U.S. patent number 3,849,291 [Application Number 05/186,639] was granted by the patent office on 1974-11-19 for high temperature catalytic cracking with low coke producing crystalline zeolite catalysts.
This patent grant is currently assigned to Mobil Oil Corporation. Invention is credited to Hartley Owen.
United States Patent |
3,849,291 |
Owen |
November 19, 1974 |
HIGH TEMPERATURE CATALYTIC CRACKING WITH LOW COKE PRODUCING
CRYSTALLINE ZEOLITE CATALYSTS
Abstract
A method and arrangement of catalyst handling steps is described
for practicing a selective high temperature catalytic cracking of
hydrocarbons which will take advantage of the low coke producing
catalyst of high activity and selectivity.
Inventors: |
Owen; Hartley (Belle Mead,
NJ) |
Assignee: |
Mobil Oil Corporation (New
York, NY)
|
Family
ID: |
22685724 |
Appl.
No.: |
05/186,639 |
Filed: |
October 5, 1971 |
Current U.S.
Class: |
208/78;
208/DIG.2; 208/67; 208/164; 208/120.01 |
Current CPC
Class: |
B01J
8/26 (20130101); B01J 29/90 (20130101); C10G
11/18 (20130101); B01J 29/40 (20130101); B01J
2229/40 (20130101); Y02P 30/446 (20151101); Y10S
208/02 (20130101); Y02P 30/40 (20151101) |
Current International
Class: |
C10G
11/18 (20060101); C10G 11/00 (20060101); B01J
8/26 (20060101); B01J 8/24 (20060101); B01J
29/40 (20060101); B01J 29/90 (20060101); B01J
29/00 (20060101); C10g 037/02 (); C10g 011/18 ();
B01j 009/20 () |
Field of
Search: |
;208/78,80,120,155,164 |
References Cited
[Referenced By]
U.S. Patent Documents
Other References
Voorhies "Carbon Formation in Catalytic Cracking," Ind. Eng. Chem.
37, 318-322, (1945). .
Shankland and Schmitkons, ""Determination of Activity and
Selectivity of Cracking Catalyst," Proc. API 27, (III), 1947, pp.
57-77..
|
Primary Examiner: Gantz; Delbert E.
Assistant Examiner: Schmitkons; G. E.
Attorney, Agent or Firm: Gaboriault; Andrew L. Farnsworth;
Carl D.
Claims
I claim:
1. In a dual riser hydrocarbon conversion operation employing
fluidizable catalyst particles, the improved method of operation
which comprises,
regenerating a catalyst mixture comprising a large pore crystalline
zeolite of the faujasite type in admixture with from 10 to about 90
weight percent of a smaller pore crystalline zeolite of the ZSM-5
type under conditions to heat the catalyst mixture to an elevated
temperature in the range of 1,000.degree.F. up to about
1400.degree.F.,
passing heated catalyst to the inlet of a first riser conversion
zone for admixture with gaseous materials such as dry gas and/or
C.sub.5 and lighter gaseous hydrocarbons to form a first suspension
therewith,
passing said first suspension in admixture with a fresh gas oil
feed upwardly through said first conversion zone under elevated
temperature cracking conversion conditions of at least
1000.degree.F. and a hydrocarbon residence time less than about 12
seconds to cyclonic separation of the suspension into a hydrocarbon
phase and a low coke containing catalyst phase, recovering the
hydrocarbon phase and stripping the catalyst phase,
passing stripped catalyst mixture of relatively low coke level
separated from said first conversion zone to the inlet of a second
riser conversion zone in combination with an increment of the
freshly regenerated catalyst,
forming a second suspension of catalyst with gaseous materials such
as dry gas and/or C.sub.5 and lighter gaseous hydrocarbons for
passage upwardly through said second riser conversion zone,
passing said second suspension with higher coke producing
hydrocarbon feed upwardly through said second conversion zone at a
temperature of at least 950.degree.F. for a hydrocarbon residence
time in the range of 2 to 15 seconds and a higher catalyst to oil
ratio than employed in said first conversion zone,
cyclonically separating the second suspension at the discharge of
the second conversion zone into a second hydrocarbon phase and a
second catalyst phase,
recovering the second hydrocarbon phase and stripping the second
catalyst phase of entrained hydrocarbons, and
passing stripped second phase catalyst to said catalyst
regeneration.
2. The process of claim 1 wherein an amount of metal oxide which
will enhance regeneration of the catalyst by converting carbon
monoxide to carbon dioxide is admixed with the zeolite catalyst
mixture and particularly the ZSM-5 crystalline component when
employed as separate and discrete particles in admixture with
particles of the large pore crystalline zeolite.
3. The method of claim 1 wherein catalyst discharged from each
conversion zone is dumped into a common fluid bed of catalyst
particles undergoing stripping with a stripping gas, a portion of
the stripped catalyst is passed to catalyst regeneration and
another portion of the stripped catalyst is combined with freshly
regenerated catalyst for passage to the inlet of the second
conversion zone.
4. The method of claim 1 wherein contaminated catalyst particles
recovered from the second conversion zone stripping operation is
regenerated in stages with at least one stage being a dense fluid
bed regeneration operation.
5. The method of claim 1 wherein regeneration of the catalyst is
accomplished in at least one dispersed phase regeneration operation
discharging adjacent the upper interface of a dense fluid bed of
catalyst being regenerated.
6. In a process for cracking hydrocarbons with crystalline zeolite
cracking catalyst and regenerating coke deactivated catalyst
wherein a plurality of riser conversion zones are relied upon for
processing a first high boiling hydrocarbon feed in one conversion
zone and a second high boiling hydrocarbon feed in another
conversion zone and wherein catalyst separated from the product of
the first conversion zone is passed in admixture with hot
regenerated catalyst to the second conversion zone, the improvement
which comprises,
using as the hydrocarbon conversion catalyst in the combination
operation a mixture of a large pore low coke producing crystalline
zeolite in combination with from 10 to 90 weight percent of a
crystalline zeolite of the ZSM-5 type, forming a suspension of the
catalyst mixture with gaseous hydrocarbons selected from the group
comprising C.sub.3 and lighter dry gases and/or wet gases
comprising C.sub.5 and lighter hydrocarbons in the inlet portion of
each riser conversion zone, converting a first fresh gas oil feed
in admixture with said suspension in the first conversion zone
under operating conditions of temperature, catalyst/oil ratio and
hydrocarbon residence time designed to particularly restrict the
deposition of coke on the low coke producing zeolite catalyst,
converting a second higher coke producing hydrocarbon feed in
admixture with the suspension in the second conversion zone under
operating conditions of temperature, catalyst/oil ratio and
hydrocarbon residence time designed to particularly deposit coke
upon the catalyst during the conversion operation, cyclonically
separating the suspension at the outlet of each conversion zone
into a hydrocarbon phase and a catalyst phase, recovering the
separated hydrocarbon phase, and passing catalyst separated from
the second conversion zone to a catalyst regeneration zone.
7. The method of claim 6 wherein the amount of regenerated catalyst
combined with the catalyst separated from the first conversion zone
is sufficient to provide a catalyst mixture having a temperature
selected from within the range of 950.degree.F. up to about
1,100.degree.F.
8. The method of claim 6 wherein a greater catalyst to oil ratio is
employed in the second conversion zone than employed in the first
conversion zone.
9. The method of claim 6 wherein the catalyst mixture comprises
less than 50 percent ZSM-5 type of conversion catalyst.
10. The method of claim 6 wherein a portion of the catalyst
separated from the first conversion zone is combined with catalyst
separated from the second conversion zone and thereafter passed to
catalyst regeneration.
11. The method of claim 6 wherein a portion of the catalyst
separated from the second conversion zone is recycled thereto and
another separated portion is passed to catalyst regeneration.
12. The method of claim 6 wherein the first hydrocarbon feed is a
low molecular weight fraction and said second hydrocarbon feed is a
high molecular weight fraction.
13. The method of claim 6 wherein said first hydrocarbon feed is a
virgin gas oil and said second hydrocarbon feed is selected from
one or more hydrocarbon materials comprising the group of coker gas
oils, recycle oil and residual stocks.
14. The method of claim 6 wherein preheating of the first and
second hydrocarbon feeds is relied upon to provide a portion of
conversion heat requirements in each conversion zone.
15. The method of claim 6 wherein the concentration of active ZSM-5
type crystalline zeolite in the second conversion zone with respect
to coke deactivated faujasite conversion catalyst is varied as a
function of the quantity of previously used stripped catalyst
passed thereto.
16. The process of claim 6 wherein the catalyst mixture comprises
from 40 to 80 weight percent of the ZSM-5 type crystalline zeolite.
Description
BACKGROUND OF THE INVENTION
The role of catalytic cracking in fluidized and moving bed systems
is well known at this stage of the art. Until recent years
catalytic cracking operations have been forced to use a
silica-alumina cracking catalyst which by today's standards is
considerably less active and particularly is considerably less
selective for performing the catalytic cracking of the hydrocarbon
charge to produce gasoline product. Thus considerable difficulty
has been encountered in the prior systems in obtaining high yields
of conversion products without excess production of the
carbonaceous contaminants.
The present trend in catalytic cracking operations is concerned
with those systems which will use more active and selective
cracking catalysts such as those comprising crystalline zeolites
for performing the conversion of one or more high boiling
hydrocarbon fractions of the same or different boiling range and
coke producing characteristics to gasoline boiling range products.
Thus crystalline zeolite cracking technology necessarily requires
using much more sophisticated cracking systems than those known or
disclosed in the prior art in order to take fully advantage of the
catalyst's conversion capabilities. Many prior art systems and
those converted for the use of high activity crystalline zeolite
cracking catalysts have produced an inefficient operation causing
undue catalyst regeneration, excessive recycle of unconverted
charge and generally inefficient use of the catalyst
composition.
The invention defined herein is concerned with an improved sequence
of conversion steps which will more efficiently utilize the
capabilities of a crystalline zeolite cracking catalyst of high
activity and high selectivity.
SUMMARY OF THE INVENTION
This invention relates to the catalytic cracking of hydrocarbon
oils having finely divided catalyst particles suspended in gasiform
material comprising hydrocarbon reactant material. It relates more
particularly to the catalytic cracking of selected hydrocarbon oils
or fractions thereof with suspended catalyst particles of high
activity and selectivity under conditions selected to produce
gasoline and improve upon the heat balance of the operation.
In a more particular aspect, the method and arrangement of
processing steps of this invention is designed to make maximum use
of low coke forming type crystalline zeolite cracking catalyst in
short contact time riser reactors at elevated cracking
temperatures.
The present invention is concerned with a method of operation which
will circumvent the problems associated with using high activity
crystalline zeolite cracking catalysts at high temperatures and low
catalyst to oil ratios resulting in a used catalyst of relatively
low carbon content.
The essence of the method and system herein described comprises a
dual riser fluid catalytic cracking system effected with freshly
regenerated catalyst in a first riser reactor and a mixture of
freshly regenerated catalyst with catalyst separated from the
conversion products of the first riser in a second riser reactor.
Thus, the present invention is particularly designed and directed
to the use of a relatively high catalyst to oil ratio in a second
riser reactor so as to increase the severity of the cracking
operation therein over that employed in a first riser reactor in
addition to providing more optimum utilization of the available
crystalline zeolite cracking catalyst activity in the processing
combination.
The method and arrangement of catalyst contact steps described
herein may be used to advantage in the following respects:
a. A low coke producing high selective crystalline zeolite cracking
catalyst may be used much more efficiently.
b. The hydrocarbon charge stock and regeneration gas preheat
facilities may be reduced over that required in a prior art system
using the same low coke producing catalyst.
c. More efficient cracking of low molecular weight and higher
molecular weight hydrocarbon fractions may be effected.
d. The processing of high molecular weight charge fractions and/or
heavy recycle oils may be accomplished in a short contact time
riser reactor at a higher or lower average temperature and a higher
catalyst to oil ratio than used for processing a lower molecular
weight oil charge in a separate riser reactor.
e. The high activity low coke producing catalyst may be more
selectively employed for converting hydrocarbon charge materials at
temperatures in the range of 900.degree.F. to 1200.degree.F.
employing catalyst-oil residence time less than 15 seconds and
catalyst to oil ratios as high as 25 to 1.
f. The conversion selectively of the catalyst is enhanced in the
processing sequence by the use of diluent gasiform material to
atomize and/or vaporize either partially or completely the
hydrocarbon charge so as to form relatively fine droplets in
intimate contact with suspended catalyst fines, and effect a
desired reduction in the hydrocarbon partial pressure so as to
optimize the catalytic conversion thereof in the system.
g. The catalyst returned to the regenerator at least from the
second riser reactor is more suitable for providing a major portion
of the endothermic reaction heat requirements of the dual riser
conversion system.
DESCRIPTION OF THE INVENTION
The present invention relates to the catalytic conversion of
hydrocarbons with a selective high activity CAS crystalline
aluminosilicate catalytic composition under elevated temperature
conversion conditions maintained within restricted catalyst-oil
residence contact time and catalyst to oil ratios in order to
produce gasoline. The present invention is concerned with and
relates to the method and means for effecting a selective cracking
of hydrocarbon charge materials to gasoline boiling product with a
very selective low coke producing crystalline alumino-silicate
catalyst composition. In the method and system of this invention
the contact time between catalyst and hydrocarbon varies with the
hydrocarbon charge passed to the selective cracking operation.
Generally the cracking operation effected in a dispersed catalyst
phase relation zone is restricted to orders of magnitude amounting
to only a few seconds up to about 15 seconds and in most instances
the contact time will be restricted depending upon composition of
the hydrocarbon charge to within the range of 4 to 12 seconds. Thus
the concepts essential to practicing the present invention includes
the method and sequence of catalyst cascade which will permit
employing cracking temperatures in the range of 900.degree.F. to
about 1,200.degree.F. at a number of different catalyst to oil
ratios and contact times herein identified. Further salient
features of the present concept include the use of low coke
producing catalysts in the riser reactors, desired catalyst-oil
suspension relationships in a relatively low catalyst inventory
system, and maximizing the use of heat available in the system to
effect the catalytic conversion desired.
The utilization of highly selective low coke providing catalyst
compositions comprising selected crystalline aluminosilicate
catalyst compositions particularly suitable for accomplishing the
processing concept are herein discussed. The processing concepts of
this invention include a restricted contact time between a
suspension of high activity catalyst and hydrocarbon feed being
converted before discharge of the suspension into suitable
separation equipment. Separation equipment particularly suitable
for this purpose comprises one or more cyclone separators at the
discharge end of transfer line reactors which will minimize the
time for separating catalyst particles and hydrocarbon material
without substantially cooling upon discharge from the transfer line
cracking zone. The use of several small cyclones in series reduces
the separation time even though such small cyclones are subject to
the constraint of some oil-catalyst mixing.
Effecting the dispersed catalyst phase cracking operation at
conditions of high temperature in the range of 900.degree. to
1,200.degree.F. and a short contact time less than 15 seconds as
advocated herein provides a low catalyst inventory operation which
is economically attractive and favors the production of superior
product quality in improved quantity.
To enhance the product quality and product distribution or
selectively encountered in the operation discussed herein, a
relatively high boiling virgin feed hydrocarbon material is passed
to a first riser reactor with freshly regenerated catalyst for a
limited residence time therein. A relatively high coke making
hydrocarbon material such as coker gas oil and recycle hydrocarbons
of the cracking operation is passed to a second riser reaction in
contact with a mixture of catalyst comprising catalyst used in the
first reactor and freshly regenerated catalyst. A diluent fluid
such as wet or dry recycle gas or stream may be used in either or
both reactors to reduce hydrocarbon partial pressure and/or control
contact time by increasing total velocity of vapors and coke
particles in the reactor. The improved method of operation herein
defined for converting low and high boiling hydrocarbons either in
a vapor, liquid and/or a partially vaporized condition depending
upon the boiling range of the hydrocarbon charge relies upn a
gasiform diluent material being mixed with the charge to control
the hydrocarbon partial pressure. The diluent also assists in
breaking up the hydrocarbon feed into relatively fine droplets
which more uniformly distribute themselves in intimate contact with
the fine catalyst particles. Gasiform diluent materials which may
be employed in the processing combination of this invention with
varying degrees of success include steam, light gaseous hydrcarbons
known as dry gas (C.sub.3 and lighter hydrocarbons) or wet gaseous
hydrocarbon streams such as those comprising C.sub.4 and C.sub.5
hydrocarbons.
A further important aspect going to the very essence of this
invention concerns itself with an apparatus and catalyst system of
greatly improved flexibility going beyond that heretofore provided
for utilizing the highly selective catalyst and heat available
therefrom in a manner which permits conversion of hydrocarbon
charge material of different coke producing characteristics to more
acceptable products with greater efficiency. Thus it is clearly
evident from that expressed above that operating flexibility to
achieve desired levels of conversion on a per pass basis is
contemplated and such variation may be had by providing one or
more, for example, a plurality of suitably spaced apart hydrocarbon
feed inlet nozzles along the length of riser reactor processing
fresh or virgin feed material. In this embodiment preheated diluent
material is combined with the hot freshly regenerated catalyst to
form a suitable high temperature suspension which is caused to flow
through the riser or confined reaction zone to which point of inlet
of the hydrocarbon feed to be converted. The hydrocarbon feed is
mixed with the flowing high temperature suspension and converted
during transverse of the remaining portion of the reaction zone to
a suspension separator such as one or more cyclone separators. In
any event, it is important that the suspension be separated
substantially immediately after a predetermined residence time so
that overcracking will be avoided. Thus, the prior art methods of
discharging the suspension into large disengaging vessels wherein
separation is effected primarily by velocity reduction and hindered
settling is to be avoided since such separating techniques
encompass high temperature after cracking in a more dilute catalyst
phase. While it is true that temperature quenching of the
suspension may be effected after a predetermined residence time,
such quenching operations are particularly to be avoided. Thus it
is preferred that the suspension in the riser be discharged into
highly efficient cyclone separator equipment so that separated
catalyst of relatively high heat content may be recovered for use
as herein provided.
In one arrangement of this invention the catalyst separated from
the riser reactor processing fresh or virgin feed being of
relatively high temperature approaching 1,000.degree.F. is
subjected to a limited stripping action without substantially
reducing the temperature of the separated catalyst. The stripped
catalyst containing a small amount of residual carbonaceous
material because of the low coke producing characteristics of the
catalyst at the high temperatures employed is then combined with a
quantity of freshly regenerated catalyst in an amount sufficient to
form a mixture of catalyst having an elevated temperature
sufficient for converting recycle feed in a second riser reactor as
herein provided. The suspension passed through the second riser
reactor under cracking conversion conditions is discharged into
suitable separator equipment arranged to effect quick separation of
catalyst particles from hydrocarbon vapors. The hydrocarbon vapors
of the first and second riser reaction zones are recovered and
passed to suitable fractionator equipment. The catalyst separated
from the second suspension is stripped and then regenerated in one
or more regeneration steps which will be adequate for removing
deposited carbonaceous material from the catalyst by burning
thereby heating the catalyst to an elevated temperature suitable
for effecting catalytic conversion reactions.
It is contemplated in the method and arrangement of processing
steps of this invention of introducing the hydrocarbon reactant in
a downstream portion of the riser reactor and relying upon gasiform
material for transporting the cracking catalyst to the point of
hydrocarbon reactant inlet. Thus a plurality of spaced apart
hydrocarbon feed inlets is provided in each riser.
In yet another embodiment, the catalyst separated from the
different riser reactors is collected as a dense fluid bed of
catalyst in a vessel or hopper separated by a common separator
baffle extending upwardly from the vessel bottom and provided with
catalyst flow through slots or passageways so that by adjustment of
catalyst bed level on either side of the baffle, catalyst may be
caused to flow through the slots in either direction as desired to
the catalyst bed exerting a lower pressure. This variation in the
basic arrangement of catalyst flow provides a further flexiblity in
the method of operating the system and such flexibility is most
desirable in present day catalytic operations using the more active
and selective conversion catalysts.
A still further embodiment of the present invention concerns itself
with collecting the separated catalyst from all the riser reactors
in a single dense fluid bed of catalyst wherein the catalyst
particles are mixed and stripped of hydrocarbon vapors with a
suitable stripping gas. In this particular embodiment, provisions
are made for passing the stripped catalyst to one or more stages of
regeneration. In addition, provisions are made for withstanding
catalyst particles from either the upper or lower portion of this
catalyst bed for admixture with fresh regenerated catalyst and use
in the second riser reactor as discussed before.
It will be noted from an understanding of the above expressed
concepts that in all of the embodiments contemplated and discussed
herein, that one feature of the improved processing combination
discussed herein is in the use of a high activity low coke
producing catalyst composition and maximizing, with such a low coke
producing catalyst, the recovery of heat available in the process
by providing catalyst cascading provisions and techniques in the
processing system. That is, the heat retained by the catalyst after
the short time conversion of fresh feed is supplemented by
admixture with freshly regenerated catalyst and thereafter used to
process more refractory materials such as high molecular weight
material fractions, coker gas oils and recycle stocks.
In the method and processing system of this invention it is also
contemplated employing catalyst compositions in admixture with one
another as separate discrete particles or combined as a single
particle of catalyst. Thus, the essential requirement of a low coke
producing crystalline alumino silicate cracking catalyst
composition may be combined with a metal oxide which will enhance
regeneration of the catalyst by converting CO to CO.sub.2 thereby
improving upon the available heat recovery from the system by the
catalyst. Also the low coke producing cracking catalyst composition
may be combined with a ZSM-5 type of catalyst particle which is
also of low coke producing characteristics as a homogeneous
particle mixture or the different catalyst compositions may be used
as separate and discrete particles in the system so that the amount
of one charged to the system may be different than the amount of
the other charged and circulated in the system. In one particular
embodiment, it is contemplated combining the metal oxide such as
chromium oxide with the ZSM-5 type of catalyst particle since such
a composition is more likely to contain less coke than the more
conventional or low coke producing catalyst particles used
therewith to effect primarily cracking of the hydrocarbon charge to
each riser reactor. Under the cracking temperature conditions
employed particularly in the fresh feed riser reactor, olefinic
constituents formed during the cracking step are cyclized to form
aromatics particularly by the ZSM-5 type catalyst. The ZSM-5 type
catalyst may be combined with the low coke producing zeolite
catalyst in amounts ranging from about 10 up to about 90 percent
but more usually will be in a range of from about 40 to about 80
percent. The ZSM-5 type of catalyst is more fully discussed
hereinafter.
The method of operation herein defined operates to take advantage
of heat available in regenerated catalyst, minimizes reaction time
within desired limits and may be used to adjust the residual coke
on catalyst eventually recycled to the regeneration zone. The
amount of coke on the catalyst is controllable and controlled by
using the catalyst separated from the products of the virgin feed
conversion step to effect conversion of the recycle and coker gas
oil in the second transfer line cracking step. Thus the catalyst
used in the first transfer line reactor is cascaded to the second
and the conversion heat requirements of the second is provided by
feed preheat and by adding hot freshly regenerated catalyst to the
catalyst cascaded thereto from the first or virgin feed conversion
step.
The catalyst regeneration techniques envisioned for use in the
concept of the present invention are those which will permit
removal of hydro-carbonaceous material or residue carbonaceous
material from the catalyst particles such as by burning in the
presence of oxygen rich gas or air whereby the catalyst will be
heated to an elevated temperature in the range of 1,100.degree.F.
up to about 1,400.degree.F. It is important to the operation of
this invention that the catalyst be stable at the elevated
temperatures above defined so as to provide a major portion of the
heat required in the catalytic cracking or conversion steps above
discussed. Catalyst regeneration temperatures above 1,200.degree.F.
and up to 1,400.degree.F. are therefore contemplated. The catalyst
regeneration techniques envisioned for use with the cracking steps
of this invention encompass riser or transfer lines regeneration
zones either alone or in combination with dense fluidized catalyst
bed regeneration techniques through which the catalyst moves as a
suspension in oxygen containing regeneration gases under conditions
to maximize retention of generated heat in the catalyst. A
combination of dispersed-catalyst-phase regeneration zones alone or
in combination with a dense fluid catalyst bed operation may be
employed, in which arrangement the catalyst and regeneration gases
will move to achieve the desired removal of carbonaceous deposits.
It is contemplated employing the sequence of riser dispersed
catalyst phase operations through which the catalyst sequentially
moves as a suspension in regeneration gas under regenerating
temperature conditions with the amount of regeneration gas, such as
air, being passed to each reactor being controlled or limited so as
to limit temperatures encountered during regeneration of the
catalyst. Regeneration temperatures may also be raised by combining
a carbon monoxide conversion metal oxide with the catalyst as
mentioned above and by increasing the amount of coke or
hydrocarbonaceous material deposited on the catalyst during the
cracking step before being passed to the regeneration step. Since
air is a common regeneration gas, varying the amount of
carbonaceous material on catalyst passed to the regenerator is most
effective for temperature control. In another arrangement one or
more riser regenerators may be used in conjunction with a more
dense catalyst bed regeneration arrangement, such as provided by a
dense fluid catalyst bed or the counter-current flow of catalyst
downwardly through a regeneration zone and counter-current to
upwardly flowing regeneration gas. In such arrangements the more
dense catalyst suspension as it moves downwardly through the system
will encounter an ever increasing concentration of oxygen rich
regeneration gas. In these systems and arrangement of regeneration
zones, it is preferred for economic reasons that the equipment
utilized be sized to minimize the catalyst inventory of the system
consistent with providing a commercially attractive operating
arrangement.
It is clear from the above discussion that a considerably greater
utilization of available heat is provided in the cracking operation
discussed and that more than one such combination of riser reactors
may be employed with a common catalyst regeneration system. The
arrangement herein discussed is particularly attractive for
processing the same or a combination of hydrocarbon charge
materials (i.e., heavy gas oil, light gas oil, recycle, straight
run gasoline or naphtha) and each combination relies upon and
maintains its own reaction parameters such as temperature, space
velocity reaction time and catalyst-oil ratio. The arrangement is
also attractive for multiple passing a hydrocarbon charge to be
converted so as to increase gasoline product yields.
In the cracking operation discussed herein the temperatures are
preferably of a high order of magnitude in the range of
1,000.degree.F. up to about 1,300.degree. or 1400.degree.F.
However, it is contemplated under some circumstances employing a
lower temperature for some specific applications. Generally the
temperature will not be below 900.degree.F. Operating pressures may
be in the range of from about atmospheric pressure up to several
atmospheres depending upon the system employed. Generally the
pressure will be selected from within the range of from about 15
psig. up to about 50 or 75 psig. Operating space velocities are
relatively high and are selected to provide catalyst-oil contact
times in the virgin feed cracking step in the range of from about 1
second up to about 12 seconds and more usually in the range of 4 to
8 seconds. Contact times used in the recycle cracking step will be
in the range of about 2 to 15 seconds and may be a few seconds
longer than employed for of virgin feed cracking step. On the other
hand, the catalyst to oil ratio used for converting virgin feed
will be from about 4 to 10 to 1 and from about 6 to 25 to 1 for
converting the recycle and coker feed material.
The method of the present invention is particularly suitable for
effecting high temperature, above 1000.degree.F., catalytic
cracking operations which have been found to yield products of
superior quality and quantity relative to that attainable by more
convetional operations now relied upon. Thus the concepts discussed
herein are particularly directed and designed for high temperature,
short contact time cracking operations in the presence of a low
coke producing CAS catalyst so as to take advantage of product
quality and quantity.
The hydrocarbon conversion catalyst employed in the method and
arrangement of process step of this invention comprises a
crystalline aluminosilicate cracking catalyst. In particular it is
a crystalline aluminosilicate cracking catalyst of unusual
properties in that at high temperature conversion of virgin charge
stocks, a smaller than usual amount of carbonaceous material is
deposited on the conversion catalyst.
In the catalytic cracking process of the present invention, there
may be employed, as mentioned herein, catalyst composition
comprising two distinct cracking components. One component is a
catalytically active form of a ZSM-5 type zeolite. The other
component may be any other suitable zeolite catalytic cracking
component which is active with respect to converting the molecular
constituents of the hydrocarbon charge and products of cracking to
desired components. Generally speaking, the cracking component may
be a porous cracking component such as silica/alumina and more
particularly a crystalline aluminosilicate zeolite cracking
component having relatively uniform pore dimensions and a pore size
selected from within the range of 6 to 15 Angstrom units. These
relatively large pore zeolite components will admit both normal and
iso-aliphatics. Particularly desirable zeolites include the
synthetic faujasites known as zeolite X and zeolite Y. Other large
pore zeolites may also be employed. The weight ratio of the porous
siliceous cracking component, (e.g., synthetic faujasite) to the
ZSM-5 component may be selected from between 0.1 and 20.
The ZSM-5 type zeolites which may be used in combination with
relatively large pore zeolites in the novel cracking process of
this invention are considered to be of a relatively intermediate
pore size. Thus, the ZSM-5 type catalysts used in the novel process
of this invention will allow entry into their internal pore
structure or normal aliphatic compounds and slightly branched
aliphatic compounds, particularly monomethyl substituted compounds,
yet substantially exclude all compounds containing at least a
quaternary carbon atom or having a molecular dimension equal to or
substantially greater than a quaternary carbon atom. Additionally,
aromatic compounds having side chains similar to the normal
aliphatic compounds and slightly branched aliphatic compounds above
described could have side chains enter the internal pore structure
of the instant catalysts. Thus, if one were to measure the
selectivity of the ZSM-5 type materials employed in the process of
this invention with regard to their ability to sorb n-hexane in
admixture with 2-methyl pentane, i.e., the ability to selectively
sorb hexane from a mixture of the same with isohexane, these
catalysts would have to be stated as being non-shape selective. It
should be immediately apparent, however, that the term selectivity
has a far greater significance than merely the ability to
preferentially distinguish between normal paraffins and
isoparaffins. Selectivity on shape is theoretically possible at any
shape or size although, quite obviously, such selectivity might not
result in an advantageous catalyst for any and all hydrocarbon
conversion processes.
While not wishing to be bound by any theory of operation
nevertheless, it appears that the crystalline zeolitic materials of
the ZSM-5 type employed in the instant invention cannot simply be
characterized by the recitation of a pore size or a range of pore
sizes. It appears that the uniform pore openings of this new type
of zeolite are not approximately circular in nature, as is usually
the case in the heretofore employed zeolites, but rather, are
approximately elliptical in nature. Thus, the pore openings of the
instant zeolitic materials have both a major and a minor axes, and
it is for this reason that the unusual and novel molecular sieving
effects are achieved. This elliptical shape can be referred to as a
"keyhole." It appears that the minor axis of the elliptical pores
in the zeolites apparently have an effective size of about 5.5
Angstrom units. The major axis appears to be somewhere between 6
and about 9 Angstrom units. The unique keyhole molecular sieving
action of these materials is presumably due to the presence of
these approximately elliptically shaped windows controlling access
to the internal crystalline pore structure.
A test method has been devised in order to determine whether or not
a zeolite possess the unique molecular sieving properties desired
to be combined with a large pore zeolite in order to carry out the
novel conversion process of this invention. In the test method a
candidate zeolite free from any matrix or binder is initially
converted to the so-called acid or hydrogen form. This procedure
involves exhaustive exchange with an ammonium chloride solution in
order to replace substantially all metallic cations originally
present. The sample is then dried, sized to 20-30 mesh and calcined
in air for 16 hours at 550.degree.F. One gram of the so-treated
zeolite is then contacted with benzene at a pressure of 12 mmHg. at
a temperature of 25.degree.C. for a time period of 2 hours. Another
gram sample is contacted with mesitylene at a pressure of 15 mmHg.
at a temperature of 25.degree.C. for a period of 6 hours. A
preferred zeolite is one whose acid form will adsorb at least 3.0
weight percent benzene and less than 1.5 weight percent mesitylene
at the above-recited conditions.
Examples of zeolitic materials which are operable in the process of
this invention are ZSM-5 type which family includes not only ZSM-5
but also ZSM-8 zeolites. ZSM-5 type materials are disclosed and
claimed in copending application Ser. No. 865,472, filed October
10, 1969, now U.S. Pat. No. 3,702,886 and ZSM-8 is disclosed and
claimed in copending application Ser. No. 865,418 filed Oct. 10,
1969 now abandoned. A process utilizing a combination of ZSM-5 type
zeolites and large pore zeolites is disclosed in Ser. No. 78,573
filed Oct. 6, 1970.
The family of ZSM-5 compositions has the characteristic X-ray
diffraction pattern set forth in Table 1, hereinbelow. ZSM-5
compositions can also be identified, in terms of mole ratios of
oxides, as follows:
0.9 .+-. 0.2 M.sub.2/n O : M.sub.2 O.sub.3 : 5-100 YO.sub.2 :
zH.sub.2 O
wherein M is a cation, n is the valence of said cation, W is
selected from the group consisting of aluminum and gallium, Y is
selected from the group consisting of silicon and germanium, and z
is from 0 to 40. In a preferred synthesized form, the zeolite has a
formula, in terms of mole ratios of oxides, as follows:
0.9 .+-. 0.2 M.sub.2/n O : Al.sub.2 O.sub.3 : 5-100 SiO.sub.2 : z
H.sub.2 O
and M is selected from the group consisting of a mixture of alkali
metal cations, especially sodium, and tetraalkylammonium cations,
the alkyl groups of which preferably contain 2-5 carbon atoms.
In a preferred embodiment of ZSM-5, W is aluminum, Y is silicon and
the silica/alumina mole ratio is at least 10 and ranges up to about
60.
Members of the family of ZSM-5 zeolites possess a definite
distinguishing crystalline structure whose X-ray diffraction
pattern shows the following significant lines:
TABLE 1 ______________________________________ Interplanar Spacing
d(A) Relative Intensity ______________________________________ 11.1
.+-. 0.2 S 10.0 .+-. 0.2 S 7.4 .+-. 0.15 W 7.1 .+-. 0.15 W 6.3 .+-.
0.1 W 6.04.+-. 0.1 W 5.97.+-. 0.1 W 5.56.+-. 0.1 W 5.01.+-. 0.1 W
4.60.+-. 0.08 W 4.25.+-. 0.08 W 3.85.+-. 0.07 VS 3.71.+-. 0.05 S
3.64.+-. 0.05 M 3.04.+-. 0.03 W 2.99.+-. 0.02 W 2.94.+-. 0.02 W
______________________________________
These values as well as all other X-ray data were determined by
standard techniques. The radiation was the K-alpha doublet of
copper, and scintillation counter spectrometer with a strip chart
pen recorder was used. The peak heights, I, and the positions as a
function of 2 times theta, where theta is the Bragg angle, were
read from the spectrometer chart. From these the relative
intensities, 100 I/I.sub.o, where I.sub.o is the intensity of the
strongest line or peak, and d (obs.), the interplanar spacing in A,
corresponding to the recorded lines, were calculated. In Table 1
the relative intensities are given in terms of the symbols S =
strong, M = medium, MS = medium strong, MW = medium weak and VS =
very strong. It should be understood that this X-ray diffraction
pattern is characteristic of all the species of ZSM-5 compositions.
Ion exchange of the sodium ion with cations reveals substantially
the same pattern with some minor shifts in interplanar spacing and
variation in relative intensity. Other minor variations can occur
depending on the silicon to aluminum ratio of the particular
sample, as well as if it has been subjected to thermal treatment.
Various cation exchanged forms of ZSM-5 have been prepared. X-ray
powder diffraction patterns of several of these forms are set forth
below. The ZSM-5 forms set forth below are all
aluminosilicates.
TABLE 2 ______________________________________ X-Ray Diffraction
ZSM-5 Powder in Cation Exchanged Forms d Spacing Observed
______________________________________ As Made HCl NaCl CaCl.sub.2
RECl.sub.3 AgNO.sub.3 ______________________________________ 11.15
11.16 11.19 11.19 11.19 11.19 10.01 10.03 10.05 10.01 10.06 10.01
9.74 9.78 9.80 9.74 9.79 9.77 -- -- 9.01 9.02 -- 8.99 8.06 -- -- --
-- -- 7.44 7.46 7.46 7.46 7.40 7.46 7.08 7.07 7.09 7.11 -- 7.09
6.70 6.72 6.73 6.70 6.73 6.73 6.36 6.38 6.38 6.37 6.39 6.37 5.99
6.00 6.01 5.99 6.02 6.01 5.70 5.71 5.73 5.70 5.72 5.72 5.56 5.58
5.58 5.57 5.59 5.53 5.37 -- 5.38 5.37 5.38 5.37 5.13 5.11 5.14 5.12
5.14 -- 4.99 5.01 5.01 5.01 5.01 5.01 -- -- 4.74 -- -- -- 4.61 4.62
4.62 4.61 4.63 4.62 -- -- 4.46 4.46 -- 4.46 4.36 4.37 4.37 4.36
4.37 4.37 4.26 4.27 4.27 4.26 4.27 4.27 4.08 -- 4.09 4.09 4.09 4.09
4.00 4.01 4.01 4.00 4.01 4.01 3.84 3.85 3.85 3.85 3.85 3.86 3.82
3.82 3.82 3.82 3.83 3.82 3.75 3.75 3.75 3.76 3.76 3.75 3.72 3.72
3.72 3.72 3.72 3.72 3.64 3.65 3.65 3.65 3.65 3.65 -- 3.60 3.60 3.60
3.61 3.60 3.48 3.49 3.49 3.48 3.49 3.49 3.44 3.45 3.45 3.44 3.45
3.45 3.34 3.35 3.36 3.35 3.35 3.35 3.31 3.31 3.32 3.31 3.32 3.32
3.25 3.25 3.26 3.25 3.25 3.25 3.17 -- -- 3.17 3.18 -- 3.13 3.14
3.14 3.14 3.15 3.14 3.05 3.05 3.05 3.04 3.06 3.05 2.98 2.98 2.99
2.98 2.99 2.99 -- -- -- -- 2.97 -- -- 2.95 2.95 2.94 2.95 2.95 2.85
2.87 2.87 2.87 2.87 2.87 2.80 -- -- -- -- -- 2.78 -- -- 2.78 --
2.78 2.73 2.74 2.74 2.73 2.74 2.74 2.67 -- -- 2.68 -- -- 2.66 -- --
2.65 -- -- 2.60 2.61 2.61 2.61 2.61 2.61 -- 2.59 -- 2.59 -- -- 2.57
-- 2.57 2.56 -- 2.57 2.50 2.52 2.52 2.52 2.52 -- 2.49 2.49 2.49
2.49 2.49 2.49 -- -- -- 2.45 -- -- 2.41 2.42 2.42 2.42 2.42 -- 2.39
2.40 2.40 2.39 2.40 2.40 -- -- -- 2.38 2.35 2.38 -- 2.33 -- 2.33
2.32 2.33 -- 2.30 -- -- -- -- -- 2.24 2.23 2.23 -- -- -- 2.20 2.21
2.20 2.20 -- -- 2.18 2.18 -- -- -- -- -- 2.17 2.17 -- -- -- 2.13 --
2.13 -- -- -- 2.11 2.11 -- 2.11 -- -- -- -- 2.10 2.10 -- -- 2.08
2.08 -- 2.08 2.08 -- -- 2.07 2.07 -- -- -- -- -- 2.04 -- -- 2.01
2.01 2.01 2.01 2.01 2.01 1.99 2.00 1.99 1.99 1.99 1.99 -- -- --
1.97 1.96 -- 1.95 1.95 1.95 1.95 1.95 -- -- -- -- -- 1.94 -- --
1.92 1.92 1.92 1.92 1.92 1.91 -- -- -- 1.92 -- -- -- -- -- 1.88 --
1.87 1.87 1.87 1.87 1.87 1.87 -- 1.86 -- -- -- -- 1.84 1.84 -- --
1.84 1.84 1.83 1.83 1.83 1.83 1.83 -- 1.82 -- 1.81 -- 1.82 -- 1.77
1.77 1.79 1.78 -- 1.77 1.76 1.76 1.76 1.76 1.76 1.76 -- -- 1.75 --
-- 1.75 -- 1.74 1.74 1.73 -- -- 1.71 1.72 1.72 1.71 -- 1.70 1.67
1.67 1.67 -- 1.67 1.67 1.66 1.66 -- 1.66 1.66 1.66 -- -- 1.65 1.65
-- -- -- -- 1.64 1.64 -- -- -- 1.63 1.63 1.63 1.63 1.62 -- 1.61
1.61 1.61 -- 1.61 1.58 -- -- -- -- -- -- 1.57 1.57 -- 1.57 1.57 --
-- 1.56 1.56 1.56 -- ______________________________________
Zeolite ZSM-5 can be suitably prepared by preparing a solution
containing tetrapropyl ammonium hydroxide, sodium oxide, an oxide
of aluminum or gallium, an oxide of silica and water and having a
composition, in terms of mole ratios of oxides, falling within the
following ranges:
TABLE 3 ______________________________________ Particularly Broad
Preferred Preferred ______________________________________
OH.sup.-/SiO.sub.2 0.07-1.0 0.1-0.8 0.2-0.75 R.sub.4 N+/(R.sub.4
N.sup.+ +Na.sup.+) 0.2-0.95 0.3-0.9 0.4-0.9 H.sub.2 O/OH.sup.-
10-300 10-300 10-300 YO.sub.2 /W.sub.2 O.sub.3 5-100 10-60 10-40
______________________________________
wherein R is propyl, W is aluminum and Y is silicon maintaining the
mixture until crystals of the zeolite are formed. Thereafter the
crystals are separated from the liquid and recovered. Typical
reaction conditions consist of heating the foregoing reaction
mixture to a temperature of from about 75.degree.C to 175.degree.C
for a period of time of from about six hours to 60 days. A more
preferred temperature range is from about 90.degree. to
150.degree.C with the amount of time at a temperature in such range
being from about 12 hours to 20 days.
The digestion of the gel particles is carried out until crystals
form. The solid product is separated from the reaction medium, as
by cooling the whole to room temperature, filtering, and water
washing.
ZSM-5 is preferably formed as an aluminosilicate. The composition
can be prepared utilizing materials which supply the appropriate
oxide. such compositions include for an aluminosilicate, sodium
aluminate, alumina, sodium silicate, silica hydrosol, silica gel,
silicic acid, sodium hydroxide and tetrapropylammonium hydroxide.
It will be understood that each oxide component utilized in the
reaction mixture for preparing a member of the ZSM-5 family can be
supplied by one or more initial reactants and they can be mixed
together in any order. For example, sodium oxide can be supplied by
an aqueous solution of sodium hydroxide, or by an aqueous solution
of sodium silicate; tetrapropylammonium cation can be supplied by
the bromide salt. The reaction mixture can be prepared either
batchwise or continuously. Crystal size and crystallization time of
the ZSM-5 composition will vary with the nature of the reaction
mixture employed. ZSM-8 can also be identified, in terms of mole
ratios of oxides, as follows:
0.9 .+-. 0.2 M.sub.2/n O : Al.sub.2 O.sub.3 : 5-100 SiO.sub.2 : z
H.sub.2 O
wherein M is at least one cation, n is the valence thereof and z is
from 0 to 40. In a preferred synthesized form, the zeolite has a
formula, in terms of mole ratios of oxides, as follows:
0.9 .+-. 0.2 M.sub.2/n O : Al.sub.2 O.sub.3 : 10-60 SiO.sub.2 : z
H.sub.2 O
and M is selected from the group consisting of a mixture of alkali
metal cations, especially sodium, and tetraethylammonium
cations.
ZSM-8 possesses a definite distinguishing crystalline structure
having the following X-ray diffraction pattern:
TABLE 4 ______________________________________ dA.degree. I/I.sub.o
I/I.sub.o dA.degree. ______________________________________ 11.1 46
4 2.97 10.0 42 3 2.94 9.7 10 2 2.86 9.0 6 1 2.78 7.42 10 4 2.73
7.06 7 1 2.68 6.69 5 3 2.61 6.35 12 1 2.57 6.04 6 1 2.55 5.97 12 1
2.51 5.69 9 6 2.49 5.56 13 1 2.45 5.36 3 2 2.47 5.12 4 3 2.39 5.01
7 1 2.35 4.60 7 1 2.32 4.45 3 1 2.28 4.35 7 1 2.23 4.25 18 1 2.20
4.07 20 1 2.17 4.00 10 1 2.12 3.85 100 1 2.11 3.82 57 1 2.08 3.75
25 1 2.06 3.71 30 6 2.01 3.64 26 6 1.99 3.59 2 2 1.95 3.47 6 2 1.91
3.43 9 3 1.87 3.39 5 1 1.84 3.34 18 2 1.82 3.31 8 3.24 4 3.13 3
3.04 10 2.99 6 ______________________________________
Zeolite ZSM-8 can be suitably prepared by reacting a solution
containing either tetraethylammonium hydroxide or
tetraethylammonium bromide together with sodium oxide, aluminum
oxide, and an oxide of silica and water.
The relative operable proportions of the various ingredients have
not been fully determined and it is to be immediately understood
that not any and all proportions of reactants will operate to
produce the desired zeolite. In fact, completely different zeolites
can be prepared utilizing the same starting materials depending
upon their relative concentration and reaction conditions as is set
forth in U.S. Pat. No. 3,308,069. In general, however, it has been
found that when tetraethylammonium hydroxide is employed, ZSM-8 can
be prepared from said hydroxide, sodium oxide, aluminum oxide,
silica and water by reacting said materials in such proportions
that the forming solution has a composition in terms of mole ratios
of oxides falling within the following range
SiO.sub.2 /Al.sub.2 O.sub.3 -- from about 10 to about 200
Na.sub.2 O/tetraethylammonium hydroxide -- from about 0.05 to
0.20
Tetraethylammonium hydroxide/SiO.sub.2 -- from about 0.08 to
1.0
H.sub.2 O/tetraethylammonium hydroxide -- from about 80 to about
200
Thereafter, the crystals are separated from the liquid and
recovered. Typical reaction conditions consist of heating the
foregoing reaction mixture to a temperature of from about
100.degree.C to 175.degree.C for a period of time of from about 6
hours to 60 days. A more preferred temperature range is from about
150.degree. to 175.degree.C with the amount of time at a
temperature in such range being from about 12 hours to 8 days.
The foregoing product is dried, e.g., at 230.degree.F. for from
about 8 to 24 hours. Of course, milder conditions may be employed
if desired, e.g., room temperature under vacuum.
As has heretofore been stated, a zeolite of the ZSM-5 type
above-described is used in conjunction with a large pore zeolite,
i.e., one having a pore size greater than 7 Angstrom units which
has the ability to act upon substantially all the components
usually found in a commercial gas oil. Large pore aluminosilicates
of this type are well known and include natural and synthetic
faujasite of both the X and Y type, as well as zeolite L. Of these
materials, zeolite Y is particularly preferred.
Both the large pore zeolites and the ZSM-5 type zeolites used in
the instant invention usually have the original cations associated
therewith replaced by a wide variety of other cations according to
techniques well known in the art. Typical replacing cations would
include hydrogen, ammonium and metal cations including mixtures of
the same. Of the replacing metallic cations, particular preference
is given to cations of rare earth, Mg.sup.+.sup.+, Zn.sup.+.sup.+,
Mn.sup.+.sup.+, Al.sup.+.sup.+.sup.+, and Ca.sup.+.sup.+.
Typical ion exchange techniques would be to contact the particular
zeolite with a salt of the desired replacing cation or cations.
Although a wide variety of salts can be employed, particular
preference is given of chlorides, nitrates and sulfates.
Representative ion exchange techniques are disclosed in a wide
variety of patents including U.S. Pat. Nos. 3,140,249; 3,140,251;
and 3,140,253.
Following contact with the salt solution of the desired replacing
cation, the zeolites may be washed with water and dried at a
temperature ranging from 150.degree.F. to about 600.degree.F. and
thereafter heated in air or other inert gas at temperatures ranging
from about 500.degree. to 1500.degree.F for periods of time ranging
from 1 to 48 hours or more. It has been further found in accordance
with the invention that catalysts of improved selectivity and
having other beneficial properties in catalytic cracking are
obtained by subjecting the zeolite to treatment with steam at
elevated temperatures ranging from 800.degree. to 1,600.degree.F
and preferably 1,000.degree.F and 1,500.degree.F. The treatment may
be accomplished in atmospheres consisting partially or entirely of
steam. This treatment may be accomplished within a commercial
cracking unit, e.g., by gradual addition of the unsteamed catalyst
to the unit.
A similar treatment can be accomplished at lower temperatures and
elevated pressures, e.g. 350.degree.-700.degree.F at 10 to about
200 atmospheres.
The novel catalyst composites of this invention, in a particular
embodiment, comprise a physical mixture of at least two different
cracking components, one being an aluminosilicate having a pore
size greater than about 7 Angstrom units. In one embodiment, a
mixture of catalyst particles is used in which each particle
contains only one of the two types of zeolites. Thus, for example,
a mixture of spray dried particles comprising ZSM-5 type crystals
in a matrix and particles comprising faujasite crystals in a matrix
may be added as make-up to the cracking unit. Alternatively, the
catalyst components may be pelleted, cast, molded, spray-dried or
otherwise formed into pieces of desired size and shape such as
rods, spheres, pellets, etc.
The compositing of the aluminosilicate with an inorganic oxide can
be achieved by several methods wherein the aluminosilicates are
reduced to a particle size less than 40 microns, preferably less
than 10 microns, and intimately admixed with an inorganic oxide
while the latter is in a hydrous state such as in the form of
hydrosol, hydrogel, wet gelatinous precipitate, or in a dried
state, or a mixture thereof. Thus, finely divided aluminosilicates
can be mixed directly with a siliceous gel formed by hydrolyzing a
basic solution of alkali metal silicate with an acid such as
hydrochloric, sulfuric, acetic, etc. The mixing of the three
components can be accomplished in any desired manner, such as in a
ball mill or other types of mills. The aluminosilicates also may be
dispersed in a hydrosol obtained by reacting an alkali metal
silicate with an acid or alkaline coagulant. The hydrosol is then
permitted to set in mass to a hydrogel which is thereafter dried
and broken into pieces of desired shape or dried by conventional
spray drying techniques or dispersed through a nozzle into a bath
of oil or other water-immiscible suspending medium to obtain
spheroidally shaped " bead" particles of catalyst such as described
in U.S. Pat. No. 2,384,946. The aluminosilicate siliceous gel thus
obtained is washed free of soluble salts and thereafter dried
and/or calcined as desired.
In a like manner, the aluminosilicates may be incorporated with an
aluminiferous oxide. Such gels and hydrous oxides are well known in
the art and may be prepared, for example, by adding ammonium
hydroxide, ammonium carbonate, etc. to a salt of aluminum, such
aluminum chloride, aluminum sulfate, aluminum nitrate, etc., in an
amount sufficient to form aluminum hydroxide which, upon drying, is
converted to alumina. The aluminosilicate may be incorporated with
the aluminiferous oxide while the latter is in the form of
hydrosol, hydrogel, or wet gelatinous precipitate or hydrous oxide,
or in the dried state.
The catalytically inorganic oxide matrix may also consist of a
plural gel comprising a predominant amount of silica with one or
more metals or oxides thereof selected from Groups IB, II, III, IV,
V, VI, VII, and VIII of the Periodic Table. Particular preference
is given to plural gels of silica with metal oxides of Groups IIA,
III and IVa of the Periodic Table, especially wherein the metal
oxide is rare earth oxide, magnesia, alumina, zirconia, titania,
beryllia, thoria, or combination thereof. The preparation of plural
gels is well known and generally involves either separate
precipitation or coprecipitation techniques, in which a suitable
salt of the metal oxide is added to an alkali metal silicate and an
acid or base, as required, is added to precipitate the
corresponding oxide. The silica content of the siliceous gel matrix
contemplated herein is generally within the range of 55 to 100
weight percent with the metal oxide content ranging from 0 to 45
percent.
The catalyst product can be heated in steam or in other
atmospheres, e.g., air, near the temperature contemplated for
conversion but may be heated to operating temperatures initially
during use in the conversion process. Generally, the catalyst is
dried between 150.degree.F. and 600.degree.F. and thereafter may be
calcined in air, steam, nitrogen, helium, flue gas or other gases
not harmful to the catalyst product at temperatures ranging from
about 500.degree.F. to 1,600.degree.F. for periods of time ranging
from 1 to 48 hours or more. It is to be understood that the
aluminosilicate can also be calcined prior to incorporation into
the inorganic oxide gel. It is also to be understood that the
aluminosilicate or aluminosilicates need not be ion exchanged prior
to incorporation in a matrix but can be so treated during and/or
after incorporation into the matrix. Preferably, the zeolite is
metal exchanged, calcined and thereafter given a second exchange
with a metal or hydrogen precursor.
It has been further found in accordance with the invention that
catalysts of improved selectivity and having other beneficial
properties in gas oil cracking are obtained by subjecting the
catalyst product to a mild steam treatment carried out at elevated
temperatures of 800.degree.F. to 1,600.degree.F. and preferably at
temperatures of about 1,000.degree.F. to 1,500.degree.F. The
treatment may be accomplished in an atmosphere of 100 percent steam
or in an atmosphere consisting of steam and air or a gas which is
not harmful to the aluminosilicate. The steam treatment apparently
provides beneficial properties in the aluminosilicate compositions
and can be conducted before, after or in place of the calcination
treatment.
The particle size of each type of zeolite making up the catalyst
system is not narrowly critical but should be less than 100 microns
and particle sizes within the range of from less than 0.1 to 10
microns are preferred. It is also to be noted that each individual
component in the catalyst system need not be of the same particle
size.
The particular proportion of one component to the other in the
catalyst system is also not narrowly critical and can vary over an
extremely wide range. However, it has been found that for most
purposes the weight ratio of the ZSM-5 type aluminosilicate to the
larger pore size aluminosilicate with which it is mixed can range
from 0.05:1 up to 10:1 and preferably from 1:3 up to 2:1 and still
more preferably 1:2 to 1:1.
The ZSM-5 type crystalline aluminosilicates and the crystalline
aluminosilicates with pores greater than 7 Angstroms may be added
to a cracking unit as a mixture of crystallites within the same
particles of catalyst composite, whether the particles are beads,
extrudates, or spray-dried microspheres. Alternatively, a mixture
of particles of fluidizable particle size may be added to the
cracking unit, some particles containing only the ZSM-5 type
aluminosilicate crystallites and the other particles containing
only the large pore aluminosilicate crystallites. In either case,
the ratio of ZSM-5 type aluminosilicates to large pore
aluminosilicates should be within the range of 1:20 to 10:1. The
ratio of aluminosilicates within this range is controlled to
produce the most desirable balance of high octane gasoline and
C.sub.3 and C.sub.4 olefin yields.
Within the above description of the aluminosilicates which can be
physically admixed in a porous matrix to prepare the catalysts of
this invention, it has been found that certain aluminosilicates
provide superior results when employed in catalytic cracking
operations.
First of all, it is preferred that there be a limited amount of
alkali metal cations associated with the aluminosilicates since the
presence of alkali metals tends to suppress or limit catalytic
properties, the activity of which as a general rule decreases with
increasing content of alkali metal cations. Therefore, it is
preferred that the aluminosilicates contain no more than 0.25
equivalents per gram atom of aluminum and more preferably no more
than 0.15 equivalents per gram atom of aluminum of alkali metal
cations.
With regard to the metal cations associated with the large pore
aluminosilicate, the general order of preference is first cations
of trivalent metals, followed by cations of divalent metals, with
the least preferred being cations of monovalent metals. Of the
trivalent metal cations, the most preferred are rare earth metal
cations, either individually or as a mixture of rare earth metal
cations.
Additionally, it is particularly preferred to have at least some
protons or proton precursors associated with the
aluminosilicate.
It is also preferred that both the aluminosilicates have an atomic
ratio of silicon to aluminum of at least 1.25 preferably 1.8 and
even more desirably at least 2.0.
It is to be understood, however, that this invention includes the
use of catalyst compositions wherein both aluminosilicates are of
the same class, e.g., both metal aluminosilicates; of different
classes, e.g., one metal and one acid aluminosilicate; in the same
matrix or in different matrixes, i.e., one aluminosilicate in
silica-alumina and the other in silica-zirconia.
In the process of the present invention, the cracking catalyst has
a particle size such that it can be passed in fluid flow through
the risers, catalyst separators, stripper, transfer conduits and
the regenerator. The particle size will generally be between 10 and
100 microns in diameter, preferably 40 to 80 microns. A particle
size of about 60 microns diameter is considered optimum.
Low coke producing crystalline aluminosilicate cracking catalysts
of the type which may be used with particular advantage in the
method and processing schemes of this invention are more
particularly exemplified by the following table and FIGS. 1 and 2
presented herewith. The table provides the chemical properties of
several crystalline zeolite containing cracking catalysts having
different selectivity characteristics and particularly coke
producing characteristics.
TABLE 5
__________________________________________________________________________
CHEMICAL PROPERTIES OF CRYSTALLINE ZEOLITE CRACKING CATALYSTS
__________________________________________________________________________
Sieve Chemical Analysis, % Wt Loss Calcined 3 hr/1200.degree.F/Air
Mean Cata- Type SiO.sub.2.sup.1 Al.sub.2 O.sub.3 RE.sub.2 O.sub.3
Na.sub.2 O ZrO.sub.2 SO.sub.4 on Surface Pore Packed Dia- lyst and
Igni- Area, Vol Density, meter Level tion m.sup.2 /g cc/g g/cc
Micron
__________________________________________________________________________
A 7.5% REX 84.4 13.8 1.84 0.05 415 0.93 0.47 77 B 5% REY 85.0 14.0
0.96 0.04 13-15 442 0.89 0.46 C 10% REY 77.4 18.0 2.5 0.12 2.0
<0.2 12-14 328 0.67 0.56 D 10-20% HY 85.5 14.0 -- 0.05 486 0.79
E 10% REX 66.0 31.0 2.9 0.09 280 0.58 0.63 79 F 12% REX 64.6 32.6
2.8 0.04 269 0.69 0.55 70 G 12-14% REX 58.7 35.0 4.5 0.07 0.12 15.4
270 0.66 H 15% REY.sup.2 66.6 29.3 3.6 0.46 329 0.58 0.58 I 15% REY
60.5 36.8 2.5 0.22 203 0.52 0.65 71 J 3% REX 73.0 25.7 0.87 0.04
295 0.55 0.61 65 K 6.7% REX 64.0 33.0 1.67 0.11 227 0.57 0.64 L 5%
REY 65.7 33.2 1.00 0.05 260 0.55 0.64 M 4 -5% REY 67.5 31.7 0.66
0.14 N 7-8% REY 82.4 15.2 2.30 0.07 0.68 0.46 O 11-12% REX.sup.3
51.7 45.5 2.20 0.62 169 0.38 0.79 51 P 10% REY 53.4 43.8 2.01 0.77
Q 5% REY 84.8 14.0 1.18 0.05 0.18 14 561 0.72 0.54 R > 10% REY
71.1 26.7 2.06 0.09 S 20% Y 61.6 35.9 0.0 0.68 325 0.37 0.80 T --
87 13 -- 0.02 450 0.68 0.58 70 U -- 75 25 -- 0.03 340 0.92 0.50
__________________________________________________________________________
.sup.1 By difference on a 100% basis. .sup.2 Crystallinity shows
30% shift toward X-type material. .sup.3 Crystallinity indicates
sieve of an X-Y type.
BRIEF DESCRIPTION OF THE DRAWINGS
FIG. 1 is a plot of data obtained, coke on charge versus conversion
after thermal treatment of the identified catalyst
compositions.
FIG. 2, on the other hand, is a plot of data obtained, coke on
charge versus conversion after steam treatment of identified
catalyst compositions.
FIG. 3 diagrammatically represents one arrangement of processing
steps comprising a dual riser operation for catalytic conversion of
hydrocarbon reactant and interconnecting catalyst transfer means
for conveying catalyst particles to the regenerator and from the
regenerator to the riser conversion operations.
FIG. 4 diagrammatically shows a modified embodiment of FIG. 3
wherein the dual riser catalyst collection hopper is separated by a
vertical baffle provided with catalyst flow through passageways for
passing catalyst in either direction through the vertical
baffle.
FIG. 5 diagrammatically shows a further modification of FIG. 3
comprising a common fluid bed of catalyst discharged from each
riser and catalyst withdrawal means above the stripping section for
providing the catalyst to the second riser reactor.
DISCUSSION OF SPECIFIC EMBODIMENTS
The operating conditions employed in obtaining the data plotted in
FIGS. 1 and 2 are identified on the Figures. Hydocarbon charge
employed in obtaining these data was a wide cut Mid-Continent gas
oil (WCMCGO). It will be observed from these figures that catalyst
C identified particularly by the curve on the table and considered
as a low coke producing catalyst gave a much lower coke yield for
any given conversion level than the other catalyst compositions
after either steam or thermal treatment. Close examination of the
data shows that several of the catalyst compositions are much
higher coke producers after thermal treatment. The thermally
treated catalyst is designed to represent a freshly prepared
catalyst. The steam treated catalyst composition is designed to
represent equilibrium catalyst existing in a commercial operation.
It is to be particularly observed that catalyst compositions I, H,
B and F form substantially a straight line on FIG. 1 and are
representative of a higher coke making catalyst than that
experienced with catalyst C represented by the curve on the
figure.
The above observation is also made with respect to the data plotted
on FIG. 2. Thus, it is directionally clear from these informations
just how the coke on charge varies with conversion and for
different catalyst compositions. Furthermore, it is clear that
merely reciting that a cracking catalyst comprises a crystalline
zeolite as a cracking component is insufficient to identify its
activity and/or selectivity or the operating parameters in which it
is most desirably employed.
It is also evident from these data of the need for highly
sophisticated catalyst conversion systems such as that described
herewith which will permit one to take advantage of a particular
catalyst composition in an optimum manner. This does not mean to
say, however, that other catalyst compositions such as those
producing more coke than the low coke producing catalyst C of this
invention cannot be used with a considerably greater efficiency in
the method and system of this invention. On the contrary such
catalyst may also be employed with a high degree of sufficiency and
in many instances at a greater efficiency than permitted in now
available in fluid catalyst systems.
Referring now to FIGS. 3, 4 and 5 by way of example, there is shown
schematically arrangements of catalyst systems for practicing the
processing concepts herein described and particularly going to the
essence of this invention.
Referring now specifically to FIG. 3, there is shown
diagrammatically an arrangement of means for providing plural
stages of riser cracking, separation of the riser effluent into a
hydrocarbon phase and a catalyst phase, stripping of the separated
catalyst, passing one of the stripped catalyst phases to
regeneration, and using part or all of the other of these catalyst
phases in combination with freshly regenerated catalyst in one of
the risers provided for converting heavy hydrocarbons such as high
molecular weight hydrocarbons, residual hydrocarbons, coker gas
oils, or recycle feed materials either alone or combined with other
materials difficult to crack and producing high levels of coke. In
the arrangement of FIG. 3, a hydrocarbon feed such as a virgin
feed, fresh feed or fractions thereof is introduced to the
processing arrangement by conduit 2 to furnace 4 wherein the feed
is preheated to a desired elevated temperature within the range of
300.degree.F. to 800.degree.F. or higher. Preheating of the
hydrocarbon feed may occur in the presence of or absence of a
suitable diluent gasiform material described herein in such
quantities as to effect a desired reduction in the hydrocarbon
partial pressure. The diluent employed may be steam, dry
hydrocarbon gases or gasiform material comprising C.sub.4 +
hydrocarbons and particularly C.sub.5 + hydrocarbons. The amount of
diluent employed will, of course, vary with the type and boiling
range of the fresh feed charge employed and will be used in a
lesser amount when converting, for example, a low molecular weight
hydrocarbon feed as distinguished from a virgin feed material
boiling up to about 1,000.degree.F. or 1,200.degree.F. The
preheated feed with combined diluent at a desired elevated
temperature in conduit 6 is combined with freshly regenerated
catalyst withdrawn from the regenerator by provided standpipe 8 to
form a suspension having a catalyst to oil ratio selected from
within the range of 4 to 10/1 at a temperature of at least about
900.degree.F. and preferably at a temperature in the range of
1,000.degree.F. up to about 1,200.degree.F. The temperature
selected for accomplishing conversion in the riser will, of course,
depend upon the specific product desired from a given hydrocarbon
charge material. For example, a product comprising significant
amounts of olefins and/or isobutane may be had with very little
coke production particularly when employing high temperatures and
the low coke producing crystalline zeolite cracking catalyst such
as catalyst C discussed hereinbefore used alone or in combination
with the ZSM-5 type of catalyst. Operating conditions will be
employed which generally maximize the yield of gasoline boiling
range product. Under some conditions it may be desirable to produce
significant amounts of olefins suitable for use in alkylation
reactions to produce gasoline boiling range product in which case
the low coke producing catalyst C would be used alone in the
absence of ZSM-5 type of catalyst particles. In any event, the
suspension formed, as described above, is caused to move through
riser 10 at a velocity designed to provide a residence time therein
selected from within the range from 1 to 12 seconds and preferably
restricted to within the range of 4 to 8 seconds. Thus, when a very
short residence time in the riser is desired, for example, in the
range of 1 to 4 seconds, it is contemplated relying upon the
gasiform diluent material to form a suspension with freshly
regenerated catalyst which suspension is caused to flow through an
initial portion of the riser reactor before bringing the
hydrocarbon reactant material in contact therewith in a downstream
portion of the reactor. Thus the fresh feed riser may have a
plurality of spaced apart hydrocarbon feed inlets 7' and 7"
throughout the length thereof so as to facilitate varying the
residence time that the hydrocarbon reactant is in contact with the
catalyst suspension in the riser reactor. In any event the
hydrocarbon charge is converted in the riser for a predetermined
residence time before being separated from suspended catalyst by
discharge into a hopper 14 wherein the separation is facilitated by
one or more cyclone separator positioned in hopper 14. It is
preferred that separation of the suspension be initially completed
in one or more cyclone separators positioned within or external to
the catalyst collecting hopper to avoid undesired prolonged
cracking. Thus discharging the suspension into the dilute catalyst
phase of a hopper is not sufficient to obtain the separation
desired by this invention. It is also important to the method and
concept of this invention that the catalyst separated from the
suspension discharged from the riser be recovered at its highest
temperature so that the available heat of the catalyst can be
utilized to maximum advantage as hereinafter provided. Thus
quenching of the suspension or separated catalyst should be
avoided.
In the arrangement of FIG. 3, the suspension discharged from the
fresh feed riser is separated as by cyclone separator 11 with the
separated hydrocarbons being passed to a collection chamber 12 in
the upper portion of the catalyst hopper 14. The collected
hydrocarbon vapors are passed from chamber 12 by conduit 16 to
suitable fractionation equipment not shown but conventional in the
art.
The catalyst separated from the suspension is passed from the
cyclone by diplet 18 to a dense fluid bed of catalyst particles 20
therebelow. Stripping gas such as steam is introduced at an
elevated stripping temperature to the lower portion of the catalyst
bed 20 by conduit 22 so as to flow upwardly therethrough without
significant cooling thereof to remove entrained vaporous
hydrocarbons from the hot catalyst particles. Provisions are made
for passing stripped hydrocarbons and stripping gas through, for
example, a separate cyclone separator before the stripping gasiform
material passes into the collection chamber 12. Any entrained
catalyst fines separated in such a cyclone separator are returned
to the catalyst bed by a suitably provided dipleg. A substantially
vertical baffle 24 is provided in the hopper and extending upwardly
from the bottom thereof for keeping the catalyst separated from
riser 10 collected as a separate bed of catalyst in the lower
portion of the hopper. The upper portion of the hopper and above
the upper dense phase of the separated catalyst beds may be in open
communication with one another. The catalyst recovered as catalyst
bed 20 and being at a relatively high temperature not substantially
below the riser discharge temperature also contains a relatively
small amount of deposited carbonaceous material thereon and
therefore has a considerable amount of residual activity and heat
available for further use, as herein described. The catalyst in bed
20 is thus withdrawn from the lower portion thereof by standpipe
26. The catalyst withdrawn by standpipe 26 is mixed with hot
freshly regenerated catalyst withdrawn from regenerator 50 by
standpipe 28. The amount of freshly regenerated catalyst combined
with partially spent catalyst will vary depending upon the
temperature of the partially spent catalyst and the temperature
selected for converting hydrocarbon feed in the second riser. Thus
the temperature of the catalyst mixture will depend upon the
conversion temperature desired in the second riser reactor
generally above about 900.degree.F. for processing heavy charge
materials such as recycle hydrocarbons, coker gas oils and/or high
molecular weight charge materials either alone or in admixture with
one another. In any event the hot catalyst mixture comprising
partially used catalyst and freshly regenerated catalyst is at a
temperature sufficient to supply a major portion of the endothermic
heat requirements of the recycle conversion riser. In some
instances the amount of regenerated catalyst combined with
partially spent catalyst will be sufficient to make up for a
substantial portion of not all of the endothermic heat loss of the
partially spent catalyst and this heat makeup in combination with
the hydrocarbon charge preheat will be sufficient to effect the
elevated temperature conversion desired in the second riser reactor
herein referred to as the recycle riser reactor. Thus the second
riser reactor confines a moving suspension of catalysthydrocarbon
charge and diluent gasiform material having a temperature selected
from within the range of 950.degree.F. up to about 1050.degree.F.
and a catalyst to oil ratio in the range of 6 to 25/1. The
suspension is caused to move through the second riser reactor for a
predetermined residence time selected from within the range of 2 to
15 seconds. It is contemplated, as hereinbefore discussed with
respect to the fresh feed riser, of relying upon the gasiform
diluent material for conveying the mixed catalyst phase through an
initial portion of the recycle riser reactor so that the higher
molecular weight hydrocarbon charge may be introduced thereto at a
downstream portion of the riser as by inlets 37', 37" and 37'" so
as to provide contact times of about 2 seconds and higher. Thus the
plurality of hydrocarbon feed inlets spaced along the length to the
recycle riser facilitate obtaining very low hydrocarbon residence
time therein. Generally the residence time of the
hydrocarbon/catalyst suspension in the recycle riser will be
selected from within the range of 2 to 15 seconds with residence
times of 4 seconds or more at the higher catalyst to oil ratios
generally preferred. Thus it is contemplated effecting conversion
of a light gas oil fraction at a longer time period than that
relied upon to convert a heavy gas oil fraction.
In the specific arrangement of FIG. 3, the recycle feed is charged
by conduit 32 to furnace 34 wherein it is heated to an elevated
cracking temperature either in the presence of or absence of
diluent gasiform material. The amount of gasiform diluent material
employed will depend, as hereinbefore discussed, upon the
hydrocarbon charge employed and the extent to which it is desired
to lower the partial pressure of the hydrocarbon charge to be
converted. In any event cracking of a recycle hydrocarbon charge is
enhanced by relatively low hydrocarbon partial pressures and
sufficient diluents should be employed to optimize this effect. The
charge preheated to an elevated temperature and combined with the
catalyst mixture as above discussed is then passed through recycle
riser reactor 36 at a temperature selected from within the range of
900.degree.F. up to about 1,200.degree.F. The suspension undergoes
conversion of the hydrocarbon charge during passage through the
riser reactor before being separated as by discharge into hopper 14
containing cyclone separator equipment or by direct discharge into
a cyclone separator 38 attached to the end of the riser and located
within hopper 14. The suspension in riser 36 is separated in a
specific example in cyclone separator 38 into a hydrocarbon vapor
phase and a catalyst particle phase. The hydrocarbon vapors are
collected in chamber 12 and thence passed to a fractionator not
shown. The separated catalyst is passed by a suitable dipleg to a
dense fluid catalyst bed 40. The dense fluid bed of catalyst
particles 40 are stripped of entrained hydrocarbon vapors by steam
or other suitable stripping medium introduced to the lower portion
of the bed by conduit 42. The stripped catalyst comprising
deposited carbonaceous material is removed by standpipe 44 for
passage to catalyst regeneration.
In the arrangement of FIG. 3, the catalyst in standpipe 44 is
combined with regeneration gas such as air or other suitable oxygen
containing regeneration gas mixture introduced by conduit 46 to
form a suspension which is then passed through the riser
regenerator 48 for discharge in the upper dispersed phase of
regeneration vessel 50. This riser regenerator may also discharge
in the upper or intermediate portion of the dense bed of catalyst
in the regenerator. The catalyst in riser regenerator 48 undergoes
partial regeneration therein and upon discharge, for example, into
the dilute phase is separated as by settling and becomes a part of
the dense bed of catalyst therebelow. It is important to recover
all available heat generated in the catalyst phases and flue gases.
Regeneration gas such as air or oxygen supplemented gasiform
material is introduced to the bottom or lower portion of catalyst
bed 52 by inlet means 56. Gaseous products of combustion or flue
gases pass through one or more catalyst cycle separator not shown
in the upper portion of the regenerator where entrained catalyst
fines are separated from the flue gas before the flue gases are
removed from the upper portion of the regenerator by conduit
54.
FIGS. 4 and 5 differ from the processing arrangement of FIG. 3
disussed above primarily in the catalyst hopper arrangement or
design housing the cyclone separators at the discharge of each
riser and relied upon for cascading catalyst particles separated
from each riser effluent to the separate catalyst streams removed
therefrom.
Specifically FIG. 4 departs from FIG. 3 by providing a vertical
baffle 24 similar to that described in FIG. 3 and containing
catalyst flow through slots 60. In this arrangement transfer of
catalysts between that discharged from riser 10 and forming
catalyst bed 20 and that discharged from riser 36 forming catalyst
bed 40 is provided so as to provide a further control on the
catalyst/oil ratio particularly desired in riser 36. For example,
in an operation where higher temperatures of the catalyst passed to
the regenerator are desired the catalyst discharged from fresh feed
riser 10 may be caused to flow from bed 20 through the slots in the
vertical partition and be combined with the more contaminated
catalyst in bed 40 discharged from riser 36 and recycled thereafter
to the regenerator. Of course the reverse flow of catalyst through
slots 60 may be had by changing the upper bed level of catalyst bed
40 discharged from riser 36 to be above the upper bed level of
catalyst 20 on the other side of baffle 24 and discharged from
riser 10. This latter catalyst flow would be used in the case where
a high catalyst to oil ratio is desired in riser 36 in addition to
increasing the level of carbonaceous material on the catalyst
eventually passed to the regeneration zone.
In the arrangement of FIG. 4, catalyst withdrawn by standpipe 26 is
normally combined with freshly regenerated catalyst and the thus
formed catalyst mixture is employed in riser 36 similarly to that
discussed with respect to FIGS. 3 and 4. In addition, catalyst
withdrawn by standpipe 44 is passed to the regenerator in the
manner similar to that discussed with respect to FIGS. 3 and 4. The
variation in operation of FIG. 4 over that of FIG. 3 resides
primarily in providing catalyst flow between the separated catalyst
beds 20 and 40 respectively and as desired.
FIG. 5, on the other hand, is a further variation on the concept of
FIG. 4 in that the catalyst discharged from each riser is
accumulated as a common dense fluid bed of catalyst which is
stripped in the lower portion thereof by stripping gas in a
relatively confined stripping section provided with alternately
staggered baffling to cause intimate contact between catalyst
particles and stripping medium.
In the arrangement of FIG. 5, catalyst is withdrawn by standpipe 44
for transfer to the regenerator as discussed above. In this
arrangement catalyst withdrawal standpipe 26 is extended upwardly
into the dense fluid bed of catalyst particles to a catalyst
withdrawal cone 62 positioned above the relatively confined
catalyst stripping section. Stripping steam is introduced to the
lower portion of the bed of catalyst and will effect considerably
stripping of the catalyst in the upper portion of the bed before it
enters withdrawal cone 62. It is contemplated adding in this
arrangement hydrocarbon material such as recycle oil, slurry oil or
other high mol weight hydrocarbon material to be partially
converted at a point substantially intermediate the withdrawal cone
and an upper portion of the catalyst stripping section. In this
embodiment additional carbonaceous material for heat balance
purposes will be laid down on the catalyst and prior to its return
to the regenerator.
The arrangement of FIG. 5 may be further modified by withdrawing
all catalysts from beneath the stripping gas inlet whether it is
returned directly to the regenerator as by standpipe 44 or passed
to riser 36 as a mixture of catalyst particles with freshly
regenerated catalyst. The amount of carbonaceous material deposited
on the catalyst may be varied considerably.
In yet a further embodiment, hopper 14 of FIG. 5 may be used in a
manner resembling an elutriator wherein catalyst particles of
reduced density by virtue of carbonaceous deposits therein will be
caused to locate in an upper portion of the dense fluid catalyst
bed with the more dense catalyst particles because of a small
amount of carbonaceous deposits being located in a lower portion of
the dense fluid catalyst bed. Thus the more dense particles will be
withdrawn from the lower portion of the catalyst bed for passage to
riser 36 in admixture with freshly regenerated catalyst and the
catalyst particles containing a higher level of carbonaceous
material being withdrawn from an upper portion of the catalyst bed
for return to the regenerator as by withdrawal standpipe 44.
It is clearly evident from the discussions hereinbefore presented
that the processing schemes discussed and variations thereto are
directed to and designed particularly for the purpose of maximizing
utilization of the activity and selectivity characteristics of low
coke forming crystalline zeolite catalyst compositions in the
conversion of hydrocarbon charge materials. Furthermore the
combinations improved upon the heat balance of a coke deficient
operation in a manner which will minimize extraneous heat
requirements of the process, for example, minimize the preheat
requirements for the hydrocarbon feed and regeneration gas. Thus in
any of the embodiments discussed hereinbefore it is contemplated
contacting the catalyst to be passed to regeneration either before
or after final stripping thereof with a heavy hydrocarbon fraction
such as a residual oil fraction which will operate to lay down a
further increment of carbonaceous material on the catalyst for the
purpose of improving the overall heat balance of the processing
arrangement. For example, slurry oil containing catalyst fines and
recovered from the bottom of the fractionator may be combined with
the catalyst to be passed to regeneration as a means for recovering
catalyst fines and providing heat producing carbonaceous material
on the catalyst particles.
Regeneration of the catalyst particles may be accomplished as
briefly discussed hereinbefore in any one of several different
arrangements of carbonaceous material burning sequences with oxygen
containing gas which will be effective in removing contaminating
carbonaceous deposits from the catalyst particles to restore
substantially, if not completely, the activity and selectivity of
the catalyst particles. Thus the catalyst may be regenerated in a
plurality of sequentially arranged dilute-phase catalyst bed
regeneration zones to which oxygen may be added in separate
increments to avoid overheating of the catalyst during regeneration
of the catalyst at a temperature selected within the range of from
about 1000.degree.F. to as high as 1400.degree.F. On the other
hand, the dilute phase regeneration zones may be employed in
combination with a more dense phase regeneration step to accomplish
the same purpose. In any event the regeneration sequence selected
should be one which will recover the major portion of heat
available to burning of carbonaceous material and its ultimate use
in the conversion processing steps of the arrangements discussed
above.
Having thus provided a general discussion of the improved method
and concepts of this invention and described specific examples in
support thereof going to the essence of the invention, it is to be
understood that no undue restrictions are to be imposed by reason
thereof except as defined in the following claims.
* * * * *