U.S. patent number 3,617,497 [Application Number 04/836,404] was granted by the patent office on 1971-11-02 for fluid catalytic cracking process with a segregated feed charged to the reactor.
This patent grant is currently assigned to Gulf Research & Development Company. Invention is credited to Millard C. Bryson, James R. Murphy.
United States Patent |
3,617,497 |
Bryson , et al. |
November 2, 1971 |
**Please see images for:
( Certificate of Correction ) ** |
FLUID CATALYTIC CRACKING PROCESS WITH A SEGREGATED FEED CHARGED TO
THE REACTOR
Abstract
A hydrocarbon is cracked in the presence of a fluid zeolite
catalyst or a catalyst of comparable activity which produces a
transient maximum gasoline yield at a residence time of 5 seconds
of less and in the presence of a diluent vapor or vapors which
lower the partial pressure of the hydrocarbon feed or feeds and
increase gasoline selectivity. Residence time is established by
controlling the total charge rate of hydrocarbons and diluent
vapors. The ratio of diluent vapors to hydrocarbon feed is also
controlled so that a greater yield of gasoline is recovered from
the process than could be recovered in the absence of the diluent
vapor or Vapors. Gasoline yield is further enhanced by segregating
the hydrocarbon feed and charging the relatively lower molecular
weight feed fraction or fractions near the bottom of an elongated
riser or transfer line rector and the relatively higher molecular
weight feed fraction or fractions progressively further up the
riser or transfer line.
Inventors: |
Bryson; Millard C. (Conway,
PA), Murphy; James R. (Huntington Station, NY) |
Assignee: |
Gulf Research & Development
Company (Pittsburgh, PA)
|
Family
ID: |
25271898 |
Appl.
No.: |
04/836,404 |
Filed: |
June 25, 1969 |
Current U.S.
Class: |
208/80; 208/128;
208/157; 422/211; 208/120.01; 208/74; 208/130; 208/160 |
Current CPC
Class: |
C10G
11/18 (20130101); Y02P 30/40 (20151101); Y02P
30/446 (20151101) |
Current International
Class: |
C10G
11/18 (20060101); C10G 11/00 (20060101); C01b
033/28 (); C10g 011/18 (); C10q 011/20 () |
Field of
Search: |
;208/80,49,128,130,153,160 ;23/288 |
References Cited
[Referenced By]
U.S. Patent Documents
Primary Examiner: Gantz; Delbert E.
Assistant Examiner: Schmitkons; G. E.
Claims
We claim:
1. In a process for cracking at least one relatively low molecular
weight gas oil hydrocarbon feed stream and at least one relatively
high molecular weight gas oil hydrocarbon feed stream to gasoline
in the presence of a fluid zeolite cracking catalyst the
improvement comprising charging the relatively low molecular weight
hydrocarbon feed stream to said process at a relatively upstream
position and charging the relatively high molecular weight feed
steam to said process at a relatively downstream position along the
reaction path, performing said process at a temperature between
900.degree. and 1,100.degree. F. and a residence time of less than
five seconds during which catalyst and hydrocarbon both flow
concurrently through the process under conditions such as to avoid
formation of a catalyst bed in the reaction flow stream, the
cracking of said low molecular weight hydrocarbon feed stream
performed in the presence of an added diluent vapor which reduces
the partial pressure of said low molecular weight hydrocarbon feed
and produces a net increase in debutanized gasoline yield in said
process, and recovering debutanized gasoline from said process in
an amount including said net increase.
2. The process of claim 1 wherein the reactor is enlarged near the
position of introduction of said high molecular weight feed stream
so that the linear velocity before and after the enlargement is
between about 25 and 75 feet per second and the total reactor
length to diameter ratio is above about 20.
3. The process of claim 1 wherein the catalyst to low molecular
weight hydrocarbon weight ratio is between about 4:1 to about
15:1.
4. The process of claim 1 wherein the density of the material at
the low molecular weight feed inlet is about 1 to 4.5 pounds per
cubic foot.
5. The process of claim 1 wherein said diluent vapor is steam and
is present in an amount between about 0.5 to 10 weight percent
based on the low molecular weight feed.
6. The process of claim 1 wherein said catalyst is charged to the
process at a temperature of at least about 1,240.degree. F.
7. The process of claim 1 wherein the diluent vapor is steam,
nitrogen, methane or ethylene.
8. The process of claim 1 wherein the effluent stream discharges
from the cracking reactor in a lateral direction.
9. The process of claim 1 wherein said low molecular weight feed
stream and said high molecular weight feed stream are fractions of
a common gas oil hydrocarbon stream.
10. The process of claim 1 wherein the pressure is about 5 to 50
pounds per square inch gauge.
11. The process of claim 1 wherein the space velocity based upon
all feed streams is at least about 100 weight of hydrocarbon feed
per hour per weight of catalyst.
Description
This invention relates to the cracking of a petroleum hydrocarbon
feed stock to gasoline in the presence of a highly active fluid
cracking catalyst such as a crystalline aluminosilicate zeolite or
a catalyst of comparable activity, or selectivity, or both.
Natural or synthetic zeolite aluminosilicate cracking catalysts
exhibit high activity in the cracking of hydrocarbon oils both in
terms of total conversion of feed stock and in terms of selectively
towards gasoline production. The present invention relates to a
method for improving the selectivity to gasoline production in
cracking processes utilizing a fluidized zeolitic cracking catalyst
or a catalyst of comparable activity and/or selectivity.
In fluid catalytic cracking operations it is generally advantageous
to operate the cracking reactor at pressures in the range of about
20 to 30 pounds per square inch gauge and it is undesirable in
terms of the integrated operation including catalyst regeneration
and power recovery from regenerator flue gases for reactor
pressures to fall significantly below this level. For example,
catalyst regeneration is generally favorably influenced by elevated
temperatures and pressures. Furthermore, in systems where
regenerator flue gas is utilized to drive a turbine to compress
combustion air to be supplied to the regenerator, it is important
to maintain an elevated pressure in the regenerator in order to
obtain efficient turbine operation. Since spent catalyst must flow
from the reactor zone to the regenerator, a correspondingly high
pressure is consequently required in the reactor in order to urge
catalyst towards the regenerator. However, as shown below,
relatively high reactor hydrocarbon feed pressures are less
favorable to gasoline selectivity in the cracking operation than
relatively low pressures.
In accordance with the present invention a method is presented for
advantageously improving operation of a reaction process employing
a zeolitic or similar fluidized cracking catalyst without lowering
the pressure in the reaction zone or catalyst disengaging or
stripping vessel. We have discovered that an unexpected advantage
occurs by charging a diluent gas to the inlet of the cracking
reaction zone to lower the partial pressure of the charge
hydrocarbon in the reaction zone without disturbing the total
pressure in the system. Any diluent which is a vapor or becomes a
vapor under the conditions of the reaction zone can be used. An
inert gas such as steam or nitrogen is a suitable diluent. A
mixture of gases can be employed. If the diluent is a hydrocarbon,
it should desirably have a boiling point below about 430.degree.
F., i.e. it should be a gasoline range hydrocarbon or lighter. If
it boils above the gasoline range it will itself be a portion of
the cracking feed. Recycle methane or ethylene could be employed.
We have found that a lower hydrocarbon feed partial pressure at any
given reaction zone total pressure produces the unexpected effect
of increasing the selectivity to gasoline production at a given
conversion level of fresh feed or, conversely, requiring a lower
conversion of total feed to produce a given gasoline yield.
Although it has been known that the use of an inert diluent such as
steam at the hydrocarbon feed zone accomplishes certain
advantageous effects in a fluid catalytic cracking operation such
as assisting in fluidization of catalyst, vaporization of liquid
feed, dispersal of catalyst into hydrocarbon feed, increasing
reaction rate, etc., the improvement in gasoline selectivity has
not heretofore been appreciated. We have further discovered that
the gasoline selectivity advantage is transient and is lost if the
cracking process is not terminated in a timely manner, as explained
below. Because of its transient nature the selectivity advantage
has heretofore been effectively masked.
It has previously been considered that the amount of steam to be
employed in a fluid catalytic cracking process should not be great
in order to avoid a reduction in residence time, and thereby a loss
in conversion. However, in accordance with the present invention
the amount of steam or other inert gas must be sufficient to
produce a significant reduction in partial pressure of the incoming
hydrocarbon capable of being cracked to gasoline. Although the
initial increments of partial pressure reduction exert a greater
effect upon gasoline selectivity than later increments, the greater
the amount of steam or other inert gas introduced relative to
hydrocarbon feed the greater will be the effect upon selectivity.
For example, 10 mol percent steam based on hydrocarbon charge will
reduce the partial pressure of the hydrocarbon charge 10 percent,
15 mol percent steam will reduce the partial pressure of the
hydrocarbon charge 15 percent, etc., and the greater the reduction
in partial pressure the greater the gasoline selectivity advantage
it is possible to achieve in accordance with this invention.
In accordance with the present invention it has further been
discovered that the selectivity advantage due to the presence of an
inert gas, which is not itself capable of being cracked to
gasoline, is most significant in the very early stages of the
cracking reaction, which is also the period in which most of the
cracking of fresh feed occurs. In fact, the curve of production of
cracked hydrocarbon vapors from fresh feed with time is exponential
with the greatest rate of cracking occurring at the outset of the
reaction so that the cracked vapors themselves quickly reduce the
partial pressure of the unreacted feed. However, by the time these
vapors are produced most of the cracking has been completed. The
extent of cracking of fresh hydrocarbon feed with a zeolite
catalyst is considerably greater in the first 0.1 second interval
in the reaction zone than in the second 0.1 second interval.
Similarly, the extent of cracking of fresh hydrocarbon charge is
considerably greater in the first 0.2 second interval in the
reaction zone than in the second 0.2 second interval. For example,
after the hydrocarbon feed has been in the reaction zone for about
0.1 second it is about 40 percent converted and after about 1.0
second conversion increases only to about 70 to 80 percent.
In control methods for fluid catalystic cracking operations
according to the prior art, a vapor such as steam was added to the
inlet of an elongated riser or reaction zone to assist dispersal of
catalyst into hydrocarbon. The amount of steam was not considered
particularly critical. Reactor residence time (space velocity) was
then adjusted to control gasoline yield in the reactor effluent. If
analysis of reactor effluent indicated an adjustment of the
residence time was required, the hydrocarbon flow rate was
adjusted. But no criticality was attached to the fact that this
adjustment varied the ratio of steam to hydrocarbon at the reaction
zone inlet. In accordance with the present invention reaction zone
residence time is established not only by establishing the total
charge rate including both hydrocarbon and steam but also by
establishing the ratio of steam to hydrocarbon in the charge in the
manner described below. We have now discovered and it is shown
below that control of the ratio of steam to hydrocarbon in the
charge and control of the total charge rate including both steam
and hydrocarbon are interdependent and interdependently exert a
critical effect on gasoline yield.
Although zeolitic aluminosilicates are especially useful catalysts
for purposes of the present invention, any silica alumina or other
cracking catalyst which is sufficiently active and/or selective to
be capable of producing a transient maximum or peak gasoline yield
from the total fresh hydrocarbon feed capable of being cracked to
gasoline at residence times of 5 seconds or less are within the
purview of this invention. The maximum gasoline yield obtained at
residence times within 5 seconds is transient and rapidly
diminishes. After a residence time of 1 second, most of the fresh
hydrocarbon feed is converted and there is a sharp drop in rate of
conversion of fresh feed. However, if the hydrocarbon continues to
remain in contact with the catalyst, products of the earlier
cracking operation themselves in turn undergo cracking. This
occurrence is termed "aftercracking." Since there is a greater
abundance of cracked material than uncracked material after only
about one-half to 1 second of reaction zone residence time or less
the situation rapidly arises wherein considerably more cracking of
cracked than uncracked material can occur. When this situation
prevails, the desired gasoline product initially produced at a high
selectivity in accordance with the present invention becomes
depleted due to aftercracking at a faster rate than it is
replenished due to cracking of remaining uncracked feed so that the
selectivity advantage initially achieved is subsequently lost at a
significant rate. If timely disengagement of hydrocarbon and
catalyst does not occur prior to the occurrence of a significant
amount of aftercracking the very existence of the earlier
advantageous selectivity effect can be entirely masked. This
invention requires substantially instantaneous disengagement of
catalyst and hydrocarbon as these materials exit from the reaction
zone into a disengaging vessel.
In accordance with the present invention a preheated liquid
hydrocarbon charge and a fluid zeolite or comparable cracking
catalyst is added to a cracking reaction zone together with an
inert gaseous diluent such as steam, nitrogen, recycle methane or
ethylene, etc. The liquid hydrocarbon charge is substantially
instantaneously vaporized and the quantity of inert diluent is
sufficient to accomplish a substantial reduction in the partial
pressure of the hydrocarbon charge. The selectivity to gasoline
production is enhanced due to the lower hydrocarbon partial
pressure at the onset of cracking of the fresh feed due to the
presence of the diluent. In order not to subsequently lose the
selectivity advantage the hydrocarbon is permitted to remain in the
presence of the catalyst only as long as further conversion of
uncracked hydrocarbon produces a significant increase in gasoline
yield. The system is controlled so that substantially at the time
when further conversion of uncracked hydrocarbon produces no
significant net increase in gasoline yield or at the time when some
decrease in gasoline yield ensues the catalyst and hydrocarbon are
substantially instantaneously disengaged from each other to prevent
aftercracking of gasoline product from destroying the selectivity
advantage initially achieved due to the diluent partial pressure
effect. Analysis of the product to measure total conversion of
fresh feed or gasoline yield or both will aid in controlling the
reactor in accordance with this invention. These analyses will
provide a measure of gasoline selectivity for controlling the
reactor. Reaction time duration can be adjusted by regulation of
total feed rate, including hydrocarbon and steam, where the reactor
height is fixed.
In accordance with this invention, the reactor is operated so that
there is a continual increase in gasoline throughout substantially
the entire length of the reactor coupled with a decrease in fresh
feed, which means that the reaction is terminated at or near the
time of maximum gasoline yield. There is a substantial absence of
backmixing in the reactor since this would be conducive to
aftercracking. Backmixing can be caused by an excessive linear
velocity which gives rise to turbulence or by the formation of a
dense catalyst bed which induces turbulence in flowing vapors. The
hydrocarbon remains in the reactor only until a decrease in fresh
feed content is not accompanied by any substantial further net
increase in gasoline. Maximum gasoline yield is accompanied by
maximum gasoline selectivity.
The overall time of contact between hydrocarbon and catalyst can be
as low as about 0.5 second or less but not greater than about 5
seconds and will depend upon many variables in a particular process
such as the boiling range of the charge, the particular catalyst,
the amount of carbon on the regenerated catalyst, the catalyst
activity, the reaction zone temperature, the polynuclear aromatic
content of the hydrocarbon feed, etc. Some of these variables can
affect one another. For example, if the fresh hydrocarbon charge
includes a considerable quantity of polynuclear aromatics, the
reaction should be permitted to proceed long enough to crack any
mono- or di-aromatics or naphthenes because these compounds produce
relatively high gasoline yields and are the most readily crackable
aromatics but the reaction should be terminated before significant
cracking of other polynuclear aromatics occurs because cracking of
these latter compounds occurs at a slower rate and results in
excessive deposition of carbon on the catalyst. It is clear, that
no fixed cracking time duration can be set forth but the time will
have to be chosen with the range of this invention depending upon
the particular system. In one system, slightly exceeding a 1.0
second residence time might result in such severe aftercracking
that the selectivity advantage would be lost while in another
system unless a 1.0 second residence time is appreciably exceeded
there might not be sufficient cracking of charge hydrocarbon to
render the process economic. Generally, the residence time will not
exceed 2.5 or 3 seconds and 4 second residence times will be
rare.
Reference to FIG. 1 will illustrate the significance of the present
invention.
FIG. 1 contains curves semiquantitatively relating the amount of
unreacted charge and gasoline, as percent based on fresh feed, to
reaction zone residence time. The curve of unreacted charge which
is typical of most fluid cracking charge stocks shows that the
amount of unreacted charge asymptotically approaches a value
somewhat less than 20 percent of fresh feed within residence times
of this invention. The curves showing quantity of gasoline produced
show that the quantity of gasoline produced rapidly reaches a
somewhat flat maximum or peak which generally coincides with the
time at which the cracking of unreacted charge is substantially
diminished. The gasoline yield at the peak for a given feed will be
determined mostly by reactor temperature, to an extent by the level
of carbon on the catalyst and to an extent by the catalyst to oil
ratio. After reaching a peak, the gasoline level diminishes because
the aftercracking of gasoline predominates over production of
gasoline from the unreacted feed. The lower of the two gasoline
curves shown in FIG. 1 indicates the level of gasoline in the
reaction zone assuming substantially no inert diluent such as steam
is introduced to the inlet zone of the reactor. The upper of the
two gasoline curves schematically shows the higher gasoline level
achieved by adding an inert diluent such as steam to the inlet of
the reaction zone which lowers the hydrocarbon feed partial
pressure and thereby increases selectivity to gasoline.
Assuming a fluid cracking process is operating with steam addition
and the gasoline yield is at point A shown in FIG. 1 where
significant aftercracking has occurred. In order to reduce the
extent of aftercracking it is decided to increase the charge rate
of hydrocarbon into the reaction zone, thereby reducing hydrocarbon
residence time. Residence time is usually adjusted by adjustment of
hydrocarbon charge rate rather than steam charge rate since for any
given percentage increase or decrease in charge rate of steam or
hydrocarbon, the effect upon reaction residence time will be much
greater in the case of the hydrocarbon adjustment because the total
amount of hydrocarbon charged is so much greater than the total
amount of steam charged. Due to the shorter residence time and
concomitant reduction in aftercracking a higher gasoline yield B is
achieved. However, because the hydrocarbon partial pressure at the
reaction zone inlet has been increased by an increase in
hydrocarbon flow rate, the point B is removed from the upper
gasoline curve in the direction of the lower gasoline curve and is
outside the cross-hatched zone which denotes the range of this
invention. The cross-hatched zone shown in FIG. 1 denotes the
transient elevated gasoline yields of this invention which can be
recovered by the use of an inert vapor but which could not be
recovered absent an inert vapor. On the other hand, if the same
decrease in hydrocarbon residence time were achieved by increasing
both hydrocarbon and steam flow rate in the same ratio so that the
partial pressure of hydrocarbon at the reaction zone inlet remained
unchanged at the new residence time, the new operating point would
be at B', instead of B, which is within the range of the present
invention. (Of course, if the same total flow rate were achieved by
increasing the ratio of steam to hydrocarbon the new operating
point would be above B' and the area covered by the cross-hatched
zone of this invention would be enlarged.) Now, if the hydrocarbon
charge rate is again increased to further reduce residence time,
the point C is reached which is further removed from the upper
gasoline curve in the direction of the lower gasoline curve than is
point B because the hydrocarbon partial pressure is further
increased in going from point B to point C. Again, because of the
increase in hydrocarbon partial pressure, point C is outside the
range of the invention. On the other hand, if the same residence
time indicated at point C is achieved by increasing the flow rate
of both steam and hydrocarbon, rather than hydrocarbon alone, so
that the hydrocarbon partial pressure at the new residence time is
the same as it was at point A, the point C' is achieved which is
within the range of this invention.
It is seen from FIG. 1, that operating points B and C represent
essentially similar gasoline conversion levels occurring at
different residence times apparently indicating that these points
lie close to a flat maximum gasoline yield. However, points B and C
lie outside the range of the present invention while operating
points B' and C', which are within the range of this invention, lie
at higher gasoline yield levels than points B and C, even through
points B and B' and points C and C' represent the same residence
times, respectively. Starting from point A, point B' is reached by
the method of lowering residence time via a change in both steam
flow rate and hydrocarbon flow rate while, also starting from point
A, point B is reached by the method of changing hydrocarbon flow
rate only to achieve the same residence time as point B'. Starting
from point B', point C' is reached by changing both steam flow rate
and hydrocarbon flow range to lower the residence time, while point
C is reached by the simpler method of changing hydrocarbon flow
rate only to achieve the same residence time as at point C'. It is
apparent that to achieve the gasoline selectivity advantage of the
present invention, the residence time and the apportioning of steam
and hydrocarbon flow rates to achieve said residence time are
interdependent and represent a critical combination for purposes of
process control.
While the partial pressure effect of this invention tends to
increase selectivity to gasoline, there is a competing effect in a
cracking process which tends to oppose and thereby mask the partial
pressure effect. This competing effect arises due to carbon laydown
on the catalyst as the catalyst travels through the reaction zone.
As the amount of carbon on the catalyst increases along the
reaction path the gasoline selectivity from the feed decreases. The
higher the molecular weight of the feed hydrocarbon the greater the
carbon on catalyst competing effect because the high molecular
weight components tend to contain more polynuclear aromatic
compounds which yield more coke on cracking than other compounds.
Of the aromatic compounds, the polynuclear compounds not only crack
at a slower rate but also have a much higher selectivity to C.sub.2
and lighter gases and coke, while the mono- and di-aromatics and
the alkyl side chains of naphthenes tend not only to crack at a
faster rate but also to exhibit a higher selectivity to gasoline.
Therefore, the heavier hydrocarbon feed components should be
subjected to a reduced residence time, such as only about 0.5 to
1.5 seconds, in order to limit the cracking thereof as much as
possible to paraffinic side chains and mono- and di-aromatics in
general.
In accordance with this invention, the feed hydrocarbon in
fractionated and a fraction containing the relatively lower
molecular weight components (predominantly paraffins, naphthenes
and mono- and di-aromatics) to be cracked is charged together with
catalyst to the bottom of an elongated reaction zone and permitted
to undergo substantial cracking before reaching the position in the
reaction path of entry of a fraction containing the relatively
higher molecular weight components (which contain more
predominantly the polynuclear aromatics). After the major portion
of the cracking of the lower molecular weight fraction has
occurred, the higher molecular weight fraction is introduced to the
reactor without additional catalyst. In this manner most of the
lighter hydrocarbon feed is cracked in the absence of the heavy
hydrocarbon feed and thus on a low carbon content catalyst. The
cracking operation for the lower molecular weight feed is optimized
(maximum gasoline selectivity) under the combined influence of the
reduced partial pressure effect of the inert diluent, low carbon on
catalyst effect, and a somewhat more severe cracking operation
(i.e. high catalyst to oil ratio). The heavy hydrocarbon feed is
then subjected to a much shorter residence time than the lighter
feed. If desired, one or more relatively light hydrocarbon feed
streams can be introduced near the bottom of the reactor and one or
more relatively heavy hydrocarbon feed streams derived from the
same or a different source than that from which the light feed is
derived can be introduced relatively downstream along the reaction
path, the heavier the feed (or more polynuclear aromatic) the
further downstream its position of introduction. A heavy charge
stream can comprise recycle in whole or in part. At each position
of introduction of heavy feed, the diameter of the reactor can
increase so that the velocity at the inlet of the reactor and at
the outlet of the reactor will be about the same. If desired, the
reactor can be tapered to provide increasing diameters along the
reaction path to provide a uniform velocity throughout. A high
degree of control in the reactor is achieved by varying the amount
and position of introduction of the heavier hydrocarbon feed or
feeds relative to the amount and position of the lighter feed or
feeds in order to vary the residence time of all material flowing
through the reactor. In accordance with this invention, the
downstream position of introduction of the high molecular weight
hydrocarbon feed stream is established so that a greater percentage
yield of gasoline from the high molecular weight feed is recovered
from the process in the presence of the low molecular weight
reaction products stream (partial pressure effect) than could be
recovered in the absence of the low molecular weight reaction
products stream.
It is shown below that segregating the total hydrocarbon feed into
relatively high and low molecular weight fractions as described
provides an increased selectivity to gasoline as compared to
charging the full range hydrocarbon feed to a single position at
the bottom of the reactor. Data presented below indicates that a
heavy carbon laydown on the catalyst (such as is contributed by
heavier feeds) is a greater detriment to gasoline selectivity when
cracking a relatively low boiling feed than when cracking a
relatively high boiling feed, although it is a detriment with both.
Therefore, a net advantage in terms of gasoline selectivity is
achieved by permitting the low molecular weight feed to undergo
most of its cracking in the absence of the heavy feed and thus with
a catalyst having a low level of carbon. Thereupon, when the heavy
feed stream is introduced at a position downstream along the
reaction path, the reaction products of the lighter feed serve to
lower the partial pressure of the heavy feed stream to a great
extent, which in turn tends to offset the gasoline selectivity
disadvantage the heavy feed experiences due to being cracked in the
presence of a used and unregenerated carbon-containing catalyst. It
is seen that the carbon on catalyst effect and the vapor pressure
effect arising due to employing a segregated feed as described
constitute interdependent effects which cooperate to enhance
gasoline selectivity in the over-all process.
FIG. 2 illustrates the control method of this invention for a
reactor wherein a relatively low molecular weight hydrocarbon
fraction is added to the bottom of the reactor and a relatively
high molecular weight hydrocarbon fraction is charged to the
reactor at a position above the bottom of the reactor and
downstream along the reaction path from the position of entry of
the low molecular weight fraction. In an advantageous embodiment a
full range feed is fractionated to segregate it into two fractions
and the lower molecular weight fraction is charged to the bottom of
the reactor while the higher molecular weight fraction is charged
to a higher position in the reactor. The two fractions can be equal
or unequal in volume. Substantial cracking (but not the optimum) of
the low molecular weight fraction occurs in advance of the position
of charging of the high molecular weight fraction.
As shown in FIG. 2, the curved dashed lines indicate the gasoline
yield at varying residence times for the relatively light charge,
the lower curved dashed line indicating gasoline yield without
added vapor and the upper curved dashed line indicating gasoline
yield with added vapor. The enclosed unhatched region above the
horizontal dashed line M indicates the additional gasoline yield
achievable due to the use of a vapor with the light charge because
of the hydrocarbon partial pressure reduction at the reactor inlet.
The vertical line X indicates the maximum allowable residence time
if this additional gasoline yield is to be actually
recoverable.
The dotted lines of FIG. 2 indicate the addition of the heavy
hydrocarbon fraction at a position in the reactor which is so high
that the heavy feed does not have time to reach a maximum gasoline
yield before reaching residence time line X.
The curved solid lines of FIG. 2 indicate the addition of the heavy
hydrocarbon fraction at a position in the reactor which is above
the bottom of the reactor but which is sufficiently close to the
bottom that the heavy charge is in the reactor for a sufficiently
long time duration to achieve a maximum gasoline yield. The
unhatched enclosed area above the horizontal solid line N
represents the additional gasoline yield achievable from the heavy
charge due to the presence of the vapors from the reaction stream
derived from the light charge which lower the partial pressure of
the heavy hydrocarbon feed at the position of admission of said
heavy hydrocarbon. The vertical line Y demarcates the lowest
residence time permissible if this additional gasoline yield
obtainable from the heavy charge is to be actually recoverable.
It is seen from FIG. 2 that the residence time interval bracketed
by vertical lines X and Y is the only interval in which the
additional gasoline yield derived from both the light and heavy
charge due to vapor pressure reduction in the feed zone of each can
actually be recovered from the reactor. Therefore, the position
denoted by B' which was discussed above in regard to FIG. 1 lies
between lines X and Y and also lies above dashed horizontal line M
with which it is associated so that at this position the additional
gasoline yield obtainable from both the light and the heavy charge
can be recovered from the process. On the other hand, the position
denoted by C' which was also discussed above in regard to FIG. 1
lies outside the bracket established by the lines X and Y so that
while the additional gasoline yield from the light charge is
recoverable the additional gasoline yield from the heavy charge is
not recoverable. Therefore, the residence time corresponding to the
point C' in FIG. 2 is not suitable when charging a segregated feed
in accordance with the present invention.
It will be evident that only those gasoline yield points in FIG. 2
lying between lines X and Y which also lie above the dashed or
solid horizontal line M or N with which they are associated fall
within the purview of this invention. For example, the position B
in FIG. 2, which position was also discussed above in regard to
FIG. 1, although it falls between residence time lines X and Y it
does not fall above horizontal dashed line M with which it is
associated, and therefore lies outside the purview of this
invention. Therefore, the positions B and C which fall outside the
limits of this invention in accordance with the discussion
regarding the single hydrocarbon feed system of FIG. 1, remain
outside the confines of this invention according to the dual feed
system illustrated in FIG. 2.
In any particular process the gasoline yield and residence time
values which encompass the gasoline selectivity advantage of the
present invention will depend upon many variables peculiar to the
particular process. These variables include the particular catalyst
which is employed, the level of carbon on the regenerated catalyst,
catalyst activity and/or selectivity, the temperature, the
refractory characteristics of the feed, etc. The extent of this
selectivity advantage of this invention might be as low as one-half
percent to 1 percent or as high as 3 to 5 percent depending upon
the ratio of diluent vapor to hydrocarbon feed at the reactor inlet
and the apportionment of charges and their respective feed
locations. Where gasoline is the most economically desirable
product of the cracking operation, the economic value of a
selectivity advantage of even one-half or 1 percent actually
recovered as effluent is considerable in a commercial reactor unit
which processes 100,000 or 150,000 barrels per day of hydrocarbon
feed.
The reaction temperature in accordance with this invention is at
least about 900.degree. F. The upper limit can be about
1,100.degree. F., or more. The preferred temperature range is
950.degree. to 1,050.degree. F. The reaction total pressure can
vary widely and can be, for example, 5 to 50 p.s.i.g., or,
preferably, 20 to 30 p.s.i.g. The maximum residence time is 5
seconds, and for most charge stocks the residence time will be
about 1.5 seconds or 2.5 seconds or, less commonly, 3 or 4 seconds.
For high molecular weight charge stocks which are rich in aromatics
a 0.5- to 1.5-second residence time could be suitable in order to
crack mono- and di-aromatics and naphthenes which are the aromatics
which crack most easily and which produce the highest gasoline
yield, but to terminate the operation before appreciable cracking
of polyaromatics occurs because these materials produce high yields
of coke and C.sub.2 and lighter gases. The length to diameter ratio
of the reactor can vary widely, but the reactor should be elongated
to provide a high linear velocity, such as 25 to 75 feet per
second, and to this end a length to diameter ratio above 20 or 25
is suitable. The reactor can have a uniform diameter or can be
provided with a continuous taper or a stepwise increase in diameter
along the reaction path to maintain a nearly constant velocity
along the flow path. The amount of diluent can vary depending upon
the rate of hydrocarbon to diluent desired for control purposes. If
steam is the diluent employed, a typical amount to be charged can
be about 10 percent by volume, which is about 1 percent by weight,
based on hydrocarbon charge. A suitable but nonlimiting proportion
of diluent gas, such as steam or nitrogen, to fresh hydrocarbon
feed can be 0.5 to 15 percent by weight.
A zeolite catalyst is a highly suitable catalytic material in
accordance with this invention. A mixture of natural and synthetic
zeolites can be employed. Also a mixture of crystalline zeolitic
organosilicates with nonzeolitic amorphous silica aluminas is
suitable as a catalytic entity. Any catalyst containing zeolitic
material or otherwise which provides a transient maximum gasoline
yield within a 5-second residence time is suitable. The catalyst
particle size must render it capable of fluidization as a disperse
phase in the reactor. Typical and nonlimiting fluid catalyst
particle size characteristics are as follows:
Size (Microns) 0-20 20-45 45-75 >75
__________________________________________________________________________
Weight percent 0-5 20-30 35-55 20-40
These particle sizes are usual and are not peculiar to this
invention. A suitable weight ratio of catalyst to total oil charge
is about 4:1 to about 12:1 or 15:1 or even 25:1, generally, or 6:1
to 10:1, preferably. The fresh hydrocarbon feed is generally
preheated to a temperature of about 600.degree. to 700.degree. F.
but is generally not vaporized during preheat, and the additional
heat required to achieve the desired reactor temperature is
imparted by hot, regenerated catalyst.
The weight ratio of catalyst to hydrocarbon in the feed is varied
to affect variations in reactor temperature. Furthermore, the
higher the temperature of the regenerated catalyst the less
catalyst is required to achieve a given reaction temperature.
Therefore, a high regenerated catalyst temperature will permit the
very low reactor density level set forth below and thereby help to
avoid backmixing in the reactor. Generally, catalyst regeneration
can occur at an elevated temperature of about 1,240.degree. F. or
1,250.degree. F. or more to reduce the level of carbon on the
regenerated catalyst from about 0.6 to 1.5 to about 0.05 to 0.3
percent by weight. At usual catalyst to oil ratios in the feed, the
quantity of catalyst is more than ample to achieve the desired
catalytic effect and therefore if the temperature of the catalyst
is high, the ratio can be safely decreased without impairing
conversion. Since zeolitic catalysts are particularly sensitive to
the carbon level on the catalyst, regeneration advantageously
occurs at elevated temperatures in order to lower the carbon level
on the catalyst to the stated range or lower. Moreover, since a
prime function of the catalyst is to contribute heat to the
reactor, for any given desired reactor temperature the higher the
temperature of the catalyst charge the less catalyst is required.
The lower the catalyst charge rate the lower the density of the
material in the reactor. As sated, low reactor densities help to
avoid backmixing.
The reactor linear velocity, while not being so high that it
induces turbulence and excessive backmixing, must be sufficiently
high that substantially no catalyst accumulation or buildup occurs
in the reactor because such accumulation itself leads to
backmixing. (Therefore, the catalyst to oil weight ratio at any
position throughout the reactor is about the same as the catalyst
to oil weight ratio in the charge.) Stated another way, catalyst
and hydrocarbon at any linear position along the reaction path both
flow concurrently at about the same linear velocity, thereby
avoiding significant slippage of catalyst relative to hydrocarbon.
A buildup of catalyst in the reactor leads to a dense bed and
backmixing which in turn increases the residence time in the
reactor for at least a portion of the charge hydrocarbon and
induces aftercracking. Avoiding a catalyst buildup in the reactor
results in a very low catalyst inventory in the reactor, which in
turn results in a high space velocity. Therefore, a space velocity
of over 100 or 120 weight of hydrocarbon per hour per weight of
catalyst inventory is highly desirable. The space velocity should
not be below 35 and can be as high as 500. Due to the low catalyst
inventory and low charge ratio of catalyst to hydrocarbon, the
density of the material at the inlet of the reactor in the zone
where the low molecular weight feed is charged can be only about 1
to less than 5 pounds per cubic foot, although these ranges are
nonlimiting. An inlet density in the zone where the low molecular
weight feed and catalyst is charged below 4 or 4.5 pounds per cubic
foot is desirable since this density range is too low to encompass
dense bed systems which induce backmixing. Although, conversion
falls off with a decrease in inlet density to very low levels, we
have found the extent of aftercracking to be a more limiting
feature than total conversion of fresh feed, even at an inlet
density of less than 4 pounds per cubic foot. At the outlet of the
reactor the density will be about half of the density at the inlet
because the cracking operation produces about a fourfold increase
in mols of hydrocarbon. The decrease in density through the reactor
can be a measure of conversion.
A wide variety of hydrocarbon oil charge stocks can be employed. A
suitable charge is a gas oil boiling in the range of 430.degree. to
1,100.degree. F. As much as 5 to 20 percent of the fresh charge can
boil above this range. Some residual oil can be charged. A 0 to 5
percent recycle rate can be employed. Generally, the recycle will
comprise 650.degree. F. + oil from the product distillation zone
which contains catalyst slurry. If there is no catalyst
entrainment, recycle can be omitted.
EXAMPLE 1
A series of tests were conducted which illustrate the effect of
reducing hydrocarbon partial pressure upon selectivity to
debutanized gasoline and to C.sub. 3 + liquid yield. The tests were
conducted in an elongated reactor and the hydrocarbon partial
pressure was reduced by addition of steam and nitrogen with the
feed hydrocarbon. The ranges of conditions of the various tests
were as follows:
Charge Stock Inspections
__________________________________________________________________________
Gravity: .degree.API 25.6 Sulfur: Weight Percent 0.8 Ramsb. Carbon
Residue: Weight Percent 0.42 Vac. Distillation (corres. to 760 mm.
Hg) .degree. F. at 10 % 580 30% 692 50% 767 70% 847 90% 969 C.sub.A
(percent of aromatic atoms) 0.18 Catalyst Zeolite (50-60 Kellogg 2
Hour Activity)
Cracking Conditions Temperature: .degree.F. 950 Contact Time:
Seconds 0.1-2.0 Cat-to-Oil Ratio 6.5-9.0 Recycle none Riser Total
Pressure: p.s.i.g. 23-30 Riser Gas Composition (Inlet): Mol Percent
Hydrocarbon 5-80 Steam 5-90 Nitrogen 2-31
The results of the tests are illustrated in FIG. 3 in which
debutanized gasoline yield and total C.sub. 3 +liquid yield, both
reported as percent by volume of fresh feed, are plotted against
total conversion at various partial pressures of hydrocarbon in the
system and at various residence times. The pressure ranges given on
the face of the graphs indicate the partial pressure in the system
of all hydrocarbon vapors, cracked and uncracked, with the
remainder of the reactor pressure accounted for by nitrogen and
steam, both nitrogen and steam being used in all tests. For each
partial pressure, conversion data is indicated for one or more
residence times.
As shown in FIG. 3, at any given conversion level the selectivity
to gasoline as well as to total C.sub. 3 +liquid increases with
decreasing hydrocarbon partial pressure. Taking a 60 percent
conversion level for purposes of example, when the hydrocarbon
partial pressure is 16-20 p.s.i.g., the gasoline yield is 47.5
percent; when the hydrocarbon partial pressure is 10-14 p.s.i.g.
the gasoline yield increases to almost 50 percent; and when the
hydrocarbon partial pressure is 2-5 p.s.i.g. the gasoline yield
increases still further to about 51.5 percent. Advantageously, a
greater improvement in gasoline selectivity occurred in reducing
hydrocarbon partial pressure from 16-20 p.s.i.g. to 10-14 p.s.i.g.
than occurred in reducing hydrocarbon partial pressure from 10-14
p.s.i.g. to the very low partial pressure level of 2-5 p.s.i.g.
This shows that the gasoline selectivity advantage of this
invention was realized to a very significant extent in the initial
partial pressure reduction step of the tests and the effect was not
as great but still substantial in the second partial pressure
reduction step of the tests.
EXAMPLE 2
Tests were conducted to illustrate the advantage of a crystalline
zeolite aluminosilicate catalyst over an amorphous silica-alumina
catalyst in a fluid catalytic cracking system. Both catalysts were
tested under sufficiently low space velocity conditions that a
dense phase bed formed in the reactor. The results are shown in
table 1.
---------------------------------------------------------------------------
TABLE
1 Charge Stock
__________________________________________________________________________
Characterization Factor 12.09 11.95 Gravity: .degree.API 29.7 29.4
Sulfur: Percent 0.42 0.36 Viscosity, SUS at: .degree. F. 130 60.3
-- 150 51.1 -- 210 38.6 37.3 Carbon Residue, Ramsbottom: Percent
ASTM D524 0.23 0.21 Aniline Point; .degree. F. 188 184 Bromine
Number, D1159 2.8 3.0 Pour Point, D97: .degree. F. 90 -- Nitrogen:
p.p.m. 710 450 Metals: p.p.m. Vanadium 0.2 0.4 Nickel 0.2 0.1
Distillation Vac. (Corres. to 760 mm. Hg) 10% over at: .degree. F.
568 556 30 659 622 50 744 699 70 845 809 90 979 939 95 -- 991
catalyst 100 percent 60 percent Amorphous silica- zeolite, 40
alumina percent silica- alumina Kellogg Activity (2-Hour) 33.8
50.6
Operating Conditions: Reactor
__________________________________________________________________________
Fresh Feed Rate: B/D 13,571 13,704 Reactor Bed Temperature:
.degree. F. 926 935 Feed Preheat Temperature: .degree.F. 700 649
Reactor Bed Pressure: p.s.i.g. 11.5 11.0 Space Velocity, (Total
Feed): Wt./Hr./Wt. 3.94 3.07 Catalyst to Oil Ratio (Total Feed):
Wt./Wt. 12.5 9.8 Recycle: Percent by Volume of Fresh Feed 74.3 31.4
Carbon on Regenerated Cat. % by Wt. 0.4 0.38 Conversion: % by
Volume of Fresh Feed 75.5 85.5
Operation Conditions: Regenerator
__________________________________________________________________________
Regen. Bed Temperature: .degree. F. 1,141 1,166 Total Regen. Air:
Mlb./Hr. 153.7 166.72 Lb. Coke Burned/Lb. Air: Wt./Wt 0.087
0.083
Yields: % by Volume of Fresh Feed
__________________________________________________________________________
Debutanized Gasoline 47.5 61.0 Butane-Butene 21.2 21.6 i-Butane 7.6
10.3 n-Butane 2.1 1.7 Butenes 11.6 9.6 Propylene Propane 4.2 5.7
Propylene 8.5 5.9 Total Liquid Recovery 105.9 108.7 C.sub.2 and
Lighter Gas: % by Wt. 4.4 2.9 Coke: % by Wt. 7.73 7.8
Inspections
__________________________________________________________________________
Motor, Clear -- 81.3 Motor, +3 cc. TEL 86.1 89.4 Research, Clear
94.0 93.4 Research, +3 cc. TEL 100.4 98.3
__________________________________________________________________________
as shown in table 1, the zeolite catalyst system exhibited a
conversion of 85.5 percent compared to only 75.5 percent for the
amorphous catalyst. In addition, the zeolite catalyst system
exhibited a 61.0 percent yield of gasoline compared to only 47.5
percent gasoline yield with the amorphous catalyst. However, while
the total yield of C.sub.3 and C.sub.4 hydrocarbons is about the
same for the zeolite and the amorphous catalyst, the proportion of
these C.sub.3 and C.sub.4 hydrocarbons which are olefinic is lower
when utilizing a zeolite catalyst in these tests. This is a
disadvantage arising when utilizing a zeolite catalyst with
extended residence times in a dense catalyst bed because C.sub.3
and C.sub.4 olefins are useful for the production of alkylate which
can be blended with the gasoline produced directly by cracking to
improve its octane value.
EXAMPLE 3
Further tests were conducted to illustrate the use of the same type
of zeolite catalyst employed in example 2 for fluid catalytic
cracking not only at relatively high residence times involving
space velocities low enough to permit a dense phase catalyst bed to
form in the reactor but also at very low residence times within the
range of this invention at which the velocity through the reactor
is sufficiently high that no bed formation within the reactor and
therefore no backmixing due to bed formation is permitted to occur.
The results are shown in table 2. ##SPC1##
A comparison of tests 1 and 2 of table 2, both conducted at
950.degree. F., shows the deleterious effect of extended residence
time when employing a zeolite catalyst. The residence time of test
2 was only 0.5 second and yet it exhibited a higher gasoline yield
and a lower C.sub.2 and lighter yield than test 1 in which the
residence time was considerably longer due to a lower space
velocity and backmixing arising in the dense catalyst bed. A
comparison of tests 1 and 2 shows that an extended residence time
gives arise to aftercracking which diminishes gasoline yield and
increases the yield of products boiling lower than gasoline.
Comparing test 3 with test 1, both involving dense bed cracking, it
is seen that raising the cracking temperature from 950.degree. to
1,000.degree. F., provided a significant increase in conversion but
very little increase in debutanized gasoline yield and a higher
yield of C.sub.2 and lighter, showing that the high degree of
aftercracking occurring in a dense bed reaction system prevents
effective control of gasoline yield via temperature adjustment.
Comparing test 4 with test 2, both involving nonbed cracking and
very low residence times within the range of this invention, it is
seen that raising the cracking temperature from 950.degree. to
1,000.degree. F. provided not only a significant increase in
conversion but also an equally significant increase in gasoline
yield coupled with a lower yield of both C.sub.2 and lighter and
coke, showing that the comparative absence of aftercracking at the
very low residence times of this invention permits control of
gasoline yield via temperature regulation. It is also noted that
test 4 provided good yields of C.sub.3 olefin and C.sub.4 olefin
which are valuable materials for preparation of alkylate
gasoline.
Since table 2 indicates that in low residence time nondense bed
systems gasoline yield can be effectively controlled via
temperature regulation, it follows that a reduction in temperature
might be useful on occasion in an operating plant to reduce
gasoline yield as required by subsequent fractionator load or to
decrease C.sub.3 olefin and C.sub.4 olefin production. However, no
matter what the operating temperature is the gasoline yield at that
temperature is increased by utilizing the control method of this
invention.
EXAMPLE 4
Table 3 shows the results of four tests including a test based upon
calculation which illustrate the advantageous effect on gasoline
yield achievable by fractionating a hydrocarbon cracking feed into
a relatively high molecular weight fraction and a relatively low
molecular weight fraction and separately cracking the fractions in
the presence of a zeolite catalyst. Test 1 of table 3 shows the
results where a full range hydrocarbon feed is charged to the
bottom of a single reactor. Test 3 shows the results where the
total feed is fractionated and the lighter 50 percent by volume is
alone charged to the bottom of a single reactor. Test 4 shows the
results where the heavier 50 percent by volume of the fresh feed is
alone charged to the bottom of a single reactor. Test 2 shows the
calculated results of an integrated process wherein a total
hydrocarbon feed is segregated so that the lighter 50 percent by
volume is charged to one reactor and the heavier 50 percent by
volume is charged to another reactor and the effluents of the two
reactors are combined. All tests were made at a sufficiently low
velocity that a dense fluid catalyst bed was formed. All the tests
were conducted at the same hydrocarbon partial pressure at the
reactor inlet. ##SPC2## ##SPC3##
Comparing test 3 and test 1 of table 3, it is seen that cracking
the light charge alone resulted in about the same conversion as was
obtained with a full range charge but at a significantly higher
gasoline yield, indicating higher gasoline selectivity.
Furthermore, the average carbon level on the catalyst in test 3 was
0.93 less 0.3, or only 0.63 percent, while the average carbon level
on the catalyst in test 1 was 1.22 less 0.3 or 0.92 percent. Again,
the total C.sub.2 and lighter plus coke yield in test 3 was only
9.1 percent while the total C.sub.2 and lighter plus cope yield in
test 1 was 12.6 percent. In all these respects the cracking of the
light fraction by itself is superior to the cracking of a full
range charge.
Opposite results are indicated by comparing test 4 with test 1,
whereby it is seen that cracking the heavy charge alone results in
a much higher conversion than was obtained with a full range charge
but at only a slightly higher gasoline yield, indicating much lower
gasoline selectively. Furthermore, the average carbon level on the
catalyst in test 4 was 1.55 less 0.3 or 1.25 percent, while the
average carbon level on the catalyst in test 1 was only 0.92
percent. Again, the total C.sub.2 and lighter plus coke yield in
test 3 was 16.9 percent while the total C.sub.2 and lighter plus
coke yield in test 1 was only 12.6 percent. In all these respects
the cracking of the heavy fraction by itself is inferior to the
cracking of a full range charge.
Now, comparing calculated test 2 with test 1, it is seen that the
combined effects of tests 3 and 4 discussed above result in an
integrated process which is favorable to gasoline selectivity in
that gasoline yield is increased from 58.8 to 59.8 percent of fresh
feed. Therefore, the segregation of the fresh field as described in
this test results in a higher gasoline yield and can cooperate with
the vapor pressure effect described above in increasing gasoline
yield with a given hydrocarbon fresh feed.
A suitable reactor-regenerator system for performing this invention
is described in reference to FIG. 4. The cracking occurs with a
fluidized zeolitic catalyst in an elongated reactor tube 10, which
is referred to as a riser. The riser has a length to diameter ratio
of above 20, or above 25. Hydrocarbon oil feed to be cracked in
line 2 is first fractionated in column 4 into a relatively low
molecular weight fraction which flows through line 6 and a
relatively high molecular weight fraction which flows through line
8. The low molecular weight fraction is passed through preheater 11
to heat it to about 600.degree. F. and then charged into the bottom
of the riser through inlet line 14. Steam is introduced into the
low molecular weight oil inlet line through line 18. Steam is also
introduced independently to the bottom of the riser through line 22
to help carry upwardly into the riser regenerated catalyst which
flows to the bottom of the riser through transfer line 26.
The high molecular weight hydrocarbon fraction is preheated to a
temperature of about 600.degree. F. in preheater 20 and is
introduced through line 24 into the upper section of the riser at
the zone wherein the diameter of the riser becomes enlarged. The
high molecular weight hydrocarbon charge is introduced at about a
45.degree. upward angle into the riser through lines 30 and 32.
Steam can be introduced into the high molecular weight hydrocarbon
inlet lines through lines 34 and 36. High molecular weight
hydrocarbon lines 30 and 32 each represent a plurality of similar
lines spaced circumferentially at the same height of the riser. Any
recycle hydrocarbon can be admitted to the upper section of the
riser through one of the upwardly inclined inlet lines designated
as 38. No catalyst is added directly to the upper section of a
riser but all of the catalyst is added at the bottom of the riser
together with the low molecular weight hydrocarbon feed. The
residence times of both the high molecular weight feed and the low
molecular weight feed can be varied by varying either the relative
amounts or positions of introduction of the high and low molecular
weight feed streams. Therefore, the high molecular weight feed
stream can be introduced, through line 30, or alternately through
higher or lower lines 30A or 30B, respectively.
The full range oil charge to be cracked in the riser is a gas oil
having a boiling range of about 430.degree. to 11,00.degree. F. As
indicated above, before being charged the gas oil is fractionated
into a low molecular weight fraction which is charged to the bottom
of the riser and a high molecular weight fraction which is charged
to the top of the riser. The steam added to the riser amounts to
about 10 weight percent based on the oil charge, but the amount of
steam can vary widely. The steam is added with both the low and
high molecular weight hydrocarbon fractions. The catalyst employed
is a fluidized zeolitic aluminosilicate and is added to the bottom
only of the riser. The riser temperature range is about 900.degree.
to 1,100.degree. F. and is controlled by measuring the temperature
of the product from the risers and then adjusting the opening of
valve 40 by means of temperature controller 42 which regulates the
inflow of hot regenerated catalyst to the bottom of the riser. The
temperature of the regenerator catalyst is above the control
temperature in the riser so that the incoming catalyst contributes
heat to the cracking reaction. The riser pressure is between about
10 and 35 p.s.i.g. Between about 0 and 5 percent of the oil charge
to the riser is normally recycled.
The residence time of both hydrocarbon and catalyst in the riser is
very small and ranges from 0.5 to 5 seconds. The lower molecular
weight hydrocarbon is usually in the riser for about two seconds
because it is introduced to the bottom of the riser but the higher
molecular weight hydrocarbon will generally be in the riser for no
more than about one second because it is introduced into the top of
the riser. The velocity throughout the riser is about 35 to 55 feet
per second and is sufficiently high so that there is little or no
slippage between the hydrocarbon and catalyst flowing through the
riser. Therefore, no bed of catalyst is permitted to build up
within the riser, whereby the density within the riser is very low.
The density within the riser is a maximum of about 4 pounds per
cubic foot at the bottom of the riser and decreases to about 2
pounds per cubic foot at the top of the riser. Since no dense bed
of catalyst is permitted to build up within the riser the space
velocity through the riser is usually high and will have a range
between 100 or 120 and 600 weight of hydrocarbon per hour per
instantaneous weight of catalyst in the reactor. No significant
catalyst build up within the reactor is permitted to occur and the
instantaneous catalyst inventory within the riser is due to a
flowing catalyst to oil weight ratio between about 4:1 and 15:1,
the weight ratio corresponding to the feed ratio.
The hydrocarbon and catalyst exiting from the top of each riser is
passed into a disengaging vessel 44. The top of the riser is capped
at 46 so that discharge occurs through lateral slots 50 for proper
dispersion. An instantaneous separation between hydrocarbon and
catalyst occurs in the disengaging vessel. The hydrocarbon which
separates from the catalyst is primarily gasoline together with
some heavier components and some lighter gaseous components. The
hydrocarbon effluent passes through cyclone system 54 to separate
catalyst fines contained therein and is discharged to a
fractionator through line 56. The catalyst separated from
hydrocarbon in disengager 44 immediately drops below the outlets of
the riser so that there is no catalyst level in the disengager but
only in a lower stripper section 58. Steam is introduced into
catalyst stripper section 58 through sparger 60 to remove any
entrained hydrocarbon in the catalyst.
Catalyst leaving stripper 58 passes through transfer line 62 to a
regenerator 64. This catalyst contains carbon deposits which tend
to lower its cracking activity and as much carbon as possible must
be burned from the surface of the catalyst. This burning is
accomplished by introduction to the regenerator through line 66 of
approximately the stoichiometrically required amount of air for
combustion of the carbon deposits. The catalyst from the stripper
enters the bottom section of the regenerator in a radial and
downward direction through transfer line 62. Flue gas leaving the
dense catalyst bed in regenerator 64 flows through cyclones 72
wherein catalyst fines are separated from flue gas permitting the
flue gas to leave the regenerator through line 74 and pass through
a turbine 76 before leaving for a waste heat boiler wherein any
carbon monoxide contained in the flue gas is burned to carbon
dioxide to accomplish heat recovery. Turbine 76 compresses
atmospheric air in air compressor 78 and this air is charged to the
bottom of the regenerator through line 66.
The temperature throughout the dense catalyst bed in the
regenerator is about 1,250.degree. F. The temperature of the flue
gas leaving the top of the catalyst bed in the regenerator can rise
due to afterburning of carbon monoxide to carbon dioxide.
Approximately a stoichiometric amount of oxygen is charged to the
regenerator and the reason for this is to minimize afterburning of
carbon monoxide to carbon dioxide above the catalyst bed to avoid
injury to the equipment since at the temperature of the regenerator
flue gas some afterburning does occur. In order to prevent
excessively high temperatures in the regenerator flue gas due to
afterburning, the temperature of the regenerator flue gas is
controlled by measuring the temperature of the flue gas entering
the cyclones and then venting some of the pressurized air otherwise
destined to be charged to the bottom of the regenerator through
vent line 80 in response to this measurement. The regenerator
reduces the carbon content of the catalyst from 1.+-.0.5 weight
percent to 0.2 weight percent, or less. If required, steam is
available through line 82 for cooling the regenerator. Makeup
catalyst is added to the bottom of the regenerator through line 84.
Hopper 86 is disposed at the bottom of the regenerator for
receiving regenerated catalyst to be passed to the bottom of the
reactor riser through transfer line 26.
* * * * *