co-
pending applications reference above and incorporated here-
in by reference. The new processes can handle reducedcrudes or crude oils containing high metals and Conradson
carbon values previously not susceptible to direct proces-
sing-
It has long been known that reduced crudes with high
nickel levels present serious problems as to catalyst !e-
activation at high metal on catalyst contents, e.g., 5000-
lO,OOO ppm and elevated regenerator temperatures. It has
now been recognized that when reduced crudes with high vana-
dium levels are processed over zeolite containing catalvsts,
especially at high vanadium levels on the catalyst, ra?id
deactivation of the zeolite can occur. ThiC deactiva~ion
manifests itself as a loss of zeolitic structure. This loss
117'~192
-8-
has been observed at vanadium levels of 1000 ppm by weight
or less. This loss of zeolitic structure becomes more ra-
pid and severe with increasing levels of vanadium and at
vanadium levels about 5000 ppm, particularly at levels ap-
proaching 10,000 ppm complete destruction of the zeolite mayoccur. Prior to the present invention, it was believed im-
possible to operate economically at vanadium levels higher
than 10,000 ppm because of this phenomenon. Previously, de-
activation of catalyst by vanadium at vanadium levels of
less than lO,000 ppm has been retarded by lowering regenera-
tor temperatures and increasing the addition rate of virgin
catalyst. Lowering regenerator temperatures has the disad-
vantage of requiring higher catalyst to oil ratios which in-
crease the amount of coke produced and adversely affect
yields. Increasing catalyst addition rates is costly and
can result in an uneconomical operation.
It has been found that vanadium is especially detrimen-
tal to catalyst life. The vanadium deposited on the cata-
lyst under the reducing conditions in the riser is in an
oxidation state less than +5. At the elevated temperatures
and oxidizing conditions encountered in the regenerator the
vanadium on the catalyst is converted to vanadium oxides, in
particular vanadium pentoxide. The vanadium pentoxide has
a melting point lower than temperatures encountered in the
regeneration zone, and it can become a mobile liquid, flow-
ing across the catalyst surface and plugging pores. This
vanadia may also enter the zeolite structure, neutralizing
the acid sites and, more significantly, irreversibly de-
stroying the crystalline aluminosilicate structure and form-
ing a less active amorphous material. In addition, thismolten vanadia can, at high vanadia levels, especially for
catalyst materials having a low surface area, coat the cata-
lyst microspheres and thereby coalesce particles which ad-
versely affects their fluidization.
Summary of the Invention
In accordance with this invention a process has been
117'~192
provided for converting a vanadium-containing hydrocarbon
oil feed to lighter product~ comprising th~ steps of con-
tacting said oil feed under conversion conditions with a
cracking catalyst to form lighter products and coke, whereby
vanadium in an oxidation state less than +5 is deposited on
said catalyst together with coke. The lighter products are
separated from the spent catalyst and the catalyst is re-
generated by contacting it with an oxygen-containinq gas
under conditions whereby said coke is burned forming CO and
CO2 and said vanadium is maintained in an oxidation state
less than +5.
This invention, by retaining vanadium in an oxidation
state wherein the vanadium has a high melting point, per-
mits the recycle of catalyst to levels of vanadium as high
as 10,000 ppm, or even 20,000 ppm or 50,000 ppm. The ad-
verse effects, such a~ clumping of the catalyst and pore clo-
sings brought about by molten pentavalent vanadium, are thus
avoided. Inasmuch as the catalyst can withstand a much high-
er vanadium loading than previously experienced the amount
of make-up catalyst is reduced.
3rief Description of the Drawinqs
Figs. l and 2 are schematic designs of catalyst regen-
eration and associated cracking apparatus which may be used
in carrying out this invention.
3est and Various Other Modes
For CarrYing Out the Invention
The invention may be carried out by controlling the re-
generatisn of the spent, vanadium-containing catalyst using
several methods, alone or in combination. The objective of
these methods is to retain vanadium in a low oxidation
state, either by not exposing the vanadium to oxidizing con-
ditions, or by exposing vanadium to oxidizing conditions for
3S too short a time to oxidize a significant amount of vanadium
to the +S state.
The concentration of vanadium on the catalyst ?articles
11711~
-- 10 --
increases as the catalyst is recycled, and the vanadium on
the caialyst introduced into the reactor becomes coated with
coke formed in the reactor. In one method of carrying out
the invention, the generator conditions are selected to
ensure that the concentration of coke is retained at at least
a minimum level on the catalyst. This coke may serve either
to ensure a reducing environment for the vandium or to provide
a barrier to the movement of oxidizing gas to underlying
vanadium. The concentration of coke on the catalyst particles
is at least about 0.05 percent and the preferred coke
concentration is at least about 0.15 percent.
In one method of carrying out this invention, which may
be combined with the foregoing method of retaining at least
about 0.05 percent coke on the catalyst or may be used to
achieve lower concentrations of coke, the regeneration is
carried out in an environment which is non-oxidizing for the
vanadium in an oxidation state less than +5. This may be
accomplished by adding reâucing gases such as, for example,
CO or ammonia to the regenerator, or by regenerating under
oxygen-deficient conditions. Oxygen-Aeficient regeneration
- increases the ratio of CO to CO2 and in this method of pro-
viding a non-oxidizing atmosphere the CO/CO2 ratio is at
least about 0.25, preferably is at least about 0.3, and most
preferably is at least about 0.4. The CO/CO2 ratio may be
controlled by controlling the extent of oxygen deficiency
within the regenerator.
The CO/CO2 ratio may be increased by providing
chlorine in an oxidizing atmosphere within the regenerator,
the concentration of chlorine preferably being from about
100 to 400 ppm. This method of increasing the CO/CO2 ratio
is disclosed in Canadian application No. 398,960 filed
1 1'7~ 2
- lOa -
March 23, 1982 for "Addition of MgC12 to Catalyst" and
U.S. Patent No. 4,375,404 for "Addition of Chlorine to
Regenerator", both in the name of George D. Myers.
The use of a reducing atmosphere within the
regenerator is especially useful in combusting coke in
zones where the coke level approaches or is reduced
below about 0.05 per-
11'7~192
--11--
cent, and it is preferred to have a CO/C02 ratio of at leastabout 0.25 in zones where the coke loading is less than
about 0.05 percent by weight.
It is especially useful to keep the vanadium in a re-
duced state under conditions wherein the particles are in
contact or in relatively frequent contact with each other.
Consequently, it is especially contemplated, in carrying out
this method, of maintaining a reducing atmosphere in zones
within the regenerator wherein the catalyst particles are in
a relatively dense bed, such as in a dense fluidized or set-
tled bed. A reducing gas such as CO, methane, or ammonia
may be added to a zone having a dense catalyst phase, such
as for example a bed having a density of about 25 to about
50 pounds per cubic foot.
In another method of carrying out this invention, a ri-
ser regenerator is used as one stage in a multi-stage regen-
erator to contact the catalyst with an oxidizing atmosphere
for a short period of time, such as for example less than
about two seconds and preferably less than about one second.
The riser stage of the regenerator has the advantage in re-
ducing the carbon concentration to a level less than about
0.15 percent or less than about 0.05 percent, that vanadium,
which is no longer protected by a coating of carbon, may not
be in an oxidizing atmosphere for a long enough time to form
molten +S vanadium. Further, the low density of the parti-
cles in the riser-regenerator, minimizes coalescence of
those particles which may have liquid pentavalent vanadia on
their surfaces.
In the preferred method of using a riser regenerator,
the particles are contacted with a reducing atmosphere, such
as one containing CO or other reducing gas, after leaving
the riser. The particles may then be accumulated, as for
example, in a settled bed, before being recycled to contact
additional fresh feed. The catalyst particles to be accu-
mulated are contacted with a reducing atmosphere, ?referablyimmediately after leaving the riser and before acc~mulat1ng
in a dense bed of regenerated particles, and in the ?re-
11~7'~
-l2-
ferred method of carrying out this process the particles are
retained in a reducing atrnosphere within such dense bed, and in
the most preferred method a reducing atmosphere is provided for
the particles until about the time they are contacted with fresh
feed.
The preferred riser regenerator is similar to the vented riser
reactor as is disclosed in U . S . Patents 4, 066, 533 and 4, 070 ,159 to
Myers et al which achieves ballistic separation of gaseous products
from catalyst. This apparatus has the advantages of achieving
virtually instantaneous separation of the regenerated catalyst, now
containing some vanadia to which any oxygen present would have
access, from the oxidizing atmosphere.
In the preferred method of reducing the coke concentration to
a level less than about 0 .15 and especially to less than 0 . 05% the
catalyst is contacted with a reducing atmosphere, preferably
immediately after its separation from the oxidizing atmosphere and
most preferably also in collection zones for the regenerated
catalyst .
This invention may be used in processing any hydrocarbon
feed containing a significant concentration of vanadium, and FCC as
well as RCC processes are contemplated. It is, however, especially
useful in processing reduced crudes having high metal and high
Conradson carbon values, and the invention will be described in
detail with respect to its use in processing an RCC feed.
The carbo-metallic feed comprises or is composed of oil which
boils above about 650F. Such oil, or at least the 650F+ portion
thereof, is characterized by a heavy metal content of at least about
4, preferably more than about 5, and most preferably at least about
5 . 5 ppm of Nickel Equivalents by weight and by a carbon residue
on pyrolysis of at least about 1% and more preferably at least about
2% by weight. In accordance with the invention, the carbo-metallic
feed, in the form of a pumpable liquid, is brought into contact with
hot conversion catalyst in a weight ratio of catalyst to feed in the
RI-6117B
1174192
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ranse of about 3 to about 18 and preferably more than about
6.
The feed in said mixture undergoes a conversion step
which includes cracking while the mixture of feed and
catalyst is flowing through a progressive flow type
reactor. The feed, catalyst, and other materials may be
introduced at one or more points. The reactor includes
an elongated reaction chamber which is at least partly
vertical or inclined and in which the feed material,
resultant products and catalyst are maintained in con-
tact with one another while flowing as a dilute phase or
stream for a predetermined riser residence time in the
range of about 0.5 to about 10 seconds.
The reaction is conducted at a temper~ture of about 900
to about 1400F, measured at the reaction chamber exit,
under a total 2ressure of about 10 to about 50 ?sia
(pounds per square inch absolute) under conditions suffi-
ciently severe to provide a conversion ?er pass in the
range of about S0~ or more and to lay down coke on the
catalyst in an amount in the range of about 0.3 to about
3~ by weight and preferably at least about 0.5g. The
overall rate of coke production, based on weight of fresh
feed, is in the range of about 4 to about 14~ by weight.
At the end of the predetermined residence time, the cata-
lyst is separated from the products, is stripped to remove
high boiling components and other entrained or adsorbed
hydrocarbons and is then regenerated with oxygen-contain-
ing combustion-suprorting gas under condi~ions of time,
temperature and atmosphere sufficient to reduce the carbon
on the regenerated catalyst to about 0.25~ or less.
Depending on how the process of the invention is ?racticed
one or more of the following additional advantages may be
reali~ed. If desired, and preferably, the process may be
1 1'7~
-14-
operated without added hydrogen in the reaction chamber. If
desired, and preferably, the process may be operated without prior
hydrotreating of the feed and/or without other process of removal
of asphaltenes of metals from the feed, and this is true even where
5 the carbo-metallic oil as a whole contains more than about 4, or
more than about 5 or even more than about 5 . 5 ppm Nickel
Equivalents by weight of heavy metal and has a carbon residue on
pyrolysis greater than about 1%, greater than about 1.4% or greater
than aboug 2% by weight. Moreover, all of the converter feed, as
10 above described, may be cracked in one and the same conversion
chamber. The cracking reaction may be carried out with a catalyst
which has previously been used (recycled, except for such
replacement as required to compensate for normal losses and
deactivation) to crack a carbo-metallic feed under the above
15 described conditions. Heavy hydr-ocarbons not cracked to gasoline
in a first pass may be recycled with or without hydrotreating for
further cracking in contact with the same kind of feed in which
they were first subjected to cracking conditions, and under the
same kind of conditions; but operation in a substantially
20 once-through or single pass mode (e . g . less than about 15% by
volume of recycle based on volume of fresh feed) is preferred.
Accordirg to one preferred embodiment or aspect of the
invention, at the end of the predetermined residence time referred
to above, the catalyst is projected in a direction established by the
25 elongated reaction chamber or an extension thereof, while the
products, having lesser momentum, are caused to make an abrupt
change of direction, resulting in an abrupt, substantially
instantaneous ballistic separation of products from catalyst. The
thus separated catalyst is then stripped, regenerated and recycled
30 to the reactor as above described.
According to another preferred embodiment or aspect of the
invention, the converter feed contains 650F+ materi-
RI-6117B
t2
-15-
al which has nvt been hydrotreated and is characterized in part by
containing at least about 5 . 5 parts per million of nickel equivalents
of heavy metals. The converter feed is brought together not only
with the above mentioned cracking catalyst, but also with additional
gaseous material including steam whereby the resultant suspension
of catalyst and feed also includes gaseous material wherein the ratio
of the partial pressure of the added gaseous material relative to
that of the feed is in the range of about 0.25 to about 4Ø The
vapor residence time is in the range of about 0. 5 to about 3
10 seconds when practicing this embodiment or aspect of the invention.
This preferred embodiment or aspect and the one referred to in the
preceeding paragraph may be used in combination with one another
or separately.
According to another preferred embodiment or aspect of the
invention, the carbo-metallic feed is not only brought into contact
with the catalyst, but also with one or more additional materials
including particularly liquid water in a weight ratio raltive to feed
ranging from about 0.04 to about 0.25, more preferably about 0.04
to about O . 2 and still more preferably about O . 05 to about O .15 .
Such additional materials, including the liquid water, may be
brought into admixture with the feed prior to, during or after
mixing the feed with the aforementioned catalyst, and either after
or, preferably, before, vaporization of the feed. The feed,
catalyst and water (e . g . in the form of liquid water or in the form
25 of steam produced by vaporization of liquid water in contact with
the feed) are introduced into the progressive flow type reactor,
which may or may not be a reactor embodying the above described
ballistic separation, at one or more points along the reactor. While
the mixture of feed, catalyst and steam produced by vaporization of
30 the liquid water flows through the reactor, the feed undergoes the
above mentioned conversion step which includes cracking. The feed
material, catalyst, steam and reasultant products
RI -6117B
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are maintained in contact with one another in the above mentioned
elongated reaction chamber while flowing as a dilute phase or stream
for the above mentioned predetermined riser residence time which is
in the range of about 0.5 to about 10 seconds.
The present invention provides a process for the continuous
catalytic conversion of a wide variety of carbo-metallic oils to lower
molecular weight products, while maximizing production of highly
valuable liquid products, and making it possible, if desired, to
avoid vacuum distillation and other expensive treatments such as
hydrotreating. The term "oils", includes not only those
predominantly hydrocarbon compositions which are liquid at room
temperature (i.e., 68F), but also those predominantly hydrocarbon
compositions which are asphalts or tars at ambient temperature but
liquify when heated to temperatures in the range of up to about
800~. The invention is applicable to carbo-metallic oils, whether
of petroleum origin or not. For example, provided they have the
requisite boiling range, carbon residue on pyrolysis and heavy
metals content, the invention may be applied to the processing of
such widely diverse materials as heavy bottoms from crude oil,
heavy bitumen crude oil, those crude oils known as "heavy crude"
which approximate the properties of reduced crude, shale oil, tar
sand extract, products from coal liquification and solvated coal,
atmospheric and vacuum reduced crude, extracts and/or bottoms
(raffinate) from solvent de-asphalting, aromatic extract from lube
oil refining, tar bottoms, heavy cycle oil, slop oil, other refinery
waste streams and mixtures of the foregoing. Such mixtures can
for instance be prepared by mixing available hydrocarbon fractions,
including oils, tars, pitches and the like. Also, powdered coal may
be suspended in the carbo-metallic oil. Persons skilled in the art
are aware of techniques for demetalation of carbo-metallic oils, and
demetalated oils may be converted using
RI-6117B
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the invention; but it is an advantage of the invention that it can
employ as feedstock carbo-metallic oils that have had no prior
demetalation treatment. Likewise, the invention can be applied to
hydrotreated feedstocks; but it is an advantage of the invention
5 that it can successfully convert carbo-metallic oils which have had
substantially no prior hydrotreatment. However, the preferred
application of the process is to red~ced crude, i . e ., that fraction
of crude oil boiling at and above 650F, alone or in admixture with
virgin gas oils. While the use of material that has been subjected
10 to prior vacuum distillation is not excluded, it is an advantage of
the invention that it can satisfactorily process material which has
had no prior vacuum distillation, thus saving on capital investment
and operating costs as compared to conventional FCC processes that
require a vacuum distillation unit.
In accordance with the invention one provides a carbo-metallic
oil feedstock, at least about 70%, more preferably at least about 85%
and still more preferably about 100% (by volume) of which boils at
and above about 650F. All boiling temperatures herein are based
on standard atmospheric pressure conditions. In carbo-metallic oil
20 partly or wholly composed of material which boils at and above
about 650F, such material is referred to herein as 650F+ material;
and 650F+ material which is part of or has been separated from an
oil containing components boiling above and below 650F may be
referred to as a 650F+ fraction. But the terms "boils above" and
25 "650F+" are not intended to imply that all of the material
characterized by said terms will have the capability of boiling. The
carbo-metallic oils contemplated by the invention may contain
material which may not boil under any conditions; for example,
certain asphalts ; and asphaltenes may crack thermally during
30 distillation, apparently without boiling. Thus, for example, when it
is said that the feed comprises at least about 70% by volume of
material which
RI-6117B
1174~92
-18-
boils above about 650F, it should be understood that the
70~ in question may include some material which will not
boil or volatilize at any temperature. These non-boilable ~
materials when present, may frequently or for the most
part be concentrated in portions of the feed which do
not boil below about 1000F, 1025F or higher. Thus,
when it is said that at least about 10~, more preferably
about 15~ and still more preferably at least about 20~
(by volume) of the 650F+ fraction will not boil below
about 1000F or 1025F, it should be understood that
all or any part of the material not boiling below about
1000 or 1025F, may or may not be volatile at and above
the indicated temperatures.
Preferably, the contemplated feeds, or at least the 650F+
material therein, have a carbon residue on ?yrolysis of
at least about 2 or greater. For example, the Conracson
czrbon content may be in the range of about 2 to about 12
and most rre~uently at least about 4. A particularly
common range is about 4 to about 8.
Preferably, the feed has an averaqe composition character-
ized by an atomic hydrogen to carbon ratio in the range of
about 1.2 to about 1.9, and preferably about 1.3 to about
1.8.
The carbo-metallic feeds employed in accordance with the
invention, or at least the 650F+ material therein, may con-
tain at least about 4 parts per million of Nickel Equiva-
lents, as defined above, of which at least about 01. ?pm isvanadium. Carbometallic oils within the above range can be
prepared from mixtures of two or more oils, some of which
do and some of which do not contain the quantities of Nic-
kel Eauivalents and vanadium set forth above. It should
also be noted that the above values for Nickel Equivalents
and nickel represent time-weighted averages for a substan-
tial period of operation of the conversiOn unit, such as one
month, for example. It should also be noted that the heavy
-- 19 --
metals have in certain circumstances exhibited some less-
ening of poisoning tendency after repeated oxidations and
reductions on the catalyst, and the literature describes
criteria for establishing "effective metal" values. For
example, see the article by Cimbalo, et al, entitled
"Deposited Metals Poison FCC Catalyst", Oil and Gas Journal,
May 15, 1972, pp 112-122. If considered necessary or
desirable, the contents of Nickel Equivalents and vanadium
in the carbometallic oils processed according to the
invention may be expressed in terms of "effective metal"
values. Notwithstanding the gradual reduction in poisoning
activity noted by Cimbalo, et al, the regeneration of
catalyst under normal FCC regeneration conditions may not,
and usually does not, severely impair the dehydrogenation,
demathanation and aromatic condensation activity of heavy
metals accumulated on cracking catalyst.
It is known that about 0.2 to about 5 weight per cent of
"sulfur" in the form of elemental sulfur and/or its
compounds ( but reported as elemental sulfur based on the weight
of feed) appears in FCC feeds and that the sulfur and modified
forms of sulfur can find that way into the resultant gasoline
product and, where lead is added, tend to reduce its
susoeptibility to octane enhancement. Sulfur in the product
gasoline often requires sweetening when processing high
sulfur containing crudes. To the extent that sulfur is
present in the coke, it also represents a potential air
pollutant since the regenerator burns it to SO2 and SO3.
i..
d
1 1 '7'~
- l9a -
However, we have found that in our process the sulfur in
the feed is on the other hand able to inhibit heavy metal
activity by maintaining metals such as Ni, V, Cu and Fe in
the sulfide form in the reactor. These sulfides are much
less active than the metals themselves in promoting
dehydrogenation and coking reactions. Accordingly, it is
acceptable to carry out the invention with a carbo-
metallic oil having at least about 0.3~,
117~19;~
_~?0_
acceptably more than about 0.8% and more acceptably at least about
1.5% by weight of sulfur in the 650F+ fraction.
The carbo-metallic oils useful in the invention may and usually
do con tain significant quantities of compounds containing nitrogen,
5 a substantial portion of which may be basic nitrogen. For example,
the total nitrogen content of the carbo-metallic oils may be at least
about 0.05% by weight. Since cracking catalysts owe their cracking
activity to acid sites on the catalyst surface or in its pores, basic
nitrogen-containing compounds may temporarily neutralize these
10 sites, poisoning the catalyst. However, the catalyst is not
permanently damaged since the nitrogen can be burned off the
catalyst during regeneration, as a result of which the acidity of the
active sites is restored.
The carbo-metallic oils may also include significant quantities
15 of pentane insolubles, for example at least about 0 . 5% by weight,
and more typically 2% or more or even about 4% or more. These
may include for instance asphaltenes and other materials.
Alkali and alkaline earth metals generally do not tend to
vaporize in large quantities under the distillation conditions
20 employed in distilling crude oil to prepare the vacuum gas oils
normally used as FCC feedstocks. Rather, these metals remain for
the most part in the "bottomst' fraction (the non-vaporized high
boiling portion) which may for instance be used in the production
of asphalt or other by-products. However, reduced crude and
25 other carbo-metallic oils are in many cases bottoms products, and
therefore may contain significant quantities of alkali and alkaline
earth metals such as sodium. These metals deposit upon the
catalyst during cracking. Depending on the composition of the
catalyst and magnitude of the reger.eration temperatures to which it
30 is exposed, these metals
RI-6117B
11741~Z
may undergo interactions and reactions with the catalyst
(including the catalyst support) which are not normally
experienced in processing VGO under conventional FCC
processing conditions. If the catalyst characteristics
and regeneration conditions so require, one will of course
take the necessary precautions to limit the amounts of
alkali and alkaline earth metal in the feed, which metals
may enter the feed not only as brine associated with the
crude oil in _ts natural state, but also as components of
water or steam which are supplied to the cracking unit.
Thus, careful desalting of the crude used to prepare the
carbo-metallic feed may be important when the catalyst is
particularly susceptible to alkali and alkaline earth metals.
In such circumstances, the content of such metals (herein-
after collectively referred to as ~sodium") in the feed can
be maintained at about 1 ppm or less, based on the weight
of the feedstock. Alternatively, the sodium level of the
feed may be keyed to that of the catalyst, so as to main-
tain the sodium level of the catalyst which is in use sub-
stantially the same as or less than that of the replacement
catalyst which is charged to the unit.
According to a particularly preferred embodiment of the
invention, the carbo-metallic oil feedstock consti'utes
at least about 70~ by volume of material which boils above
about 650F, and at least about 10~ of the material which
boils above about 650F will not boil below about 1025F.
The average composition of this 650F+ material may be fur-
ther characterized by: (a) an atomic hydrogen to carbon ra-
tio in the range of about 1.3 to about 1.8; (b) a Conrad-
son carbon value of at least about 2; (c) at least about
four parts per million of Nickel Equivalents, as defined
above, of which at least about two parts per million is
nickel (as metal, by weiqht), at least about 0.1 part per
million vanadium; and (d) at least one of the following:
(i) at least about 0.3% by weight of sulfur, (ii), at least
about 0.05% by weight of nitrogen, and (iii) at least about
11'~41~2
-22- ,
0.5% by weight of pentane insolubleg. Very comrnonly, the
preferred feed will include all of (i), (ii) and (iii), the
other components found in oil~ of petroleurn and non-petro-
leum origin may also be present in varying quantities pro-
viding they do not prevent operation of the process.
Although there is no intention of excluding the possibilityof usin~ a feedstock which has previously been subjected
to some cracking, the present invention has the definite
advantage that it can successfully product large conver-
sions and very substantial yields of liquid hydrocarbon
fuels from carbo-metallic oils which have not been subject-
ed to any substantial amount of cracking. Thus, for
example, and preferably, at least about 85%, more prefe-ably
at least about 90~ and most preferably substantially all of
of the carbo-metallic feed introluced into the present
process is oil which has not previously been contactad
with cracking catalyst under cracking conditions. ~ore-
over,the process of the invention is suitabla for oparation
in a substantially once-through or singla pass mode. Thus,
the volume of recycle, if any, based on the volume of fresh
feed is prefarably about 15% or less and more preferably
about 10~ or less.
In general, the weight ratio of catalyst to fresh feed
(feed which has not previously been exposed to cracking
catalyst under carcking conditions) used in the process
is in the range of about 3 to about 18. Preferred and
more preferred ratios are about 4 to about 12, more prefer-
ably about S to about 10 and still more preferably about6 to about 10, a ratio of about 10 presentlv being con-
sidered most nearly o?timum. ~ithin the limitations of
product quality reauirements, controlling the catalyst to
oil ~atio at relatively low levels within the afo-esaid
rarges tends to reduce the coke yield of the ?rocess, based
on 'resh feed.
117~19;~
-23-
In conventional FCC processing of VGO, the ratio between
the nu~ber of barrels per day of plant through-put and
the total number of tons of catalyst undergoing circula-
tion throughout all phases of the process can vary widely.
For purposes of this disclosure, daily plant through-put
is defined as the number of barrels of fresh feed boiling
above about 650F which that plant processes per average
day of operation to lisuit products boiling below about
430F. For example, in one commercially successful type
of FCC-VG0 operation, about 8 to about 12 tons of catalyst
are under circulation in the process ?er 1000 barrels
per day of plant through-put. In another commercially
successful process, this ratio is in the range of about 2
to 3. While the present invention may be practiced in the
range of about 2 to about 30 and more typically about 2 to
about 12 tons of catalyst inventory per 1000 barrels of
daily plant through-put, it is preferred to carry out the
process of the present invention with a very small ratio
of catalyst weight to daily plant through-put. ~ore
specifically, it is preferred to carry out the proc~ss of
the present invention with an inventory of catalyst that
is sufficient to contact the feed for the desired residence
time in the above indicated catalyst to oil ratio while
minimizing the amount of catalyst inventory, relative to
plant through-put, which is undergoing circulation or
being held for treatment in other phases of the process
such as, for example, stripping, regeneration and the li~e.
Thus, more particularly, it is preferred to carry out the
process of the present invention with about 2 to about 5
and more preferably about 2 tons of catalyst inventory or
less per thousand barrels of daily plant through-put.
In the practice of the invention, catalyst may be added
continuously or periodically, such as, for e~ample, to
make up for normal losses of catal st from the system.
.~loreover, catalyst addition may be conduc_-d in con~_nc-
tion with withdrawal of catalyst, such as, Cor e~ ?le, -o
m~intain or increase the average activity level of ~he
li7'1
~f
catalyst in the unit. ~or exannple, the rate at which virgin catalyst
is adAed to the unit may be in the range of about 0.1 to about 3,
more preferably about 0.15 to about 2, and most preferably to
about 0.2 to about 1.5 pounds per barrel of feed. If on the other
hand equilibrium catalyst from FCC operation is to be utilized,
replacement rates as high as about 5 pound per barrel can be
practiced. Where circumstances are such that the catalyst employed
in the unit is below average in resistance to deactivation and/or
conditions prevailing in the unit are such as to promote more rapid
10 deactivation, one may employ rates of addition greater than those
stated above; but in the opposite circumstances, lower rates of
addi-tion may be employed. By way of illustration, if a unit were
operated with a metal(s) loading of 5000 ppm Ni + V in parts by
weight on equilibrium catalyst, one might for example employ a
replacement rate of about 2 . 7 pounds of catalyst introduced for
each barrel (42 gallons) of feed processed. However, operation at
a higher level such as 10,000 ppm Ni + V on catalyst would enable
one to substantially reduce the replacement rate, such as for
example to about 1.3 pounds of catalyst per barrel of feed. Thus,
20 the levels of metal(s) on catalyst and catalyst replacement rates may
in general be respectively increased and decreased to any value
consistent with the catalyst activity which is available and desired
for conducting the process.
Without wishing to be bound by any theory, it appears that a
25 number of features of the process to be described in greater detail
below, such as, for instance, the residence time and optional mixing
of steam with the feedstock, tend to restrict the extent to which
cracking conditions produce metals in the reduced state on the
catalyst from heavy metal sulfide(s), sulfate(s) or oxide(s)
30 deposited on the catalyst particles by prior exposures to
carbo-metallic feedstocks and regeneration conditions. Thus, the
RI-6117B
process appears to afford significant control over the poisoning
effect of heavy metals on the catalyst even when the accu-
mulations of such metals are quite substantial.
Accordingly, the process may be practiced with catalyst5 bearing high accumulations of heavy metal(s) in the form of
elemental metal(s), oxide(s), sulfide(s) or other compounds. Thus,
operation of the process with catalyst bearing heavy metals
accumulations in the range of about 3000 or more ppm Nickel
Equivalents, on the average, is contemplated. The concentration of
10 Nickel Equivalents of metals on catalyst can range up to about
50,000 ppm or higher. More specifically, the accumulation may be
in the range of about 3000 to about 30,000 ppm, preferably in the
range of about 3000 to 20,000 ppm, and more particularly about
3000 to about 12,000 ppm . Within these ranyes just mentioned,
operation at metals levels of about 4000 or more, about 5000 or
more, or about 7000 or more ppm can tend to reduce the rate of
catalyst replacement required. The foregoing ranges are based on
parts per million of Nickel Equivalents, in which the metals are
expressed as metal, by weight, measured on and based on
regenerated equilibrium catalyst. However, in the event that
catalyst of adequate activity is available at very low cost, making
feasible very high rated of catalyst replacement, the carbo-metallic
oil could be converted to lower boiling liquid products with catalyst
bearing less than 3,000 ppm Nickel Equivalents of heavy metals.
For example, one might employ equilibrium catalyst from another
unit, for example, an FCC unit which has been used in -the
cracking of a feed, e. g. vacuum gas oil, having a carbon residue
on pyr olysis of less than 1 and containing less than about 4 ppm
Nickel Equivalents of heavy metals.
. 30 In any event, the quilibrium concentration of heavy metals in
the circulating inventory of catalyst can be controlled (including
maintained or varied as desired or needed ) by manipulation of the
rate of catalyst addition discussed above. Thus, for example,
RI-6117B
li7~1'32
addition of catalyst may ~e maintained at a rate which will control
the heavy metals
RI-6117B
117~1~Z
accumulation on the catalyst in one of the ranaes set
forth above.
In seneral, it is preferred to employ a catalyst having a
relatively high level of cracking activity, proviaing high
levels of conversion and productivity at low residence
times. The conversion capabilities of the catalvst mav
be expressed in terms of the conversion produced ~uring ac-
tual operation of the process and/or in terms of conversion
produced in standard catalyst activity tests. For exa~ple,
it is preferred to em~loy catalyst which, in the course of
extended operation under prcvailing process conditions, is
sufficientlv active for sustaining a levcl of conversion of
at least about 50~O and more preferably at least about 60~.
In this connection, conversion is expressed in liauid vol-
ume percent, based on fresh ~eed.
Also, for example, the preferred catalyst may be definedas one which, in its virgin or eouilibrium state, e~hibits
a s?ecified activity expressed as a percentage in terms
of ,~AT (micro-activity test) conversion. For pur~oses of
the present invention the foregoing percentage is the vol-
ume percentage of standard feedstock which a catalyst uncer
evaluation will convert to 430F end point gasoline, light-
er ?roducts and coke at 900F, 16 WHSV (weight hourly space
velocity, calculated on a moisture free basis, usinc clean
catalyst which has been dried at 1100F, weighed and thenconditioned, for a period of at least 8 hours at about 28C
and ;Og relative humidity, until about one hour or less
prior to contacting the feed) and 3C/O (catalyst to oil
weiqht ratio) by AST~ ~-3Z MAT test D-390~-aO, using an a?-
propriate standard feedstock, e.g. a sweet light ?rimarygas oil, such as that used bv Davison, Division of W.R.
Grace, having the following analvsis and ?ro?e-ties:
3S API Gravity at 60F, degrees 31~0
S?eci'ic Gr~vity at 50F, ~/cc 0.~-~08
-27-
Ramsbottom Carbon, wt . % 0.09
Conradson Carbon, wt. % (est. ) 0.04
Carbon, wt . % 84.92
Hydrogen, wt . % 12.94
Sulfur, wt . % 0.68
Nitrogen, ppm 305
Viscosity at 100F, centistokes 10.36
Watson K Factor 11.93
Aniline Point 182
Bromine No. 2.2
Paraffins, Vol . % 31.7
Olefins, Vol . % 1.6
Naphthenes, Vol . % 44.0
Aromatics, Vol . % 22.7
Average Molecular Weight 284
Nickel Trace
Vanadium Trace
Iron Trace
Sodium Trace
Chlorides Trace
BS&W Trace
Distillation ASTM D-1160
IBP 445
10% 601
30% 664
50% 701
70% 734
90% 787
FBP 834
The gasoline end point and boiling temperature-volume percent
relationships of the product produced in the MAT conversion test
may for example be determined by simulated distillation techniques,
for example modifications of gas chromate graphic "Sim-D", ASTM
D-2887-73. The results of such simulations are in reasonable
agreement with the results obtained hy subjecting larger samples of
material to standard laboratory distillation techniques.
RI -6117B
-2~s-
Conversion is calculated by subtracting from 100 the volume percent
(based on fresh feed) of those products heavier than gasoline which
remain in the recovered product.
On page 935-937 of Hougen and Watson, Chemical Process
Principles, John Wiley & Sons, Inc ., N . Y . (1947), the concent of
"Activity Factors " is discussed . This concept leads to the use of
"relative activity" to compare the effectiveness of an operating
catalyst against a standard catalyst. Relative activity measurements
facilitate recognition of how the quantitiy requirements of various
catalysts differ from one another. Thus, relative activity is a ratio
obtained by dividing the weight of a standard or reference catalyst
which is or would be required to produce a given level of
conversion, as compared to the weight of an operating catalyst
(whether proposed or actually used) which is or would be required
to produce the same level of conversion in the same or equivalent
feedstock under the same or equivalent conditions. Said ratio of
catalyst weights may be expressed as a numerical ratio, but
preferably is converted to a percentage basis. The standard
catalyst is preferably chosen from among catalysts useful for
conducting the present invention, such as for example zeolite fluid
cracking catalysts, and is chosen for its ability to produce a
predetermined level of conversion in a standard feed under the
conditions of temperature, WHSV, catalyst to oil ratio and other
conditions set forth in the preceding description of the MAT
conversion test and in ASTM D-32 MAT test D-39û7-80. Conversion
is the volume percentage of feedstock that is converted to 430F
end point gasoline, lighter producs and coke. ~or standard feed,
one may employ the above-mentioned light primary gas oil, or
equivalent.
For purposes of conducting relative activity determinations,
one may prepare a "standard catalyst curve", a chart or
RI-6117B
li'7~ 2
-29 -
graph of conversion (as above defined) vs. reciprocal WHSV for the
standard catalyst and feedstock. A sufficient number of runs is
made under ASTM D-3907-80 conditions (as modified above) using
standard feedstock at varying levels of WHSV to prepare an
5 accurated 'tcurve" of conversion vs. WHSV for the standard
feedstock. This curve should traverse all or substantially all of
the various levels of conversion including the range of conversion
within which it is expected that the operating catalyst will be
tested. From this curve, one may establish a standard WHSV for
10 test comparisons and a standard value of reciprocal WHSV
corresponding to that level of conversion which has been chosen to
represent 100% relative activity in the standard catalyst. For
purposes of the present disclosure the aforementioned reciprocal
WHSV and level of conversion are, respectively, 0.0625 and 75%.
15 In testing an operating catalyst of unknown relative activity, one
conducts a sufficient number of runs with that catalyst under
D-3907-80 conditions (as modified above) to establish the level of
conversion which is or would be produced with the operating
catalyst at standard recirpocal WHSV. Then, using the
20 above-mentioned standard catalyst curve, one establishes a
hypothetical reciprocal WHSV constituting the reciprocal WHSV which
would have been required, using the standard catalyst, to obtain
the same level of conversion which was or would be exhibited, by
the operating catalyst at standard WHSV. The relative activity may
25 then be calculated by dividing the hypothetical reciprocal WHSV by
the reciporcal standard WHSV, which is 1/16, or .0625. The result
is reltaive activity expressed in terms of a decimal fraction, which
may then be multiplied by 100 to convert to percent relative
activity. In applying the results of this determination, a relative
activity of 0.5, ro 50%, means that it would take twice the amount
of the operating catalyst to give the same conversion as the
standard catalyst, i . e ., the production catalyst is 50% as active as
the reference catalyst.
RI-6117B
~174192
The catalyst may be introduced into the process in its virgin
form or, as previously indicated, in other than virgin form; e. g .
one may use equilibrium catalyst withdrawn from another unit, such
as catalyst that has been employed in the cracking of a different
5 feed. Whether characterized on the basis of MAT conversion
activity or relative activity, the preferred catalysts may be
described on the basis of their activity "as introduced" into the
process of the present invention, or on the basis of their "as
withdrawn" or equilibrium activity in the process of the present
10 invention, or on both of these bases. A preferred activity level of
virgin and non-virgin catalyst "as introduced" into the process of
the present invention is at least about 60% by MAT conversion, and
preferably at least about 20%, more preferably at least about 40%
and still more preferably at least about 60% in terms of relative
15 activity. However, it will be appreciated that, particularly in the
case of non-virgin catalysts supplied at high addition rates, lower
activity levels may be acceptable. An acceptable "as withdrawn" or
equilibrium activity level of catalyst which has been used in the
process of the present invention is at least about 20% or more, but
20 about 40% or more and preferably about 60% or more are preferred
values on a relative actcivity basis, and an activity level of 60% or
more on a MAT conversion basis is also contemplated. More
preferably, it is desired to employ a catalyst which will, under the
conditions of use in the unit, establish an equilibrium activity at or
25 about the indicated level. The catalyst activities are determined
with catalyst having less than 0.01 coke, e.g. regenerated catalyst.
One may employ any hydrocarbon cracking catalyst having the
above indicated conversion capabilities. A particularly preferred
class of catalysts includes those which has pore structures into
30 which molecules of feed material may enter for adsorption and/or for
contact with active cataly-
RI-6117B ~ 3~ _
117419Z
-31-
tic sites within or adjacent the pores. Various types of
catalysts are available withi~ this classification, includ-
ing for exa~ple the layered silicates, e.g. smectites.
Although the ~ost widely available catalysts within this
classification are the well-known zeolite-containing
catalysts, non-zeolite catalysts are also contemplated.
The preferred zeolite-containing catalysts may include
any zeolite, whetl~er natural, semi-synthetic or synthetic,
alone or in admixture with other materials which do not
significantly impair the suitability of the catalyst, ?ro-
vided the resultant catalyst has the activity and pore
structure referred to above. For example, if the virgin
catalyst is a mixture, it may include the zeolite compo-
nent associated with or dispersed in a porous refr~ctory
inorganic oxide carrier, in such case the catalyst may for
e~ample cont~in about 1~ to a~out 60~, more ~referably
about 15 to about 50%, and most typically about 20 to about
45~ by weight, based on the total weisht of catalyst (water
free basis) of the zeolite, the balance of the catalyst
being the porous refractory inorganic oxide alone or in
combination with any of the known adjuvants for promoting
or suppressing various desired and undesired reactions.
For a general explanation of the genus of zeolite, mole-
cular sieve catalysts useful in the invention, attention
is drawn to the disclosures of the articles entitled
"Refinery Catalysts Are a Fluid ~usiness" and "Making Cat
Crackers Work On Varied Diet~, appearing respectively in
the July 26, 1978 and September 13, 1978 issues of Chemical
Week magazine.
For the most part, the zeolite components of the zeol: e-
containing catalysts will be those which are known to be
use'ul in FCC crac~ing processes. In general, these
are crystalline aluminosilicates, ty?ically ~ade U2 of
tetra coordinated aluminum atoms associated throush o~ysen
atoms with adjacent silicon atoms in the crystal structure.
117~1~tZ
-32-
However, the term "zeolite" as used in this disclosure
contemplates not only aluminosilicates, but also substances
in which the aluminum has been partly or wholly replaced,
such as for instance bv gallium and/or other metal atoms,
and further includes substances in which all or part of
the silicon has been replaced, such as for instance by
ge~æ-.ium. Titanium and zirconium substitution may also
be practiced
Most zeolites are prepared or occur naturally in the sodium
form, so that sodium cations are associated with the elec:ro-
negative sites in the crystal structure. The sodium
cations tend to m~e zeolites inactive and much less
stable when exposed to hydrocarbon conversion conditions,
particularly high temperatures. ~ccordingly, the zeolite
may be ion exchanged, and where the zeolite is a component
of a catalyst composition, such ion exchanging may occur
before or after incorporation of the zeolite as a component
of the composition. Suitable cations for replacement
of sodium in the zeolite crystal structure include a~onium
(decomposable to hydrogen~, hydrogen, rare earth metals,
alkaline earth metals, etc. Various suitable ion exchange
procedures and cations which may be exchanged into the
zeolite c ystal structure are well known to those skilled
ir. the art.
Examples of the naturally occuring crystalline
aluminosilicate zeolites which may be used as or included
in the catalyst for the present invention are faujasite,
mordenite, clinoptilote, chabazite, analcite, crionite,
as well as levynite, dachiardite, paulingite, noselite,
ferriorite, heulandite, scolccite, stibite, harmotome,
phillipsite, brewsterite, flarite, datolite, gmelinite,
ca~mnite, leucite, lazurite, s_721ite, mesolite, ?toL te,
ne?hline, matrolite, offretlte and sodalite.
Exa ples of the synth2tic crys_all;ne a um nosll:_ te
1 1~7~1~t2
zeolites which are useful as or in the catalyst for carry-
ing out the present invention are Zeolite X, U. S. Patent
No. 2,882,244, Zeolite Y, U. S. Patent No. 3,130,007:
and Zeolite A, U. S. Patent No. 2,882,243; as well as
Zeolite B, U. S. Patent No. 3,008,803; Zeolite D, Canada
Patent No. 661,981; Zeolite E, Canada Patent No. 614,495;
Zeolite F, V. S. Patent No. 2,996,358; Zeolite ~. U. S.
Patent No. 3,010,789; Zeolite J., U. S. Patent No.
3,011,869; Zeolite J., Belgian Patent No. 575,177; Zeolite
M., U. S. Patent No. 2,995,423, Zeolite O, U. S. Patent
No. 3,140,252; Zeolite Q, U. S. Patent No. 2,991,151;
Zeolite S, U. S. Patent No. 3,054,657, Zeolite T, U. S.
Patent No. 2,950,952; Zeolite W, U. S. Patent No.
3,012,853; Zeolite Z, Canada Patent No. 614,495; and
Zeolite Omega, Canada Patent No. 817,915. Also, ZK-4HJ,
alpha beta and ZSM-type zeolites are useful. Moreover,
the zeolites described in U. S. Patents Nos. 3,140,249,
3,140,253, 3,944,482 and 4,137,151 are also useful,
The crystaliine aluminosilicate zeolites having a fauja-
site-type crvstal structure are particularly preferred for
use in the present invention. This includes particularly
natural faujasite and Zeolite X and Zeolite Y.
The crystalline aluminosilicate zeolites, such as synthetic
faujasite, will under normal conditions crystallize as
regularly shaped, discrete particles of about one to about
ten microns in size, and, accordingly, this is the size
range frequently found in commercial catalysts which can
be used in the invention. Preferably, the particle size of
the zeolites is from about 0.1 to about 10 mic_ons and more
prefe-ably is from about 0.1 to about 2 mic-ons or less.
For example, zeolites prepared in situ from calcined ~aolin
may ~e characterized by even smaller crystallites. Crvs_al-
line zeolites exhibit both an interior and an exte-ior
11'7'~1~tZ
-3~-
surface area, which we have defined as "portal" surface area, with
the largest portion of the total surface area being internal. By
portal surface area, we refer to the outer surface of the zeolite
crystal through which reactants are considered to pass in order to
convert to lower boiling products. Blockages of the internal
channels by, for example, coke formation, blockages of entrance to
the internal channels by deposition of coke in the portal surface
area, the contamination by metals poisoning, will greatly reduce the
total zeolite surface area. Therefore, to minimize the effect of
contamination and pore blockage, crystals larger than the normal
size cited above are preferably not used in the catalysts of this
invention .
Commercial zeolite-containing catalysts are available with
carriers containing a variety of metal oxides and combination
thereof, including for example silica, alumina, magnesia, and
mixtures thereof and mixtures of such oxides with clays as e . g .
described in U. S. Patent No. 3,034,948. One may for example
select any of the zeolite-containing molecular sieve fluid cracking
catalysts which are suitable for production of gasoline from vacuum
gas oils. However, certain advantages may be attained by judicious
selection of catalysts having marked resistance to metals. A metal
resistant zeolite catalyst is, for instance, described in U.S. Patent
No. 3,944,482, in which the catalyst contains 1-40 weight percent of
a rare earth-exchanged zeolite, the balance being a refractory metal
oxide having specified pore volume and size distribution. Other
catalysts described as "metals-tolerant" are described in the above
mentioned Cimbalo et al article.
In general, it is preferred to employ catalysts having an
over-all particle size in the range of about 5 to about 16U, more
preferably about 40 to about 120, and most preferably about 40 to
about 80 microns. For example, a useful catalyst may have a
skeletal density of about 150
RI-6117~
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pounds per cubic foot and an averaae particle size of about
60-70 micronS, with less than 10~ of the particles havinq
a size less th n ~bout 40 microns and less than 80~ having
a size less than about 50-60 microns.
Although a wide variety of other catalysts, including
both zeolite-containing and non-zeolite-containing mqy be
employed in the pr~ctice of the invention the following
are e~am21es of commercially available catalysts which may
be emoloyed in practicing the invention: '
TABLE 2
Specific Weiqht Percent
Surface Zeolite
m2/Content A12O3 SiO2 Na2O Fe2O TiO2
AGZ-290 30011.0 29.5 59.0 0.40 0.11 0.59
GRZ-l 16214.0 23.4 69.0 0.10 0.4 0.9
CCZ-220 12911.0 34.6 60.0 0.60 0.57 1.9
Super DX lSS 13.0 31.0 65.0 0.80 0.57 1.6
F-87 24010.0 44.0 50.0 0.80 0.70 1.6
FOX-90 2408.0 44.0 52.0 0.65 0.65 1.1
HFZ 20 31020.0 59.0 40.0 0.47 0.54 2.75
HEZ SS 21019.0 59.0 35.2 0.60 0.60 2.5
The AGZ-290, GRZ-l, CCZ-220 and Su2er DX catalysts re'erred
to above are products of W. R. Grace and Co. F-87 and
FOC-90 are product~ of Filtrol, while HFZ-20 and HEZ-SS
are 2roducts of Engelhard/Houdry~ The above are ?roperties
of virgin catalyst and, except in the case of zeolite con-
tent, are adjusted to a water free basis, i.e. based on
material ignited at 1750F. The zeolite content is derived
by comparison of the X-ray intensities of a catalyst sam21e
and of a standard material com?osed of high ?ur-ty sodlum
Y zeolite in accordance with draft ~6, dated January 9,
1978, of proposed AST~5 Stand3rd ~5ethod er.ti~led "De~ na-
3j 'ion o' the raujasite Content of a Cat~lvst."
-36-
Among the above mentioned commercially available catalyst, the
Super D family and especially a catalyst designated GRZ-1 are
particularly preferred. For example, Super DX has given
particularly good results with Arabian Light crude. The GRZ-1,
5 although substantially more expensive than the Super DX at
present, appears somewhat more metals tolerant.
Although not yet commercially available, it is believed that the
best catalysts for carrying out the present invention will be those
which, according to proposals advanced by Dr. William P.
10 Hettinger, jr. and Dr. James E. Lewis, are characterized by
matrices with feeder pores having large minimum diameters and
large mouths to facilitate diffusion of high molecular wieght
molecules through the matrix to the portal surface area of molecular
sieve particles within the matrix. Such matrices preferably also
15 have a relatively large pore volume in order to soak up unvaporized
- portions of the carbo-metallic oil feed. Thus, significant numbers
of liquid hydrocarbon molecules can diffuse to active catalytic sites
both in the matrix and in sieve particles on the surface of the
matrix. In general it is preferred to employ catalysts with matrices
20 wherein the feeder pores have diameters in the range of about 400
to about 6000 angstrom units, and preferably about 1000 to about
6000 angstrom units.
It is considered to be an advantage that the process of the
present invention can be conducted in the substantial absence of tin
25 and/or anitmony or at least in the presence of a catalyst which is
substantially free of either or both of these metals.
The process of the present invention may be operated with the
above described carbo-metallic oil and catalyst as substantially the
sole materials char~ed to the reaction zone. But the charging of
30 additional material is not excluded.
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The charging of recycled oil to tne reaction zone has already
been mentioned. As described in greater detail below, still
other materials fulfilling a variety of functions may also
be charged. In such case, the carbo-metallic oil and
catalyst usually represent the major proportion by weight
of the total of all materials charged to the reaction zone.
Certain of the additional materials which may be used perform
functions which offer significant advantages over the process
as performed with only the carbo-metallic oil and catalyst.
Among these functions are: controlling the effects of heavy
metals and other catalyst contaminants; enhancing catalyst
activity; absorbing excess heat in the catalyst as received
from the regenerator; disposal of pollutants or conversion
thereof to a form or forms in which they may be more readily
separated from products and/or disposed of; controlling catalyst
temperature; diluting the carbo-metallic oil vapors to reduce
their partial pressure and increase the yield of desired
products; adjusting feed/catalyst contact time; donation of
hydrogen to a hydrogen deficient carbo-metallic oil feedstock,
for example as disclosed in Canadian application 398,945
entitled "Use of Naphtha in Carbo-Metallic Oil Conversion" and
filed March 22, 1982, assisting in the dispersion of the
feed; and possibly also distillation of products. Certain
of the metals in the heavy metals accumulation on the
catalyst are more active in promoting undesired reactions
when they are the form of elemental metal, than they are
when in the oxidized form produced by contact with oxygen
in the catalyst regenerator. ~owever, the time of
ll';~ ~lS~Z
- 37a -
contact between catalyst and vapors of feed and product in
past conventional catalytic cracking was sufficient so that
hydrogen released in the cracking reaction was able to
reconvert a signifcant portion of the less harmful oxides
back to the more harmful elemental heavy metals. One can
take advantage of this situation through the introduction of
additional materials
Z
-3~-
which are in gaseoue (including vaporous) form in the reaction zone
in admixture with the catalyst and vapors of feed and products.
The increased volume of material in the reaction zone resultin~ from
the presence of such additional materials tends to increase the
5 velocity of flow through the reaction zone with a corresponding
decrease in the residence time of the catalyst and oxidized heavy
metals borne thereby. Because of this reduced residence time,
there is less opportunity for reduction of the oxidized heavy metals
to elemental form and therefore less of the harmful elemental
10 metals are available for contacting the feed and products.
Added materials may be introduced into the process in any
suitable fashion, some examples of which follow. For instance, they
may be admixed with the carbo-metallic oil feedstock piror to
contact of the latter with the catalyst. Alternatively, the added
15 materials may, if desired, be admixed with the catalyst prior to
contact of the latter with the feedstock. Separate portions of the
added materials may be separately admixed with both catalyst and
carbo-metallic oil. Moreover, the feedstock, catalyst and additional
materials may, if desired, by brought together substantially
20 simultaneously. A portion of the added material smay be mixed with
catalyst and/or carbo-metallic oil in any of the above described
ways, while additional portions are subsequently brought into
admixture. For example, a portion of the added materials may be
added to the carbo-metallic oil and/or to the catalyst before they
25 reach the reaction zone, while another portion of the added
materials is introduced directly into the reaction zone. The added
materials may be introduced at a plurality of spaced locations in the
reaction zone or along the length thereof, if elongated.
The amount of additional materials which may be present in the
30 feed, catalyst or reaction zone for carrying out the
RI-6117B
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above functions, and others, may be varied as desired; but said
amount will preferably be sufficient to substantially heat balance the
process. These materials may for example be introduced into the
reaction zone in a weight ratio relative to feed of up to about 0.4,
5 preferably in the range of about 0.02 to about 0.4, more preferably
about 0 . 03 to about 0 . 3 and most preferably about 0 . 05 to about
0.25.
For example, many or all of the abcve desirable functions may
be attained by introducing H20 to the reaction zone in the form of
10 steam or of liquid water or a combination thereof in a weight ratio
relative to feed in the range of about 0 . 04 or more, or more
preferably about 0.05 to about 0.1 or more. Without wishing to be
bound by any theory, it appears that the use of H20 tends to
inhibit reduction of catalyst-borne oxides, sulfites and sulfides to
15 the free metallic form which is believed to promote
condensation-dehydrogenation with consequent promotion of coke
and hydrogen yield and accompanying loss of product. Moreover,
H20 may also, to some extent, reduce deposition of metals onto the
catalyst surface. There may also be some tendency to desorb
20 nitrogen-containing and other heavy contaminant-containing
molecules from the surface of the catalyst particles, or at least some
tendency to inhibit their absorption by the catalyst. It is also
believed that added H20 tends to increase the acidity of the
catalyst by Bronsted acid formation which in turn enhances the
25 activity of the catalyst. Assuming the H20 as supplied is cooler
than the regenerated catalyst and/or the temperature of the
reaction zone, the sensible heat involved in raising the temperature
of the H20 upon contacting the catalyst in the reaction zone or
elsewhere can absorb excess heat from the catalyst. Where the H20
30 is or includes recycled water that contains for example about 500 to
about 5000 ppm of H2S dissolved therein, a number of additional
advantages may accrue. The ecologically unattractive H2S need not
RI-6117B
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be vented to the atompshere, the recycled water does not require
further treatment to remove H2S and the H2S may be of assistance
in reducing coking of the catalyst by passivation of the heavy
metals, i . e . by conversion thereof to the sulfide form which has a
5 lesser tendency than the free metals to enhance coke and hydrogen
production. In the reaction zone, the presence of H20 can dilute
the carbo-metallic oil vapors, thus reducing their partial pressure
and tending to increase the yield of the desired products. It
has been reported that H20 is useful in combination with other
10 materials in generating hydrogen during cracking; thus it may be
able to act as a hydrogen donor for hydrogen deficient
carbo-metallic oil feedstocks. The H20 may also serve certain
purely mechanical functions such as: assisting in the atomizing or
dispersion of the feed; competing with high molecular weight
15 molecules for adsorption on the surface of the catalyst, thus
interrupting coke formation; steam distillation of vaporizable product
from unvaporized feed material; and disengagement of product from
catalyst upon conclusion of the cracking reaction. It is particularly
preferred to bring together H20, catalyst and carbo-metallic oil
20 substantially simultaneously~ For example, one may admix ~I20 and
feedstock in an atomizing nozzle and immediately direct the resultant
spray into contact with the catalyst at the downstream end of the
reaction zone.
The addition of steam to the reaction zone is frequently
25 mentioned in the literature of fluid catalytic cracking. ~ddition of
liquid water to the feed is discussed relatively infrequently,
compared to the introduction of steam directly into the reaction
zone. However, in accordance with the present invention it is
particularly preferred that liquid water be brought into intimate
30 admixture with the carbo-metallic oil in a weight ratio of about 0.04
to about 0 . 25 at or prior to the time of in troduction of the oil into
the reaction zone, whereby the water (e . g ., in the form of liquid
water or in the form of steam produced by
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vaporization of liquid water in contact with the oil) enters the
reaction zone as part of the flow of feedstock which enters such
zone. Although not wishing to be bound by any theory, it is
believed that the foregoing is advantageous in promoting dispersion
5 of the feedstock. Also, the heat of vaporization of the water,
which heat is absorbed from the catalyst, from the feedstock, or
from both, causes the water to be a more efficient heat sink than
steam alone. Preferably the weight ratio of liquid water to feeds to
about O . 04 to about O . 2 more preferably about O . 05 to about O .15 .
Of course, the liquid water may be introduced into the process
in the above described manner or in other ways, and in either
event the introduction of liquid water may be accompanied by the
introduction of additional amounts of water as steam into the same
of different portions of the reaction zone or into the catalyst
15 and/or feedstock. For example, the amount of additional steam may
be in a weight ratio relative to feed in the range of about 0.01 to
about 0.25, with the weight ratio of total H20 (as steam and liquid
water) to feedstock being about O . 3 or less . The charging weight
ratio of liquid water relative to steam in such combined use of
20 liquid water and steam may for example range from about 15 which
is preesntly preferred, to about 0.2. Such ratio may be maintained
at a predetermined level within such range of varied as necessary
or desired to adjust or maintain heat balance.
Other materials may be added to the reaction zone to perform
25 one or more of the above described functions. For example, the
dehydrogenation-condensation activity of heavy metals may be
inhibited by introducing hydrogen sulfide gas into the reaction
zone. Hydrogen may be made available for hydrogen deficient
carbo-metallic oil feedstocks by introducing into the reaction zone
30 either a conventional hydrogen donor diluent such as a heavy
naphtha relative-
F~I-6117B
1 1'7'~
-42 -
ly low molecular weight carbon-hydrogen fragment contributors,
including for example: light paraffins; low molecular weight
alcohols and other compounds which permit or favor intermolecular
hydrogen transfer; and compounds that chemically combine to
5 generate hydrogen in the reaction zone such as by reaction of
carbon monoxide with water, or with alcohols, or with olefins, or
with other materials or mixtures of the foregoing.
All of the above mentioned additional materials (including
water), alone or in conjunction with each other or in conjunction
10 with other materials, such as nitrogen or other inert gases, light
hydrocarbons, and other, may perform any of the above-described
functions for which they are suitable, including without limitation,
acting as diluents to reduce feed partial pressure and/or as heat
sinks to absorb excess heat present in the catalyst as received from
15 the regeneration step. The foregoing is a discussion of some of the
functions which can be performed by materials other than catalyst
and carbo-metallic oil feedstock introduced into the reaction zone,
and it should be understood that other materials may be added or
other functions performed without departing from the spirit of the
20 invention.
The invention may be practiced in a wide variety of apparatus.
However, the preferred apparatus includes means for rapidly
vaporizing as much feed as possibie and efficiently admixing feed
and catalyst (although not necessarily in that order), for causing
25 the resultant mixture to flow as a dilute suspension in a
progressive flow mode, and for separating the catalyst from cracked
products and any uncracked or only partially cracked feed at the
end of a predetermined residence time or times, it being preferred
that all or at least a substantial portion of the product should be
30 abruptly separated from at least a portion of the catalyst.
For example, the apparatus may include, along its elongated
RI-6117B
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~17 ~1~2
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reaction chamber, one or more points ror introduction of
carbo-metallic feed, one or more points for introduction
of catalyst, one or more points for introduction of
additional material, one or more points for withdrawal of
products and one or more points of withdrawal of catalyst.
The means for introducir.g feed, catalyst and other material
may range from open pipes to sophisticated jets or spray
nozzles, it being preferred to use means capable of breaking
up the liquid feed into fine droplets. Preferably, the
catalyst, liquid water twhen used) and fresh feed are
brought together in an apparatus similar to that disclosed
in Canadian Application No. 341,829, filed January 11, 1982.
According to a particularly preferred embodiment, the liquid
water and carbo-metallic oil, prior to their introduction into
the riser, are caused to pass through a propeller, apertured
disc, or any appropriate high shear agitating means for
forming a "homogenized mixture" containing finely divided
droplets of oil and/or water with oil and/or water present as
a continuous phase.
It is preferred that the reaction chamber, or at least the
major portion thereof, be more nearly vertical than
horizontal and have a length to diameter ratio of at least
about 10, more preferably about 20 or 25 or more. Use of
a vertical riser type reactor is preferred. If tubular,
the reactor can be of uniform diameter throughout or may
be provided with a continuous or step-wise increase in
diameter along the reaction path to maintain or vary the
velocity along the flow path.
In general, the charging means (for catalyst and feed)
and the reactor configuration are such as to provide a
117~
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relatively high velocity of flow and dilute suspension of catalyst.
For example, the vapor or catalyst velocity in the riser will be
usually at least about 25 and more typically at least about 35 feet
per second. This velocity may range up to about 55 or about 75
feet or about 100 feet per second or higher. The vapor velocity at
the top of the reactor may be higher than that at the bottom and
may for example be about 80 feet per second at the top and about
40 feet per second at the bottom. The velocity capabilities of the
reactor will in general be sufficient to prevent substantial build-up
of catalyst bed in the bottom or other portions of the riser,
whereby the catalyst loading in the riser can be maintained below
about 4 or 5 pounds, as for example about 0.5 pounds, and below
about 2 pounds, as for example 0 . 8 pound, per cubic foot,
respectively, at the upstream (e . g . bottom) and downstream (e . g .
top) ends of the riser.
The progressive flow mode involves, for example flowing of
catalyst, feed and products as a stream in a positively controlled
and maintained direction established by the elongated nature of the
reaction zone. This is not to suggest however that there must be
strictly linear flow. As is well known, turbulent flow and
"slippage" of catalyst may occur to some extent especially in certain
ranges of vapor velocity and some catalyst loadings, although it has
been reported advisable to employ sufficiently low catalyst loadings
to re~trict slippage and back-mixing.
Most preferably the reactor is one which abruptly separates a
substantial portion or all of the vaporized cracked products from
the catalyst at one or more points along the riser, and preferably
separates substantially all of the vaporized cracked products from
the catalyst at the downstream end of the riser. A preferred type
of reactor embodies ballistic separation of catalyst and products;
RI-6117B
11'7~ 2
-45 -
that is, catalyst is projected in a direction established by the risertube, and is caused to continue its motion in the general direction
so established, while the products, having lesser momentum, are
caused to make an abrupt change of direction, resulting in an
5 abrupt, substantially instantaneous separation of product from
catalyst. In a preferred embodiment referred to as a vented riser,
the riser tube is provided with a substantially unobstructed
discharge opening at i~s downstream end for discharge of catalyst.
An exit port in the side of the tube adjacent the downstream end
10 receives the products. The discharge opening communicates with a
catalyst flow path which extends to the usual stripper and
regenerator, while the exit port communicates with a product flow
path which is substantially or entirely separated from the catalyst
flow path and leads to separation means for separating the products
15 from the relatively small portion of catalyst, if any, which manages
to gain entry to the product exit port. Examples of a ballistic
separation apparatus and technique as above described, are found
in U . S . Patents 4, 066, 533 and 4, 070 ,159 to Myers et al the
disclosures of which patents are hereby incorporated herein by
20 reference in their entireties. According to the particularly
preferred embodiment, based on a suggestion understood to have
emanated from Paul W. Walters, Roger M. Benslay and Dwight F.
Barger, the ballistic separation step includes at least a partial
reversal of direction by the product vapors upon discharge from
25 the riser tube; that is, the product vapors make a turn or change
of direction which exceeds 90 at the riser tube outlet. This may
be accomplished for example by providing a cup-like member
surrounding the riser tube at its upper end, the ratio of
cross-sectional area of the cup-like member relative to the
30 cross-sectional area of the riser tube outlet being low i.e. Iess than
1 and preferably less than about 0.6. Preferably the lip of the cup
is slightly downstream of, or above the downstream end or top of
the riser tube, and the cup is preferably concentric with the riser
RI-6117B
il7~19~
-4t,-
tube. By means of a product vapor line communicating with the
interior of the cup but not the interior of the riser tube, having
its inlet positioned within the cup interior in a direction upstream
of the riser tube outlet, product vapors emanating from the riser
tube and entering the cup by reversal of direction are transported
away from the cup to catalyst and product separation equipment.
Such an arrangement can produce a high degree of completion of
the separation of catalyst from product separation equipment. Such
an arrangement can produce a high degree of completion of the
separation of catalyst from product vapors at the riser tube outlet,
so that the required amount of auxiliary catalyst separation
equipment such as cyclones is greatly reduced, with consequent
large savings in capital investment and operating cost.
Preferred conditions for operation of the process are described
below. Among these are feed, catalyst and reaction temperatures,
reaction and feed pressures, residence time and levels of
conversion, coke production and coke laydown on catalyst.
In conventional FCC operations with VGO, the feedstock is
customarily preheated, often to temperatures significantly higher
than are required to make the feed sufficiently fluid for pumping
and or introduction into the reactor. For example, preheat
temperatures as high as about 700 or 800F have been reported.
But in our process as presently practiced it is preferred to restrict
preheating of the feed, so that the feed is capable of absorbing a
larger amount of heat from the catalyst while the catalyst raises the
feed to conversion temperature, at the same time minimizing
utilization of external fuels to heat the feedstock. Thus, where the
nature of the feedstock permits, it may be fed at ambient
temperature. Heavier stocks may be fed at preheat temperatures of
up to about 600F, typically about 200F to about 500F, but higher
preheat temperatures are not necessarily exluded.
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-47-
The catalyst fed to the reactor may vary widely in
temperature, for example from about 1100 to about 1600F, more
preferably about 1200 to about 1500F and most preferably about
1300 to about 1400F, with about 1325 to about 1375 being
considered optimum at present.
As indicated previously, the conversion of the carbo-metallic
oil to lower molecular weight products may be conducted at a
temperature of about 900 to about 1400F, measured at the reaction
chamber outlet. The reaction temperature as measured at said
outlet is more preferably maintained in the range of about 965 to
about 1300F, still more preferably about 975 to about 1200F, and
most preferably about 980 to about 1150F. Depending upon the
temperature selected and the properties of the feed, all of the feed
may or may not vaporize in the riser.
Although the pressure in the reactor may, as indicated above,
range from about 10 to about 50 psia, preferred and more preferred
pressure ranges are about 15 to about 35 and about 20 to about 35.
In general, the partial (or total) pressure of the feed may be in
the range of about 3 to about 30, more preferably about 7 to about
25 and most preferably about 10 to about 17 psia. The feed partial
pressure may be controlled or suppressed by the introduction of
gaseous (including vaporous) materials into the reactor, such as for
instance the steam, water and other additional materials described
above. The process has for example been operated with the ratio
of feed partial pressure relative to total pressure in the riser in
the range of about 0.2 to about 0.8, more typically about 0.3 to
abou tO .7 and still more typically about 0.4 to about 0.6, with the
ratio of added gaseous material (which may include recyclecl gases
and/or steam resulting from introduction of H20 to the riser in the
form of steam and/or liquid water) relative to total pressure in the
riser correspondingly ranging from about
RI-6117B
11'7'~1~2
-~i~-
0 . 8 to about 0 . 2, more typically about 0 . 7 to about 0 . 3 and still
more typically about 0.6 to about 0.4. In the illustrative operations
just described, the ratio of the partial pressure of the added
gaseous material relative to the partial pressure of the feed has
been in the range of about 0.25 to about 4.0, more typically about
0 . 4 to abou-t 2. 3 and still more typically about 0 . 7 to about 1. 7 .
Although the residence time of feed and product vapors in the
riser may be in the range of about 0 . 5 to about 10 seconds, as
described above, preferred and more preferred values are about 0.5
to about 6 and about 1 to about 4 seconds, with about 1.5 to about
3 . 0 seconds currently being considered about optimum . ~or
example, the process has been operated with a riser vapor
residence time of about 2.5 seconds or less by introduction of
copious amounts of gaseous materials into the riser, such amounts
being sufficient to provide for example a partial pressure ratio of
added gaseous materials relative to hydrocarbon feed of about 0 . 8
or more. By way of further illustration, the process has been
operated with said residence time being about two seconds or less,
with the aforesaid ratio being in the range of about 1 to about 2.
The combination of low feed partial pressure, very low residence
time and ballistic separation of products from catalyst are
considered especially beneficial for the conversion of carbo-metallic
oils. Additional benefits may be obtained in the foregoing
combination when there is a substantial partial pressure of added
gaseous material, especially H20 as described above.
Depending upon whether there is slippage between the catalyst
and hydrocarbon vapors in the riser, the catalyst riser residence
time may or may not be the same as
RI-6117B
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~9
that of the vapors. Thus, the ratio of average catalyst reactor
residence time versus vapor reactor residence time, i . e . slippage,
may be in the range of about 1 to about 5, more preferably about 1
to about 4 and most preferably about 1 to about 3, with about 1 to
5 about 2 currently being considered optimum.
In practice, there will usually be a small amount of slippage,
e . g ., at least about 1. 05 or 1. 2 . In an operating unit there may
for example be a slippage of about 1.1 at the bottom of the riser
and about 1.05 at the top.
In certain types of known FCC units, there is a riser which
discharges catalyst and product vapors together into an enlarged
chamber, usually considered to be part of the reactor, in which the
catalyst is disengaged from product and collected. Continued
contact of catalyst uncracked feed (if any) and cracked products in
15 such enlarged chamber results in an overall catalyst feed contact
time appreciably exceeding the riser tube residence times of the
vapors and catalysts. When practicing the process of the present
invention with ballistic separation of catalyst and vapors at the
downstream (e.g. upper) extremity of the riser, such as is taught
20 in the above mentioned Myers et al patents, the riser residence time
and the catalyst contact time are substantially the same for a major
portion of the feed and product vapors. It is considered
advantageous if the vapor riser residence time and vapor catalyst
con~act time are substantially the same for at least about ~0%, more
25 preferably at least about 90% and most preferably at least about 95%
by volume of the total feed and product vapors passing through the
riser. By denying such vapors continued contact with catalyst in a
catalyst disengagement and collection chamber one may avoid a
tendency toward re-cracking and diminished selectivity.
RI-6117B
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In general, the combination of catalyst to oil ratio,
temperatures, pressures and residence times should be such as to
effect a substantial conversion of the carbo-metallic oil feedstock.
It is an advantage of the process that very high levels of
conversion can be attained in a single pass; for example the
conversion may be in excess of 50% and may range to about ~0% or
higher. Preferably, the aforementioned conditions are maintained at
levels sufficient to maintain conversion levels in the range of about
60 to about 90% and more preferably about 70 to about 85%. The
10 foregoing conversion levels are calculated by subtracting from 100~o
the percentage obtained by dividing the liquid volume of fresh feed
into 100 times the volume of liquid product boiling at and above
430F (tbp, standard atmospheric pressure).
These substantial levels of conversion may and usually do
15 result in relatively large yields of coke, such as for example about
4 to about 14% by weight based on fresh feed, more commonly about
6 to about 13% and most frequently about 10 to about 13%. The
coke yield can more or less quantitatively deposit upon the catalyst.
At cor.templated catalyst to oil ratios, the resultant coke laydown
20 may be in excess of about 0 . 3, more commonly in excess of about
0 . 5 and very frequently in excess of about 1% of coke by weight,
based on the weight of moisture free regeneration catalyst. Such
coke laydown may range as high about 2%, or about 3%, or even
higher .
In common with conventional PCC operations on VGO, the
present process includes stripping of spent catalyst after
disengagement of the catalyst from product vapors. Persons skilled
in the art are acquainted with appropriate stripping agents and
conditions for stripping spent catalyst, but in some cases the
30 present process may require somewhat more severe conditions than
are commonly employed. This may result, for example, from the
use of
RI -6117B
117'11~1~
-51 -
a carbo-metallic oil having constituents which do not volatilize under
the conditions prevailing in the reactor, which constituents deposit
themselves at least in part on the catalyst. Such adsorbed,
unvaporized material can be troublesome from at least two
5 standpoints. First, if the gases (including vapors) used to strip
the catalyst can gain admission to a catalyst disengagement or
collection chamber connected to the downstream end of the riser,
and if there is an accumulation of catalyst in such chamber,
vaporization of these unvaporized hydrocarbons in the stripper can
10 be followed by adsorption on the bed of catalyst in the chamber.
More particularly, as the catalyst in the stripper is stripped of
adsorbed feed material, the resultant feed material vapors pass
through the bed of catalyst accumulated in the catalyst collection
and/or disengagement chamber and may deposit coke and/or
15 condensed material on the catalyst in said bed. As the catalyst
bearing such deposits moves from the bed and into the stripper and
from thence to the regenerator, the condensed products can create
a demand for more stripping capacity, while the coke can tend to
increase regeneration temperatures and/or demand greater
20 regeneration capacity. For the foregoing reasons, it is preferred
to prevent or restrict contact between stripping vapors and catalyst
accumulations in the catalyst disengagement of collection chamber.
This may be done for example by preventing such accumulations
may be done for example by preventing such accumulations from
25 forming, e. g. with the exception of a quantity of catalyst which
essentially drops out of circulation and may remain at the bottom of
the disengagement and/or collection chamber, the catalyst that is in
circulation may be removed from said chamber promptly upon
settling to the bottom of the chamber. Also, to minimize
30 regeneration temperatures and demand for regeneration capacity, it
may be desirable to employ conditions of time, temperature and
atmosphere in the stripper which are sufficient to reduce potentially
volatile hydrocarbon material borne by
RI-6117B
~7~19f~
the stripped catalyst to about 10~i or less by weight of the total
carbon loading on the catalyst. Such stripping may for example
include reheating of the catalyst, extensive stripping with steam,
the use of gases having a temperature considered higher than
5 normal for FCC/VGO opertaions, such as for instance flue gas from
the regenerator, as well as other refinery stream gases such as
hydrotreater off-gas (H2S containing), hydrogen and others. For
example, the stripper may be operated at a temperature of about
350F using steam at a pressure of about 150 psig and a weight
10 ratio of steam to catalyst of about 0.002 to about 0.003. On the
other hand, the stripper may be operated at a temperature of about
1025F or higher.
Substantial conversion of carbo-metallic oils to lighter products
in accordance with the invention tends to produce sufficiently large
15 coke yields and coke laydown on catalyst to require some care in
catalyst regeneration. In order to maintain adequate activity in
zeolite and non-zeolite catalysts, it is desirable to regenerate the
catalyst under conditions of time, temperature and atmosphere
sufficient to reduce the percent by weight of carbon remaining on
20 the catalyst to about 0.25% or less . The amounts of coke which
must therefore by burned off of the catalysts when processing
carbo-metallic oils are usually substantially greater than would be
the case when cracking VGO. The term coke when used to
describe the present invention, should be understood to include any
25 residual unvaporized feed or cracking product, if any such material
is present on the catalyst after stripping.
Regeneration of catalyst, burning away of coke deposited on
the catalyst during the conversion of the feed, may be performed at
any suitable temperature in the range of about 1100 to about
30 160ûF, measured at the regenerator catalyst outlet. This
temperature is preferably in the range of about 1200F to about
1500F, more preferably about 1275 to about 1425F and optimally
about 1325 to about 1375F. The process has been operated, for
RI - 6117B
:;
li7'~192
.,
--~) 3--
example, with a fluidized regenerator with the temperature of the
catalyst dense phase in the range of about 130û to about 1400F.
In accordance with the invention, regeneration is conducted
while maintaining the catalyst in one or more fluidized beds in one
5 or more fluidization chambers. Such fluidized bed operations are
characterized, for instance, by one or more fluidized dense beds of
ebulliating particles having a bed density of, for example, about 25
to about 50 pounds per cubic foot. Fluidization is maintained by
passing gases, including combustion supporting gases, through the
10 bed at a sufficient velocity to maintain the particles in a fluidized
state but at a velocity which is sufficiently small to prevent
substantial entrainment of particles in the gases. For example, the
lineal velocity of the fluidizing gases may be in the range of about
0.2 to about 4 feet per second and preferably about 0.2 to about 3
15 feet per second. The average total residence time of the particles
in the one or more beds is substantial, ranging for example from
about 5 to about 30, more preferably about 5 to about 20 and still
more preferably about 5 to about 10 minutes.
Heat released by combustion of coke in the regenerator is
20 absorbed by the catalyst and can be readily retained thereby until
the regenerated catalyst is brought into contact with fresh feed.
When processing carbo-metallic oils to the relatively high levels of
conversion involved in the present invention, the amount of
regenerator heat which is transmitted to fresh feed by way of
25 recycling regenerated catalyst can substantially exceed the level of
heat input which is appropriate in the riser for heating the
vaporizing the feed and other materials, for supplying the
endothermic heat of reaction for cracking, for making up the
RI-6117B
-54 -
heat losses of the unit and so forth. Thus, the amount of
regenerator heat transmitted to fresh feed may he controlled, or
restricted where necessary, within certain approximate ranges.
The amount of heat so transmitted may for example be in the range
of about 500 to about 1200, more particularly about 600 to about
900, and more particularly about 650 to about 850 BTUs per pound
of fresh feed. The aforesaid ranges refer to the combined heat, in
BTUs per pound of fresh feed, which is transmitted by the catalyst
to the feed and reaction products (between the contacting of feed
with catalyst and the separation of product from catalyst) for
supplying the heat of reaction (e . g . for cracking) and the
difference in enthalpy between the products and the fresh feed.
Not included in the foregoing are the heat made available in the
reactor by the adsorption of coke on the catalyst, nor the heat
consumed by heating, vaporizing or reacting recycle streams and
such added materials as water, steam naphtha and other hydrogen
donors, flue gases and inert gases, or by radiation and other
losses .
One or a combination of techniques may be utilized for
controlling or restricting the amount of regeneration heat
transmitted via catalyst to fresh feed. For example, one may add a
combustion modifier to the cracking catalyst in order to reduce the
temperature of combustion of coke to carbon dioxide and/or carbon
monoxide in the regenerator. Moreover, one may remove heat from
the catalyst through heat exchange means, including for example
heat exchangers (e.g. steam coils) built into the regenerator itself,
whereby one may extract heat from the catalyst during
regeneration. Heat exchangers can be built into catalyst transfer
lines, such as for instance the catalyst return line from the
regenerator to the reactor, whereby heat may be removed from the
catalyst after it is regenerated. The amount of heat imparted to
the catalyst in the regenerator may be restricted by reducing the
amount
RI-6117B
li7419;~
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of insulation on the regenerator to permit some heat loss to the
surrounding atmosphere, especially if feeds of exceedingly high
coking potential are planned for processing; in general, such loss
of heat to the atmosphere is considered economically less desirable
5 than certain of the other alternatives set forth herein. One may
also inject cooling fluids into portions of the regenerator other than
those occupied by the dense bed, for example water and/or steam,
whereby the amount of inert gas available in the regenerator for
heat absorption and removal is increased.
Another suitable and preferred technique for controlling or
restricting the heat transmitted to fresh feed via recycled
regenerated catalyst involves maintaining a specified ratio between
the carbon dioxide and carbon monoxide formed in the regenerator
while such gases are in heat exchange contact or relationship with
15 catalyst undergoing regeneration.
Still another particularly preferred technique for controlling or
restricting the regeneration heat imparted to fresh feed via recycled
catalyst involves the diversion of a portion of the heat borne by
recycled catalyst to added materials introduced into the reactor,
20 such as the water, steam, naphtha, other hydrogen donors, flue
gases, inert gases, and other gaseous or vaporizable materials
which may be introduced into the reactor.
In most circulstances, it will be important to insure that no
adsorbed oxygen containing gases are carried into the riser by
25 recycled catalyst. Thus, whenever such action is considered
necessary, the catalyst discharged from the regenerator may be
stripped with appropriate stripping gases to remove oxygen
containing gases. Such stripping may for instance be conducted at
relatively high temperatures, for example about 1350 to about
30 1370F, using steam, nitrogen or other inert gas as the stripping
gas(es). The use of nitrogen
RI-6117B
i 1'7~ 2
-56-
and other inert gases is beneficial from the standpoint of avoiding a
tendency toward hydrothermal catalyst deactivation which may result
from the use of steam.
The following comments and discussion relating to metals
5 management, carbon management and heat management may be of
assistance in obtaining best results when operating the invention.
Since these remarks are for the most part directed to what is
considered the best mode of operation, it should be apparent that
the invention is not limited to the particular modes of operation
10 discussed below. Moreover, since certain of these comments are
necessarily based on theoretical considerations, there is no intention
to be bound by any such theory, whether expressed herein or
implicit in the operating suggestions set forth hereinafter.
Although discussed separately below, it is readily apparent
15 that metals management, carbon management and heat management
are interrelated and interdependent subjects both in theory and
practice. While coke yield and coke laydown on catalyst are
primarily the result of the relatively large quantities of coke
precursors found in carbo-metallic oils, the production of coke is
20 exacerbated by high metals accumulations, which can also
significantly affect catalyst performance. Moreover, the degree of
success experienced in metals management and carbon management
will have a direct influence on the extent to which heat management
is necessary. Moreover, some of the steps taken in support of
25 metals management have proved very helpful in respect to carbon
and heat management.
As noted previously the presence of a large heavy metals
accumulation on the catalyst tends to aggravate the problem of
dehydrogenation and aromatic condensation, resulting in increased
30 production of gases and coke for a feedstock of a given Ramsbottom
carbon value. The introduction of substantial quantities of H20
into the reactor, either in the form
RI-6117B
,~,
11 7L1 19Z
-57-
of steam or liquid water, appears highly beneficial from the
standpoint of keeping the heavy metals in a less harmful form, i . e .
the oxide rather than metallic form. This is of assistance in
maintaining the desired selectivity.
Also, a unit design in which system components and residence
times are selected to reduce the ratio of catalyst reactor residence
time relative to catalyst regenerator residence time will tend to
reduce the ratio of the times during which the catalyst is
respectively under reduction conditions and oxidation conditions.
This too can assist in maintaining desired levels of selectivity.
Whether the metals content of the catalyst is being managed
successfully may be observed by monitoring the total hydrogen plus
methane produced in the reactor and/or the ratio of hydrogen to
methane thus produced. In general, it is considered that the
hydrogen to methane mole ratio should be less than about 1 and
preferably about 0 . 6 or less, with about 0 . 4 or less being
considered about optimum. In actual practice the hydrogen to
methane ratio may range from about 0 . 5 to about 1. 5 and average
about 0.8 to about 1.
Careful carbon management can improve both selectivity (the
ability to maximize production of valuable products ), and heat
productivity. In general, the techniques of metals control
described above are also of assistance in carbon management. The
usefulness of water addition in respect to carbon management has
already been spelled out in considerable detail in that part of the
specification which relates to added materials for introduction into
the reaction zone. In general, those techniques which improve
dispersion of the feed in the reaction zone should also prove
helpful, these include for instance the use of fogging or misting
devices to assist in dispersing the feed.
Catalyst to oil ratio is also a factor in heat management. In
common with prior FCC practice on VGO, the reactor temperature
RI-6117B
;
11'~'1192
-58-
may be controlled in the practice of the present invention by
respectively increasing or decreasing the flow of hot regenerated
catalyst to the reactor in response to decreases and increases in
reactor temperature, typically the outlet temperature in the case of
5 a riser type reactor. Where the automatic controller for catalyst
introduction is set to maintain an excessive catalyst to oil ratio, one
can expect unnecessarily large rates of carbon production and heat
release, relative to the weight of fresh feed charged to the reaction
zone .
Relatively high reactor temperatures are also beneficial from
the standpoint of carbon management. Such higher temperatures
foster more complete vaporization of feed and disengagement of
product from catalyst.
Carbon management can also be facilitated by suitable
15 restriction of the total pressure in the reactor and the partial
pressure of the feed. In general, at a given level of conversion,
relatively small decreases in the aforementioned pressures can
substantially reduce coke production. This may be due to the fact
that restricting total pressure tends to enhance vaporization of high
20 boiling components of the feed, encourage cracking and facilitate
disengagement of both unconverted feed and higher boiling cracked
products from the catalyst. It may be of assistance in this regard
to restrict the pressure drop of equipment downstream of and in
communication with the reactor. But if it is desired or necessary
25 to operate the system at higher total pressure, such as for instance
because of operating limitations (e. g. pressure drop in downstream
equipment) the above descrihed benefits may be obtained by
restricting the feed partial pressure. Suitable ranges for total
reactor pressure and feed partial pressure have been set forth
30 above, and in general it is desirable to attempt to minimize the
pressures within these ranges.
The abrupt separation of catalyst from product vapors and
RI-6117B
i
11'7~:19Z
-59-
unconverted feed (if any) is also of great assistance. It is for this
reason that the so-called vented riser apparatus and technique
disclosed in U . S . Patents 4, 070 ,159 and 4, 066, 533 to George D .
Myers et al is the preferred type of apparatus for conducting this
5 process. For similar reasons, it is beneficial to reduce insofar as
possible the elapsed time between separation of catalyst from
product vapors and the commencement of stripping. The vented
riser and prompt stripping tend to reduce the opportunity for
coking of unconverted feed and higher boiling cracked products
10 adsorbed on the catalyst.
A particularly desirable mode of operation from the standpoint
of carbon management is to operate the process in the vented riser
using a hydrogen donor if necessary, while maintaining the feed
partial pressure and total reactor pressure as low as possible, and
15 incorporating relatively large amounts of water, steam and if
desired, other diluents, which provide the numerous benefits
discussed in greater detail above. Moreover, when liquid water,
steam, hydrogen donors, hydrogen and other gaseous or
vaporizable materials are fed to the reaction zone, the feeding of
20 these materials provides an opportunity for exercising additional
control over catalyst to oil ratio. Thus, for example, the practice
of increasing or decreasing the catalyst to oil ratio for a given
amount of decrease or increase in reactor temperature may be
reduced or eliminated by substituting either appropriate reduction
25 or increase in the charging ratios of the water, steam and other
gaseous or vaporizable material, or an appropriate reduction or
increase in the ratio of water to steam and/or other gaseous
materials introduced into the reaction zone.
Heat management includes measures taken to control the amount
30 of heat released in various parts of the process and/or for dealing
successfully with such heat as may be released. Unlike
conventional FCC practice using VGO, wherein it is usually a
RI-6117B
1 1'7'~
-60-
problem to generate sufficient heat during regeneration to heat
balance the reactor, the processing of carbo-metallic oils generally
produces so much heat as to require careful management thereof.
Heat management can be facilitated by various techniques
5 associated with the materials introduced into the reactor. Thus,
heat absorption by feed can be maximized by minimum preheating of
feed, it being necessary only that the feed temperature be high
enough so that it is sufficiently fluid for sucessful pumping and
dispersion in the reactor. When the catalyst is maintained in a
10 highly active state with the suppression of coking (metals control),
so as to achieve higher conversion, the resultant higher conversion
and greater selectivity can increase the heat absorption of the
reaction. In general, higher reactor temperatures promote catalyst
conversion activity in the face of more refractory and higher boiling
15 constituents with high coking potentials. While the rate of catalyst
deactivation may thus be increased, the higher tcmperature of
operation tends to offset this loss in activity. Higher temperatures
in the reactor also contribute to enhancement of octane number,
thus offsetting the octane depressant effect of high carbon lay
20 down. Other techniques for absorbing heat have also been
discussed above in connection with the introduction of water,
steam, and other gaseous or vaporizable materials into the reactor.
As noted above, the invention can be practiced in the
above-described mode and in many others. an illustrative,
25 non-limiting example is described by the accompanying schematic
diagrams in the figures and by the description of these figures
which follows.
Referring in detail to the drawings, in Figure 1 petroleum
feedstock is introduced into the lower end of riser reactor 2
30 through inlet line 1 at which point it is mixed with hot regenerated
catalyst coming from regenerator 9 through line 3.
~'.` Rl-6117B
li~7~1~?2
-~>:1 -
The feedstock is catalytically cracked in passing up riser 2
and the product vapors are separated from spent catalyst in vessel
8. The catalyst particles move upwardly from riser 2 into the
space within vessel 8 and fall downwardly into dense bed 16. The
S cracking products together with some catalyst fines pass through
horizontal line 4 into cyclone 5. The gases are separated from the
catalyst and pass out through line 6. The catalyst fines drop into
bed 16 through dipleg 19.
The spent catalyst, coated with coke and vanadium in a
10 reduced state, passes through line 7 into upper dense fluidized bed
18 within regenerator 9. The spent catalyst is fluidized with a
mixture of air, C0 and C02 passing through porous plate 21 from
lower zone 20. The spent catalyst is partially regenerated in bed
18 and is passed into the lower portion of vented riser 13 through
line 11. Air is introduced into riser 13 through line 12 where it is
mixed with partially regenerated catalyst. The catalyst is forced
rapidly upwards through the riser and it falls into dense settled
bed 17. Line 14 provides a source of reducing gas such as C0 for
bed 17 to keep the regenerated catalyst in a reducing atmosphere
20 and thus keep vanadium present in a reduced oxidation state.
Regenerated catalyst is returned to the riser reactor 2 through
line 3, which is provided with a source of a reducing gas such as
C0 through line 22.
In Figure 2, spent catalyst coated with coke and vanadium in a
25 reduced state flows into dense fluidized bed 32 of regenerator 31
through inlet line 33. Air to combust the coke and fluidize the
catalyst is introduced through line 34 into air distributor 35. Coke
is burned and passes upwardly into riser regenerator 36. The
partially regenerated catalyst which reaches the riser 35 is
30 contacted with air from line 37 which completes the regeneration.
The regenerated catalyst passes upwardly from the top of the riser
36 and falls down into dense settled bed 37. Dense bed 37 and the
zone above 37 through which the regenerated catalyst falls are
RI-6117B
1~7'~
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supplied with a reducing gas such as CO through lines 40 and 41.
The regenerated catalyst is returned to the cracking reactor
through lin~ 38. The CO-rich flue gases leave the regenerator
through line 39.
5Having thus described this invention, the following Example is
offered to illustrate it in more detail.
Example
A carbo-metallic feed at a temperature of about 400F is fed at
a rate of about 2000 pounds per hour into the bottom of a vented
10riser reactor where it is mixed with a zeolite catalyst at a
temperature of about 1275~F and a catalyst to oil ratio by weight of
about 11.
The carbo-metallic feed has a heavy metal content of about 5
ppm Nickel Equivalents, including 3 ppm vanadium, and has a
15Conradson carbon content of about 7 percent. About 86 percent of
the feed boils above 650F and about 20 percent of the feed boils
above 1025F.
The temperature within the reactor is about 1000F and the
pressure is about 27 psia. About 75 percent of the feed is
20converted to fractions boiling at a temperature less than 430F and
about 53 percent of the feed is converted to gasoline. During the
conversion, about 11 percent of the feed is converted to coke.
The catalyst containing about one percent by weight of coke
contains about 20,000 ppm Nickel Equivalents including about 12,000
25ppm vanadium. The catalyst is stripped with steam at a
temperature of about 1000F to remove volatiles and the stripped
catalyst is introduced into the upper zone of the regenerator as
shown in Figure 1 at a rate of about 23,000 pounds per hour, and
is partially regenerated to a coke concentration of about 0.2 percent
30by a mixture of air, CO and C02. The CO/C02 ratio in the
fluidized bed in the upper zone is about 0.3.
The partially regenerated catalyst is passed to the bottom of a
riser reactor where it is contacted with air in
RI-6117B
1 ~ '7'~1~2
an amount sufficient to force the catalyst up the riser with a
residence time of about. 1 second. The regenerated catalyst, having
a coke loading of about 0.05 percent exits from the top of the riser
and falls into a dense bed having a reducing atmosphere comprising
5 CO. The regenerated catalyst is recycled to the riser reactor for
contact with additional feed.
RI-6117B
