US3617497A - Fluid catalytic cracking process with a segregated feed charged to the reactor - Google Patents

Fluid catalytic cracking process with a segregated feed charged to the reactor Download PDF

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US3617497A
US3617497A US836404A US3617497DA US3617497A US 3617497 A US3617497 A US 3617497A US 836404 A US836404 A US 836404A US 3617497D A US3617497D A US 3617497DA US 3617497 A US3617497 A US 3617497A
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catalyst
hydrocarbon
feed
gasoline
process
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Millard C Bryson
James R Murphy
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Chevron Research and Technology Co
Gulf Research and Development Co
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    • CCHEMISTRY; METALLURGY
    • C10PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10GCRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • YGENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02PCLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P30/00Technologies relating to oil refining and petrochemical industry
    • Y02P30/40Ethylene production
    • Y02P30/44Cracking, e.g. steam cracking
    • Y02P30/446Catalytic cracking

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.

Description

United States Patent [72] Inventors Millard C. Bryson Conway, Pa.; James R. Murphy, Huntington Station, N.Y.

[21] Appl. No. 836,404

[22] Filed June 25,1969

[ 45] Patented Nov. 2, 1971 [73] Assignee Gulf Research & Development Company Pittsburgh, Pa.

[54] FLUID CATALYTIC CRACKING PROCESS WITH A SEGREGATED FEED CHARGED TO THE REACTOR 11 Claims, 4 Drawing Figs.

52 vs. C! 208/80, 23/288 s, 208/74, 208/120, 208/128, 208/130,

51 1111.01 ..C0lb33/28, Cl0gll/l8,Cl0qll/20 501 FieldofSearch ..208/80,49, 128, 130, 153, 160; 23/288 [5 6] References Cited UNITED STATES PATENTS 10/1960 Marshall et al.

3,246,960 4/I966 Sharpetal 3,524,809 8/1970 Hansford 23/288 208/l I l 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 of5 seconds ofless and in the presence ofa 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.

PATENTED'Mnv 2 l97| SHEET 1 IF 3 4 vy 0mm: ADDED r00 LATE REACTOR RESIDENCE TIME 0 7'0 5 SECONDS ADD/770M I I l I 0 7'0 .5 SECONDS FROM T/MELDW MOLECULAR I'VE/6H7 CHARGE ADDED I N VENTORS M/LL 420 c. BRYSD/V JAMES R. MURPHY FLUID CATALYTIC CRACKING PROCESS WITH A SEGREGATED FEED CHARGED TO THE REACTOR 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 to 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 condi-' tions 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 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 l0 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. ln 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.l second interval in the reaction zone than in the second 0.] 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 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 efiluent. 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. lf 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 max imum 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 0n 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 monoor 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 l.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 crosshatched 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 1 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. I, 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 and lighter gases and coke, while the monoand 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 com ponents 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 monoand 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 paraffms, naphthenes and monoand 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 efi'ect, 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 per centage 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 efiects 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 F. The upper limit can be about l,l00 F., or more. The preferred temperature range is 950 to l,050 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 monoand 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, 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 S-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:

o-zo 454s 75 Size (Microns) 20-45 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 l2zl 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 to 700 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 F. or 1,250 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 efi'ect 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 backmixmg.

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 bacltmixing. (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 attercracking 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 to l,100 F. As much as 5 to 20 percent of the fresh charge can boil above this range. Some residual oil can be charged. A to percent recycle rate can be employed. Generally, the recycle will comprise 650 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 -Hiquid 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 Zcolite (50-60 Kellogg 2 Hour Activity) Catalyst Cracking Conditions Temperature: "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 Gal Comporition (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;,+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 -Hiquid 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 I 1.95 Gravity: AP1 29.7 29.4 Sulfur: Percent 0.42 0.36 Viscosity, SUS at: F.

210 38.6 37.3 Carbon Residue,

Ramsbottom: Percent ASTM Aniline Point; F. 188 184 Bromine Number. D1159 2.8 3.0 Pour Point, 1397: I 90 Nitrogen: p.p.m. 710 450 Metull: p.p.m.

Vanadium 0.2 0.4

Nickel 0.2 0.1 Distillation Vac. (Corres. to

760 mm. Hg)

10% over at: F. 568 556 95 to bed formation is permitted to occur. The results are shown in table 2.

Catalyst I percent 60 percent Amorphous silicazeolite, 40 alumina percent TABLE 2 silicaalumina Test 1 2 3 4 Kellogg Activity Catal I) l st Catalyst bed formatiom. Yes No Yes No Cracking temperature. F 950 950 1.000 1. 000 Space velocity ttotalteed) 19.2 100 19.3 100 Reacm' Contact time, seconds 0. 5 (1) 2.0 Recycle. percent by volume 2. 4 5. 3 None None Conversion. percent by volume. 72. 9 77.1 76. 2 80.11 Fresh Feed Rate: BID 13.704 Yields. percent by volume of Reactor Bed Temperature: F. 926 935 fresh feed: Feed Preheat Temperature: F. 700 649 Total: R C1 9.0 10.4 11.7 .3 eactor Bed Pressure. p.s.1.g. ll.5 11.0 C 6 5 q 0 5 v1 T 1F 4- l5 pace eoc1y.( ota ce Total.

Wt./Hr./Wt. 3.94 3.07 0. 14. 2 16. 0 15.8 17. 7 Catalyst to Oil Ratio 4= 6. 8 7. 6 8. 0 7. R

(Tomi Feed); w w 12,5 93 Debutanized gasoline 545. 8 5= 8 t .1 74 3 3| 4 07 plus gasolino 44.2 47.0 44.8 50.11 Total 0. plus liquid 106.8 100.5 107. 5 111.1! Carbon on Rcgeneratcd Cat. C; and lighter. percent by k b Wt. 04 0.38 weight 3. 6 .5 4.1 3.1 Conversion: '5 by Volume G ({pke. ptercent by weight, 5.6 .0 5. 0 4. 5

850 ill! 00 Bile: Motor. clear 711. 3 .6 so. a 711. 3 lklotor. rgnslii cc 85.?! .2 86.2 esearc .c ear 92. 6 9.. .Il. Operation Cond1t1ons. Regenerator Research p u 3 cc 10 2 3 99. 5 U8. 7

2mins. Regen. Bed Temperature: F. l,l4l 1,166 3 Total Regen. Air: MlbJl-lr. 153.7 166.72 Bed backm'xmg' firm 0 087 0 083 A comparison of tests l and 2 of table 2, both conducted at 950 F., shows the deleterious effect of extended residence Yields: by volumc of Fresh Feed tlme when employing a zeohte catalyst. The resldence time of test 2 was only 0.5 second and yet It exhlblted a higher Debmnmd Gasoline 4H gasolme y1eld and a lower C1 and lighter yleld than test l in Bumwmne 2L2 2L6 wh1ch the residence ume was cons derably longer due to a i-Butanc 7.6 10.3 lower space veloc1ty and backm1x1ng ar1s1ng 1n the dense B H catalyst bed. A comparison of tests l and 2 shows that an ex- Propylene tended residence time gives arise to aftercracking which Propane 4.2 5.7 diminishes gasoline yield and increases the yield of products P em boiling lower than gasoline. Zi'gh'tifgl 40 Comparing test 3 with test l, both involving dense bed Z wL M 249 cracking, it is seen that raising the cracking temperature from Coke: 91 by m. 7.73 7.11 950 to L000 F., provided a slgmficant mcrease 1n convers1on but very little increase in debutanized gasoline yield and a Inspections higher yield of C and lighter, showing that the high degree of aftercracking occurring in a dense bed reaction system MotonClcar 111.3 prevents effective control of gasoline yield via temperature Motor. +3 cc. TEL 86.l 119.4 adjustment. Research Clear. 94.0 93.4

- o r s 4 w 5 v0 vin n nb d Search cc TEL 100A 983 C mpa mg te t 1th te t 2, both in l g o e As shown in tablev l, 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 6L0 percent yield of gasoline compared to only 47.5 percent gasoline yield with the amorphous catalyst. However, while the total yield of C and C hydrocarbons is about the same for the zeolite and the amorphous catalyst, the proportion of these C and C hydrocarbons which are oleflnic 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 and C olet'ms 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 cracking and very low residence times within the range of this invention, it is seen that raising the cracking temperature from 950 to l,000 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 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 4provided good yields of C olefin and C 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 olefin and C 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.

TABLE 3 Test 1 2 3 4 Catalyst Charge stock Charge stock inspections:

Gravity, API 25. 6 25. 6 30. 5 21. 4 Sulfur, percent by weight 0. 8 0.8 0. 65 1. Ramsb. carbon residue, percent by weight 0. 42 0. 42 0. 09 0. 73 Vacuum distillation (corres. to 760 mm. Hg) F. at, percent by volume:

10 580 580 510 809 30 692 692 629 831 50 767 767 659 873 70- 847 847 684 921 90 969 969 712 1,016 C ,\-percentage of total atoms which are aromatic atoms 0. 18 0. 17 20 Operating conditions:

Temperature, F 940 940 940 .140 Space velocity (total feed) weight/ hour/weight 6. 2 6.2 6. 2 6. 2 Gatalyst-to-oil ratio (total 1'eed) 7. 9 8. 0 8. 1 8. 0 Slurry oil recycle, percent by volume of fresh feed 5. 2 5.0 9. 1 Carbon on spent catalyst, percent by weight 1. 22 1.25 0.93 1. 55 Carbon on regen. cat., percent by weight 0.3 0.3 0.3 0.3 Gas oil conversion, percent by volurne of fresh feed 80. 7 85. 2 Yields, percent by volume ol'Iresh feed:

Debutanized gasoline 60. 4 l Butane-butene 16. 2 17. 7 Isobutane 9. 1 8. 3 n-Butane 2. 2 2. 1 Butenes 4. l 7. 3 Propane-propylene 10. 5 12. 1 Propane 4. 8 5. 2 Propylene 5. 7 6. 9 Light catalyst gas 011. 19. 3 Decantedoil- 11.9 Total 106. 4 103. 7

Gas, 02 and lighter, percent by weight 3 5 4. 4 Coke, percent by weight 5. 6 12. 5

Total 9.1 16.9

HzS,percent by weight... 0.1 0.2

1 Zeolite-bed. 2 Full range unsegregated. 3 Segregated-calculated total yield derived from both streams. 4 Light 50 percent by volume only of full range feed. 5 Heavy 50 percent by volume only of full range feed.

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 and lighter plus coke yield in test 3 was only 9.1 percent while the total C 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 and lighter plus coke yield in test 3 was 16.9 percent while the total C 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 F. and then charged into the bottom of the riser through inlet line l4. 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 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 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 to 1 l,00 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. Thc riser temperature range is about 900 to l, l00 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 and 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:] and l5:l, 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 46so 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 l,250 F. The temperature of the flue gas leaving the top of the catalyst bed in the regenerator can rise due to afterbuming 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 orderto prevent excessively high temperatures in the regenerator flue gas due to afterbuming, 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 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.

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 and l,l00 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 I 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 I wherein the catalyst to low molecular weight hydrocarbon weight ratio is between about 4:1 toabout 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 I 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 I wherein said catalyst is charged to the process at a temperature of at least about L240 F.

7. The process of claim I 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 ounds per square inch gauge.

11. he process of claim 1 wherein the space velocity based upon all feed streams is at least about weight of hydrocarbon feed per hour per weight of catalyst.

* 4 i I i November Dated Patent No. 3 617 497 Millard C. Bryson and James R. Murphy the above-identifies: patent 3; untreated as shcwn 'belsw:

It is certified that errcr appears in that aaid Letters Patent are here:

Invan'wfls) and COLUMN 8, LINE 3G, DELETE "15 A! INSERT --1U-- INSERT UNDER LINE 37 DELETE "Butenes 11.6" AND ---Butenes-- UNDER "n-Eutane" INSERT --'i.l.u-- 182 11! COLUMN l1 COLUMN 12 LIIMJ 10, IN COLUMN 1 AND 3, DELETE (1) AND INSERT ---{2) and seale this: 231%: day of 2* y W72

Claims (10)

  1. 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.
  2. 3. The process of claim 1 wherein the catalyst to low molecular weight hydrocarbon weight ratio is between about 4:1 to about 15:
  3. 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.
  4. 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.
  5. 6. The process of claim 1 wherein said catalyst is charged to the process at a temperature of at least about 1,240* F.
  6. 7. The process of claim 1 wherein the diluent vapor is steam, nitrogen, methane or ethylene.
  7. 8. The process of claim 1 wherein the effluent stream discharges from the cracking reactor in a lateral direction.
  8. 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.
  9. 10. The process of claim 1 wherein the pressure is about 5 to 50 pounds per square inch gauge.
  10. 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.
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US4166025A (en) * 1973-08-22 1979-08-28 Bocharov Jury N Process for purifying aromatic hydrocarbons
US4218306A (en) * 1979-01-15 1980-08-19 Mobil Oil Corporation Method for catalytic cracking heavy oils
US4422925A (en) * 1981-12-28 1983-12-27 Texaco Inc. Catalytic cracking
US4479870A (en) * 1984-02-29 1984-10-30 Jop Inc. Use of lift gas in an FCC reactor riser
US4483761A (en) * 1983-07-05 1984-11-20 The Standard Oil Company Upgrading heavy hydrocarbons with supercritical water and light olefins
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WO1987002695A1 (en) * 1985-10-30 1987-05-07 Chevron Research Company Gasoline octane enhancement in fluid catalytic cracking process with split feed injection to riser reactor
US4717466A (en) * 1986-09-03 1988-01-05 Mobil Oil Corporation Multiple riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4717467A (en) * 1987-05-15 1988-01-05 Mobil Oil Corporation Process for mixing fluid catalytic cracking hydrocarbon feed and catalyst
EP0259154A1 (en) * 1986-09-03 1988-03-09 Mobil Oil Corporation Upgrading naphtha in a single riser fluid catalytic cracking operation employing a catalyst mixture
EP0259156A1 (en) * 1986-09-03 1988-03-09 Mobil Oil Corporation Process for fluidized catalytic cracking with reactive fragments
US4752375A (en) * 1986-09-03 1988-06-21 Mobil Oil Corporation Single riser fluidized catalytic cracking process utilizing a C3-4 paraffin-rich co-feed and mixed catalyst system
US4764268A (en) * 1987-04-27 1988-08-16 Texaco Inc. Fluid catalytic cracking of vacuum gas oil with a refractory fluid quench
US4787967A (en) * 1986-09-03 1988-11-29 Mobil Oil Corporation Process for two-phase fluid catalytic cracking system
US4802971A (en) * 1986-09-03 1989-02-07 Mobil Oil Corporation Single riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4818372A (en) * 1985-07-10 1989-04-04 Compagnie De Raffinage Et De Distribution Total France Process and apparatus for the catalytic cracking of hydrocarbon feedstocks with reaction-temperature control
US4826586A (en) * 1986-09-03 1989-05-02 Mobil Coil Corporation Single riser fluidized catalytic cracking process utilizing a C3-4 paraffin-rich co-feed and mixed catalyst system
US4830728A (en) * 1986-09-03 1989-05-16 Mobil Oil Corporation Upgrading naphtha in a multiple riser fluid catalytic cracking operation employing a catalyst mixture
US4832825A (en) * 1985-02-07 1989-05-23 Compagnie De Raffinage Et De Distribution Total France Method for the injection of catalyst in a fluid catalytic cracking process, especially for heavy feedstocks
US4853105A (en) * 1986-09-03 1989-08-01 Mobil Oil Corporation Multiple riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4861741A (en) * 1986-09-03 1989-08-29 Mobil Oil Corporation Mixed catalyst system and catalytic conversion process employing same
US4863585A (en) * 1986-09-03 1989-09-05 Mobil Oil Corporation Fluidized catalytic cracking process utilizing a C3-C4 paraffin-rich Co-feed and mixed catalyst system with selective reactivation of the medium pore silicate zeolite component thereofo
US4865718A (en) * 1986-09-03 1989-09-12 Mobil Oil Corporation Maximizing distillate production in a fluid catalytic cracking operation employing a mixed catalyst system
US4869807A (en) * 1985-10-30 1989-09-26 Chevron Research Company Gasoline octane enhancement in fluid catalytic cracking process with split feed injection to riser reactor
US4871446A (en) * 1986-09-03 1989-10-03 Mobil Oil Corporation Catalytic cracking process employing mixed catalyst system
US4874503A (en) * 1988-01-15 1989-10-17 Mobil Oil Corporation Multiple riser fluidized catalytic cracking process employing a mixed catalyst
US4883583A (en) * 1985-07-16 1989-11-28 Compagnie De Raffinage Et De Distribution Total France Process for the catalytic cracking of hydrocarbons in a fluidized bed and their applications
US4888103A (en) * 1986-09-03 1989-12-19 Herbst Joseph A Process of stripping in a catalytic cracking operation employing a catalyst mixture which includes a shape selective medium pore silicate zeolite component
US4892643A (en) * 1986-09-03 1990-01-09 Mobil Oil Corporation Upgrading naphtha in a single riser fluidized catalytic cracking operation employing a catalyst mixture
US4990314A (en) * 1986-09-03 1991-02-05 Mobil Oil Corporation Process and apparatus for two-phase fluid catalytic cracking system
US5009769A (en) * 1989-02-06 1991-04-23 Stone & Webster Engineering Corporation Process for catalytic cracking of hydrocarbons
US5087349A (en) * 1988-11-18 1992-02-11 Stone & Webster Engineering Corporation Process for selectively maximizing product production in fluidized catalytic cracking of hydrocarbons
US5114682A (en) * 1988-11-18 1992-05-19 Stone & Webster Engineering Corporation Apparatus for recovering heat energy from catalyst regenerator flue gases
US5139748A (en) * 1990-11-30 1992-08-18 Uop FCC riser with transverse feed injection
US5271826A (en) * 1988-03-03 1993-12-21 Mobil Oil Corporation Catalytic cracking of coke producing hydrocarbons
EP0849347A2 (en) * 1996-12-17 1998-06-24 Exxon Research And Engineering Company Catalytic cracking process comprising recracking of cat naphtha to increase light olefins yields
US5976355A (en) * 1984-03-09 1999-11-02 Stone & Webster Engineering Corp. Low residence time catalytic cracking process
USRE36403E (en) * 1985-10-30 1999-11-23 Chevron Research And Technology Company Gasoline octane enhancement in fluid catalytic cracking process with split feed injection to riser reactor
US6565739B2 (en) 2000-04-17 2003-05-20 Exxonmobil Research And Engineering Company Two stage FCC process incorporating interstage hydroprocessing
US6569315B2 (en) 2000-04-17 2003-05-27 Exxonmobil Research And Engineering Company Cycle oil conversion process
US6569316B2 (en) 2000-04-17 2003-05-27 Exxonmobil Research And Engineering Company Cycle oil conversion process incorporating shape-selective zeolite catalysts
US6811682B2 (en) 2000-04-17 2004-11-02 Exxonmobil Research And Engineering Company Cycle oil conversion process
US6837989B2 (en) 2000-04-17 2005-01-04 Exxonmobil Research And Engineering Company Cycle oil conversion process
US20060163116A1 (en) * 2003-06-03 2006-07-27 Baptista Claudia Maria De Lace Process for the fluid catalytic cracking of mixed feedstocks of hydrocarbons from different sources
US20100083566A1 (en) * 2007-03-21 2010-04-08 Geir Remo Fredriksen Biogasoline
US20110062054A1 (en) * 2007-12-20 2011-03-17 Yongcan Gao Improved integrated process for hydrogenation and catalytic cracking of hydrocarbon oil
US20150275754A1 (en) * 2012-09-27 2015-10-01 Huaichao Chen Vapor cracking catalyst, preparation method thereof, and combustion method of hydrogen obtained by vapor cracking
WO2017065810A1 (en) * 2015-10-14 2017-04-20 Saudi Arabian Oil Company Processes and systems for fluidized catalytic cracking

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US3948757A (en) * 1973-05-21 1976-04-06 Universal Oil Products Company Fluid catalytic cracking process for upgrading a gasoline-range feed
US4166025A (en) * 1973-08-22 1979-08-28 Bocharov Jury N Process for purifying aromatic hydrocarbons
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US4147617A (en) * 1978-04-06 1979-04-03 Mobil Oil Corporation Processing hydrocarbon feed of high carbon residue and high metals content
US4218306A (en) * 1979-01-15 1980-08-19 Mobil Oil Corporation Method for catalytic cracking heavy oils
US4422925A (en) * 1981-12-28 1983-12-27 Texaco Inc. Catalytic cracking
US4602993A (en) * 1982-05-13 1986-07-29 Ashland Oil, Inc. Carbo-metallic oil conversion
US4483761A (en) * 1983-07-05 1984-11-20 The Standard Oil Company Upgrading heavy hydrocarbons with supercritical water and light olefins
US4479870A (en) * 1984-02-29 1984-10-30 Jop Inc. Use of lift gas in an FCC reactor riser
US5976355A (en) * 1984-03-09 1999-11-02 Stone & Webster Engineering Corp. Low residence time catalytic cracking process
US4832825A (en) * 1985-02-07 1989-05-23 Compagnie De Raffinage Et De Distribution Total France Method for the injection of catalyst in a fluid catalytic cracking process, especially for heavy feedstocks
US4818372A (en) * 1985-07-10 1989-04-04 Compagnie De Raffinage Et De Distribution Total France Process and apparatus for the catalytic cracking of hydrocarbon feedstocks with reaction-temperature control
US4883583A (en) * 1985-07-16 1989-11-28 Compagnie De Raffinage Et De Distribution Total France Process for the catalytic cracking of hydrocarbons in a fluidized bed and their applications
WO1987002695A1 (en) * 1985-10-30 1987-05-07 Chevron Research Company Gasoline octane enhancement in fluid catalytic cracking process with split feed injection to riser reactor
USRE36403E (en) * 1985-10-30 1999-11-23 Chevron Research And Technology Company Gasoline octane enhancement in fluid catalytic cracking process with split feed injection to riser reactor
US4869807A (en) * 1985-10-30 1989-09-26 Chevron Research Company Gasoline octane enhancement in fluid catalytic cracking process with split feed injection to riser reactor
US4966681A (en) * 1986-09-03 1990-10-30 Mobil Oil Corporation Multiple riser fluidized catalytic cracking process utilizing a C3 -C4 paraffin-rich co-feed and mixed catalyst system
US4802971A (en) * 1986-09-03 1989-02-07 Mobil Oil Corporation Single riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4787967A (en) * 1986-09-03 1988-11-29 Mobil Oil Corporation Process for two-phase fluid catalytic cracking system
US4826586A (en) * 1986-09-03 1989-05-02 Mobil Coil Corporation Single riser fluidized catalytic cracking process utilizing a C3-4 paraffin-rich co-feed and mixed catalyst system
US4830728A (en) * 1986-09-03 1989-05-16 Mobil Oil Corporation Upgrading naphtha in a multiple riser fluid catalytic cracking operation employing a catalyst mixture
US4752375A (en) * 1986-09-03 1988-06-21 Mobil Oil Corporation Single riser fluidized catalytic cracking process utilizing a C3-4 paraffin-rich co-feed and mixed catalyst system
EP0259156A1 (en) * 1986-09-03 1988-03-09 Mobil Oil Corporation Process for fluidized catalytic cracking with reactive fragments
US4861741A (en) * 1986-09-03 1989-08-29 Mobil Oil Corporation Mixed catalyst system and catalytic conversion process employing same
US4863585A (en) * 1986-09-03 1989-09-05 Mobil Oil Corporation Fluidized catalytic cracking process utilizing a C3-C4 paraffin-rich Co-feed and mixed catalyst system with selective reactivation of the medium pore silicate zeolite component thereofo
US4865718A (en) * 1986-09-03 1989-09-12 Mobil Oil Corporation Maximizing distillate production in a fluid catalytic cracking operation employing a mixed catalyst system
EP0259154A1 (en) * 1986-09-03 1988-03-09 Mobil Oil Corporation Upgrading naphtha in a single riser fluid catalytic cracking operation employing a catalyst mixture
US4871446A (en) * 1986-09-03 1989-10-03 Mobil Oil Corporation Catalytic cracking process employing mixed catalyst system
US4990314A (en) * 1986-09-03 1991-02-05 Mobil Oil Corporation Process and apparatus for two-phase fluid catalytic cracking system
US4717466A (en) * 1986-09-03 1988-01-05 Mobil Oil Corporation Multiple riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4888103A (en) * 1986-09-03 1989-12-19 Herbst Joseph A Process of stripping in a catalytic cracking operation employing a catalyst mixture which includes a shape selective medium pore silicate zeolite component
US4892643A (en) * 1986-09-03 1990-01-09 Mobil Oil Corporation Upgrading naphtha in a single riser fluidized catalytic cracking operation employing a catalyst mixture
US4853105A (en) * 1986-09-03 1989-08-01 Mobil Oil Corporation Multiple riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4764268A (en) * 1987-04-27 1988-08-16 Texaco Inc. Fluid catalytic cracking of vacuum gas oil with a refractory fluid quench
US4717467A (en) * 1987-05-15 1988-01-05 Mobil Oil Corporation Process for mixing fluid catalytic cracking hydrocarbon feed and catalyst
US4874503A (en) * 1988-01-15 1989-10-17 Mobil Oil Corporation Multiple riser fluidized catalytic cracking process employing a mixed catalyst
US5271826A (en) * 1988-03-03 1993-12-21 Mobil Oil Corporation Catalytic cracking of coke producing hydrocarbons
US5087349A (en) * 1988-11-18 1992-02-11 Stone & Webster Engineering Corporation Process for selectively maximizing product production in fluidized catalytic cracking of hydrocarbons
US5114682A (en) * 1988-11-18 1992-05-19 Stone & Webster Engineering Corporation Apparatus for recovering heat energy from catalyst regenerator flue gases
US5009769A (en) * 1989-02-06 1991-04-23 Stone & Webster Engineering Corporation Process for catalytic cracking of hydrocarbons
US5139748A (en) * 1990-11-30 1992-08-18 Uop FCC riser with transverse feed injection
EP0849347A2 (en) * 1996-12-17 1998-06-24 Exxon Research And Engineering Company Catalytic cracking process comprising recracking of cat naphtha to increase light olefins yields
US5846403A (en) * 1996-12-17 1998-12-08 Exxon Research And Engineering Company Recracking of cat naphtha for maximizing light olefins yields
EP0849347A3 (en) * 1996-12-17 1998-12-09 Exxon Research And Engineering Company Catalytic cracking process comprising recracking of cat naphtha to increase light olefins yields
US6837989B2 (en) 2000-04-17 2005-01-04 Exxonmobil Research And Engineering Company Cycle oil conversion process
US6569315B2 (en) 2000-04-17 2003-05-27 Exxonmobil Research And Engineering Company Cycle oil conversion process
US6565739B2 (en) 2000-04-17 2003-05-20 Exxonmobil Research And Engineering Company Two stage FCC process incorporating interstage hydroprocessing
US6811682B2 (en) 2000-04-17 2004-11-02 Exxonmobil Research And Engineering Company Cycle oil conversion process
US6569316B2 (en) 2000-04-17 2003-05-27 Exxonmobil Research And Engineering Company Cycle oil conversion process incorporating shape-selective zeolite catalysts
US20060163116A1 (en) * 2003-06-03 2006-07-27 Baptista Claudia Maria De Lace Process for the fluid catalytic cracking of mixed feedstocks of hydrocarbons from different sources
US7736491B2 (en) 2003-06-03 2010-06-15 Petroleo Brasileiro S.A. - Petrobras Process for the fluid catalytic cracking of mixed feedstocks of hydrocarbons from different sources
US20100083566A1 (en) * 2007-03-21 2010-04-08 Geir Remo Fredriksen Biogasoline
US9260667B2 (en) 2007-12-20 2016-02-16 China Petroleum & Chemical Corporation Combined process of hydrotreating and catalytic cracking of hydrocarbon oils
US9309467B2 (en) * 2007-12-20 2016-04-12 China Petroleum And Chemical Corp. Integrated process for hydrogenation and catalytic cracking of hydrocarbon oil
US20110062054A1 (en) * 2007-12-20 2011-03-17 Yongcan Gao Improved integrated process for hydrogenation and catalytic cracking of hydrocarbon oil
US10006362B2 (en) * 2012-09-27 2018-06-26 Huaichao Chen Vapor cracking catalyst, preparation method thereof, and combustion method of hydrogen obtained by vapor cracking
US20150275754A1 (en) * 2012-09-27 2015-10-01 Huaichao Chen Vapor cracking catalyst, preparation method thereof, and combustion method of hydrogen obtained by vapor cracking
WO2017065810A1 (en) * 2015-10-14 2017-04-20 Saudi Arabian Oil Company Processes and systems for fluidized catalytic cracking
US9896627B2 (en) 2015-10-14 2018-02-20 Saudi Arabian Oil Company Processes and systems for fluidized catalytic cracking

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Publication number Publication date
FR2047929A1 (en) 1971-03-19
CA930684A (en) 1973-07-24
ES381104A1 (en) 1972-11-01
FR2047929B1 (en) 1974-09-20
JPS4914321B1 (en) 1974-04-06
GB1266070A (en) 1972-03-08
DE2031448A1 (en) 1971-01-07
NL7009391A (en) 1970-12-29
CA930684A1 (en)

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