US2734850A

US2734850A - brown

Info

Publication number: US2734850A
Authority: US
Priority date: 1951-05-19
Publication date: 1956-02-14
Anticipated expiration: 1973-02-14
Also published as: BE508628A; NL81459C; GB710144A; DE937723C

Description

Feb. 14, 1956 FLUID COKING OF' HEAVY HYDROCARBONS Filed May 19. 1951 W. BROWN 2 Sheets-Sheet 1 gFLuGAs Qfczrnes fowrz nverzbor Q55 Bda Clbborne Feb. 14, 1956 J, wl BROWN 2,734,850

FLUID COKING OF HEAVY HYDROCARBONS Filed May 19, 1951 2 Sheets-Sheet 2 Livia Fra-2 Uamee rowrz rzvembor P051' Clbjborne Unite States Patent p 2,734,350 FLUID como or HEAVY HYnnocAnnoNs James Woodrow Brown, Elizabeth, N. J., assignor to Esso Research and Engineering Company, a corporation of Delaware Application May 19, 1951, Serial No. 227,169

4 Claims. (Cl. 19d-49) The present invention relates to the art of treating hydrocarbons and, more particularly, to a Vcombination of coking and catalytically cracking heavy hydrocarbons in a three-stage fluid system wherein heat for the coking step is supplied by indirect heat exchange with freshly regenerated catalyst. Y

Heretofore a process has been proposed in U. S. Patent 2,388,055 wherein heavy hydrocarbons were converted into lighter products by first thermally cracking or coking the -heavy feed in the presence of inert solids such as hot coke particles, thereby also removing ash and coke forming constituents from the feed and subsequently cracking the resulting vapors in the presence of a catalyst in a second zone. However, in such a process two separate combustion zones had to be -provided in addition to the two conversion zones, ,one combustion zone serving to` heat the coke particles to a temperature suiiicient to supply the heat required Vby the coking step while-the other combustion zone served to regenerate the catalyst by burning off the carbonaceous matter deposited on the catalyst during the catalytic conversion step.

-It is the object of the present invention to improve on the aforementioned system and to increase the yield of recoverable coke product. Another object is to improve the heat balance of the process and to take better advantage of the heat produced in regenerating the catalyst. -A further object is to reduce the `number of steps and of major reaction vessels required to carry out the process. Other objects will appear from the following detailed description and claims. v

Fig. l of the accompanying drawing diagrammatically illustrates an apparatus for carrying out a preferred modication-of the invention, according to which both a coking zone containing a heat exchanger anda catalytic cracking zone superimposed on the coking zone are located -within a single vessel.

Fig. 2 of the drawing illustrates an alternative embodiment Vof the invention wherein the coking zone and the catalytic cracking zone are contained in separate vessels, thus permitting separation and recovery of light and heavy fractions from the coker products before .a selected cut of the latter is passed to the catalytic cracker.

In practicing the preferred embodiment of the invention illustrated n Fig. l, a residual petroleum stock, as Aforlinstance a residuum having a gravity of about API and an atmospheric boiling lrange in excess of about l100 F. is preheated to about 70.0 F. by conventional -means (not shown) and is then injected through line 1 and jets 2 into the lower coker portion 3 of the reactor. The-latter may contain about 0.5 to lbs. of coke particles perlbJhr. of feed, the icoke particles ranging in size between about 40 and 150 microns. Steam is injected through line 6 into the bottom of the coker 3 so as Vto produce a total upward vapor velocity of about 1.5 to 3 ft./sec. within the Coker, thereby maintaining the coke particles in the form of-a dense uidized bed 4 having a density of about 20 to .40 lbs/cu. ft. and having an upper leverS above which there is a more dilute phase 7 having a density of only about 0.01 to l lb./cu. ft. z At the indicated velocities, the hydrocarbon vapors formed in the coking zone stay there for a period corresponding to a residence time of about 3 to 10 seconds, before passing overhead into the catalytic cracking zone as subsequently described.

A shell-and-tube heat exchanger 8 submerged in the lluid coke bed 4 maintains the latter at about 850 to 950 F., which is a temperature suited for obtaining the desired degree of thermal cracking and low temperature coking of theresiduum feed. The-necessary heat is supplied by withdrawing hot catalyst at about 1250 F. from catalyst regenerator 30 through line 37 and circulating such hot catalyst in a fluidized state through the shell of heat exchanger 8. The coke mixture of lluidized bed 4 in which the heat exchanger 8 is submerged is maintained in a fluid state within the tubes 11of the heat exchanger 8.

Steam at a rate sucient to give a total upward vapor velocity of at least 1.15 ft./sec. within the heat exchanger shell is injected through manifold 10 to the bottom of the shell below perforated distributor plate 9 which extends across the shell between the exchanger tubes 11 containing the fluidized coke particles. In this manner the hot catalyst is maintained within the shell as a dense fluidized phase having an upper level 1'7 and a dilute phase thereabove, from which the steam is withdrawn through line 12 .which may pass through a cyclone separator 13 for recovering entrained catalystnes from the steam. With the indicated aeration velocities in the heat exchanger an over-:all heat coeicient of about 75 B. t. u./hr./sq. ft. F. can'he .obtained and thus the required coking heat can-be supplied .by circulating regenerated 1250 F. catalyst through the Vexchanger to maintain 1l50 F. catalyst in the exchanger. Under such conditions a tube metal temperature of only about 1000 F. prevails in the exchanger, allowing the'use ofregular carbon steel tubes. Catalyst particles at about 12150 F. arey withdrawn from the heat exchanger shell through line 14 and returned to regenerator 3,0 via standpipe 15 and transfer line 16.

Coke is formed in the process primarily from the relatively heavy components of the feed, which component-s are most Vreadily deposited togetherwith the ash constituents on the `luidized particles in the coker and have an especial-ly great coke-forming tendency as indicated by a high Conradson carbon value.v Particles of the net coke productare withdrawn from coke bed 4 downwardly -to storage through line 17 while the aeration steam and hydrocarbon vapors liberated in coke bed 4 pass from coker 3 upward, preferably`through cyclones 29 or other entrainment separators such as inclined bailes, and through perforated distribution v 19 ofthe two-zone reactor. The diameter of the cracking portion 19 is preferably larger than the diameter of the lower, cokng portion 3 of the reactor because the'total vapor `Volume is increased during catalytic conversion. The gas Velocity in cracker -19 is again suiicient to maintain the hot catalyst particles above plate 19 in a dense lluid phase 20 having an upper level 21,*above which is a more `dilute `phase Z2. The hydrocarbon vapors are cracked on contactwith the synthetic silica-alumina gel or othersuitable cracking catalyst in dense bed 2i) and the cracked vapors together with the steam` are withdrawn overhead through cyclone 23 and passed through line 25 to :a `conventional fractionation system for recovery. The

contact time of the vapors in the cracking zone rangesfrom. about S10 lto '30 seconds. Catalyst lfines `entrained with the vapors-.are separated in the cyclone and returned to dense bed 20 through Ldip `leg`24 which extends below theilevel 21 ofthe crackingzo'ne.

Heat of reaction is supplied to the crackingzone 20' by means of'hot regenerated catalyst whichis circulated from regeneration zone 30 through line 26 a-t arate suif cient to maintain the temperature of the cracking zone grid 18 into the cracking portion 3 at about 1000 F. while pressure in the dilute phase of the cracking zone is maintained at about l p. s. i. g. Spent catalyst is withdrawn from catalyst bed through standpipe 15 and recirculated to the regeneration zone through transfer line 16 with the aid of a lift gas such as air which is injected into line 16 at the bottom of the standpipe through line 27. Minor amounts of an aeration gas such as steam may also be injected into standpipe 15 through one or more lines 28 to assure a free flowing condition for the spent catalyst.

From transfer line 16 the mixture of spent cracking catalyst and air is passed through perforated distributor plate 31 into the regenerator 30 where the carbonaceous contaminants are burned oif the catalyst while the catalyst is maintained in the form of a fluid bed at a temperature of about i250" F. and at a pressure of about 1 p. s. i. g., low regenerator pressures being preferred so as to keep down air compression costs. The flue gas produced in the regenerator is removed overhead through a cyclone 32 wherein entrained catalyst fines are separated and returned through dip leg 33 to a point below level 34 of the fluid regenerator bed 35. As described earlier, a substantial portion of the heat produced in the regenerator 30 is utilized to supply the heat of coking to coking zone 4 by means of the indirect heat exchanger 8, while another portion of the heat produced in the regenerator is supplied to the catalytic cracking zone 19 by returning hot regenerated catalyst through line 26 and mixing it directly into fluid bed 20. Excess heat may be removed from the regeneration zone by means of an indirect heat exchanger 36 through which a cooling medium such as water is circulated to generate steam for use in the process or elsewhere. Alternatively, it is also feasible to pass the heavy hydrocarbon feed through heat exchanger 36 and thus to utilize the excess heat from the regenerator for preheating the hydrocarbon feed to the desired temperature before it is fed to the coker.

According to another embodiment of the invention, separately fueled heaters for the coker solids can be eliminated by using the excess regenerator heat for operating the coker and at the same time the excess contamination of the catalyst by carbonaceous deposit during cracking is reduced by fractionating the hydrocarbon vapors produced in the coker into a coke-forming, heavy liquid fraction which may amount to about l0 to 35% of the residual hydrocarbon fed to the coker and which is recycled to the coker, and a lighter fraction which alone is then passed to the catalytic cracker. To produce minimum coke in the cracker, the catalytic cracking feed preferably consists of a heavy gas oil having a boiling range between about 650 to 950 F. and amounting to about 25 to 50% of the residual hydrocarbon fed to the coker. However, when catalyst contamination is not excessive and a better naphtha product is desired, at times it may be beneficial to include even lighter products in the catalytic cracking feed so as to reform the low grade coker naphtha simultaneously with the cracking of the gas oil, thereby increasing the total yield of high octane motor fuel. Moreover, especially when combustion of the catalytic coke deposit obtained from the narrow cut gas oil is insufficient to satisfy the heat requirements of both the catalytic cracking step and of the coking step to which a substantial amount of the heavy coke-forming feed fraction is normally recycled, it may also be desirable to reduce the amount of the heavy recycle by raising the cut-olf point between residual recycle and gas-oil range cracking feed to such an extent that coke-forming constituents are included in the catalytic cracking feed stock in a sufficient amount to deposit enough coke in the catalytic cracking step so as to balance the heat requirements of the entire process.

Referring now specifically to Fig. 2 of the drawing a vacuum pitch having a boiling range in excess of about 1100 F. (atmospheric equivalent) and a gravity of about 5 API is preheated to about 700 F. and sprayed through line 101 into a coking vessel 102. In vessel 102 finely divided coke particles ranging in size between 40 and 150 microns are maintained at a temperature of about 900 to 1000 F., preferably at about 950 F., in the form of a dense fluidized bed 103 having an upper level 104 by passing steam through line 105 below perforated distributor plate 106 at a sufficient rate to produce in coker 102 a total superficial vapor velocity of about 1.5 to 3.0 feet per second. The required heat of reaction is supplied to coking bed 103 at a rate suhicient to maintain the temperature of the bed at the desiredV level, the heat transfer being accomplished by circulating a part of the coke particles back and forth to regenerator 301 where they become heated by indirect heat exchange with the catalyst bed maintained at a temperature of about 1250 F. as later herein described.

The vapors produced in the coker and amounting to about 70 to 90 wt. percent of the feed are withdrawn over head from the dilute phase of the coker through a solids separator such as cyclone 107 wherein entrained'coke fines are separated and returned to fluid bed 103 through dip leg 108 while the vapors themselves are passed through line 109 to distillation tower 110 for separation Yinto de-` sired fractions. The heavy bottoms fraction which hasY an atmospheric boiling point above about 950 F. is recycled from the distillation tower to the coker through

lines

111 and 101. The intermediate fraction, which lis preferably a heavy gas oil having an atmospheric boiling range between about 650 and 950 F., is passed through line 112 and mixed in transfer line 113 with hot regenerated catalyst withdrawn from regeneration zone 301 through standpipe 302. The hot catalyst thus serves in a known manner to vaporize the oil as well as to supply the heat of reaction required in the catalytic cracking step. The mixture of catalyst and vaporized gas oil formed in the transfer line 113 is then introduced into cracking vessel 201 below a perforated distributor plate 202. In vessel 201 the mixture of oil vapor and catalyst is maintained at a temperature of about 800 F. to `1200 F., preferably at about 900 to 1000 F., and kept in the form of a dense ud phase by passing the vapors, and auxiliary gas if necessary, upward through the reactor at a proper velocity as is well known. From vessel 201 the cracked product vapors are withdrawn overhead through cyclone 203 and line 204 to another conventional distillation system 205 where desired fractions such as hydrocarbon gas, naphtha and gas oil are separated.

As a result of cracking the coker gas oil, relatively large quantities of carbon amounting to as much as 5 to 15 weight percent based on the pitch fed to the process, are deposited on the catalyst in cracking vessel 201, thereby reducing the activity of the catalyst.V Such spent catalyst is therefore continuously withdrawn from vessel 201 through standpipe 206, admixed with air introduced at 207, and the resulting mixture is then passed through transfer line 208 to regenerator vessel 301 wherein it is evenly distributed with the aid of perforated plate 303. In the regenerator the catalyst is likewise maintained in the forrn of a dense uid bed 304 having an upper level 30S and a more dilute phase 306 thereabove, whence the resulting flue gases are withdrawn through cyclone 307 and line 308. In this manner the carbonaceous catalyst deposit is burned oif in the regenerator at 1000 to 1500 F., preferably at about 1250 F. with the production of suicient heat to operate not only the catalytic cracking step but also to supply the heat requirements of the coking step.

The large amount of heat required for coking of the heavy feed is supplied to fluid bed 103 by circulating coke between coke bed 103 and shell-and-tube type heat exchanger 309 which is submerged in the hot regenerator bed 304. This circulation of coke is accomplished by withdrawing the coke from coke bed 103 through standpipe 114, injecting steam into the withdrawn coke through line 115 at the foot of standpipe 114 and passing the resulting .dilute suspension of coke particles steam through transfer line 116 to the -shell lof heat exchanger 309 similar to exchanger 8 illustrated in Fig. '1. The coke particles are maintained in the* heat exchanger shell as a densefluid bed for a vsuliicientlength of time to ralse their temperature to about 950-1050" F. by indirectheat transfer from the hot iiuidcatalystbed 30.4.which passes through the tubes of the heat exchanger 309. `The reheated coke is linally withdrawn from the heatexchanger through standpipe 310 and returned .to coke bed 103 after mixing with steam in transfer line 105 as previously described.

Alternatively, the heat exchanger may be located within coke bed 103, in which eventa proper amount of hot regenerated catalyst is-passed from regeneratorobed 304 to the shell of the exchanger located in the coker while the iiuid coke bed` passes through the tubes of the heat exchanger in the manner previously described -in-.connection with Fig. l. In such a design-the regenerated catalyst, maintained fluid and partially cooled in the shell of the heat exchanger Vlocated in the coker, 'may be subsequently mixed with the coker gas .oil instead `of using hot catalyst withdrawn directly from the regenerator as shown in Fig. 2, and theresu'ltingsuspension of oil vapor and catalyst may then be fed to the catalytic cracking vessel through line 113. Any excess catalyst from the heat exchanger may be mixed with air `and returned to the regenerator similarly as spent catalyst is returned. The location of the heat exchanger in the coke. bed vat about 950 F. `as just described has the advantage of lower tube metal temperatures `as compared with the design actually shown in Fig. 2 where the exchanger is submerged in the regenerator bed at about 12507 F., and hence the location of the exchanger in the relatively cool coker bed may be preferred since less expensive construction materials can be used. On the other hand, in many instances location of the heat exchanger in the high temperature regenerator bed may be preferred since the periodic heating of the coke particles to the higher temperature of about l000 F. inherent in this modification tends to dry out the coke more completely and thus prevents any uidization diliiculties in the coker, and also results in a better coke product, all of which may outweigh the increased metal cost required in this design.

Having described specific embodiments of the invention as well as suitable methods ofv operation, it will be understood that this has been done for purposes of illustration and not of limitation. On the contrary, the invention can be varied and modified in numerous ways which will occur to persons skilled in the art without departing from the scope and spirit of the present disclosure or of the appended claims.

For instance, beyond the scope of the specific examples given, the invention is broadly applicable to the treatment of heavy residual crude stocks as well as cycle stocks having a boiling range above about 900 to ll50 F. (atmospheric equivalent) and a gravity between about to 20 API, and even to lighter stocks such as gas oils. The invention is of particular value with stocks having high coke forming tendencies as indicated by Conradson carbon values between about and 35 weight percent such as crudes obtained by atmospheric or vacuum distillation and representing about 2 to 25 vol. percent of the whole crude distilled, or the invention may be applied to clarified oil from catalytic cracking, various pitches, tars from visbreaking operations and the like. Prior to feeding to the coker, the heavy feed stocks may be cut back with naphtha or other light products, and preferably preheated to temperatures ranging from 200 to 1000 F., or especially 600 to 800 F. Moreover, the hydrocarbon feed may also be diluted in the various reaction zones with steam, recycle gas or other inert gas in amounts up to about 500 to 5000 cubic feet (at coker conditions) per barrel, since such diluent increases gas velocity in the cker such fluidizing velocities may .range from about 0.5 to 5 or 10 feet per second toestablish apparent densities'in the dense solids phase of about 10-50 lbs. per cu. ft. and about 0.01 yto 5 lbs. per cu. ft. in the disperse phase asis well known per se.

The contact solids used in the coker are preferably coke particles ranging in size from about 0 to 200 or 500 microns, thoughv other inert solids such as sand, vspent claysand lthe like may similarly be used if a coke product of .high ash content can be tolerated. Y v

The contact solids inthe. .catalyst cracking zone may beany tinely divided cracking catalyst such .as activated clays, activated alumina, synthetic composites of silica with alumina, magnesia` .and/or boria, .activated carbon or .other conventional cracking catalyst. The Particle size of .thesolids inthe catalytic cracking zone as well-as inthe regeneration zone, :and also apparent densities and specified gas velocities prevailing therein are substantially within. the same limits as given .above with reference to thecoker solids.v y

Reaction conditions may include coking temperatures offabout .800 .to 1200F.preferably S50-to 950 F. catalytic -crack-ing temperaturesof about 800 to 1200 F., preferably l9-00 to'1000 F.; and catalyst regeneration temperatures of about 1000 .to1500 F., preferably 1100 to 13.00 Vl"".,lo1 1oe importantlim-itation being that the 'regencrat-iontemperature be at least F. above `the ,coker temperature so as to permit the required indirect yheat exchange Ybetween coker solids and regenerated catalystto be accomplished in an, efcient manner. Also, of course, the `.regenera-tion temperature must notrbe high enough to cause serious deterioration of the catalyst, as is well known. In the coking and cracking zones, the weight ratio of oil to total `solids may be from about 0.1 to 5 w./hr./w. (weights per hour per weight); the ratio of catalyst to oil in the dilute catalyst suspension such as is fed to cracking zone 201 through transfer line 113 in Fig. 2 may be from about l to 20 parts by weight of catalyst for one part by weight of oil.

Furthermore, the physical arrangement of the illustrated apparatus may be modified in various ways. For instance, while the heat exchanger is shown entirely submerged in the coker bed in Fig. 1 illustrating a one-vessel converter and entirely submerged in the regenerator bed in Fig. 2 illustrating a two-vessel converter, the heat exchanger may be only partially submerged in the dense fluid bed and may be located in the reverse location in either case, requiring only minor modifications in the ow of the streams as described earlier herein. Moreover, the heat exchanger itself may be of any convenient design other than of the tube-and-shell type illustrated, provided that finely divided solids can be fluidized therein on either side of the heat transfer surface. Also, while in Fig. 1 of the drawing the regenerated catalyst cooled in the heat exchanger is shown to be returned to the regenerator, it can be passed to the catalytic reactor bed instead, especially where coker and catalytic reactor are contained in two separate vessels as' in Fig. 2, and then particularly when the heat exchanger is submerged in the Coker bed so that the catalyst is on the shell side of the exchanger. Likewise, injection of the feed stock directly into the dense coker bed as shown in Fig. 1 or into the more disperse phase as shown in Fig. 2 is optional, as either procedure offers certain advantages not possessed by the other.

Having given a full description of the invention and of the manner and process of using it, the invention is particularly pointed out and distinctly claimed in the claims which follow.

The invention claimed is:

1. A process for converting a residual petroleum stock which comprises the following steps: contacting a residual stock with coke particles maintained as a fluid bed at a coking temperature in a coking zone to form additional amounts of coke and hydrocarbon vapors, withdrawing net coke product from said coking zone, passing said hydrocarbon vapors after removing entrained solids therefromV directly to the lower portion of a cracking zone, contacting said hydrocarbon vapors therein with a uid bed of particulate cracking catalyst at a cracking temperature, withdrawing cracked product vapors from said cracking zone, withdrawing and mixing a portion of said cracking catalyst with a free oxygen-containing gas in a combustion zone, forming a iluid bed of the withdrawn catalytic solids therein at a regeneration temperature at least 100 F. above said coking temperature, withdrawing flue gases from said combustion zone, circulating regenerated catalyst from said regeneration zone to said cracking zone, also circulating regenerated catalyst from said regeneration zone and coke particles from said coking zone to a heat exchange zone, maintaining the regenerated catalyst and coke particles in said heat exchange zone in a tiuidized condition in separate adjacent chambers and in indirect heat exchange relation, withdrawing and returning cooled regenerated catalyst from said heat exchange zone to said process, and withdrawing and returning heated coke particles from said heat exchange zone to said coking zone to maintain said coking temperature.

2. In a residual oil conversion process wherein a residual oil is converted to conversion products in a coking zone by contact with uidized inert solids maintained at a coking temperature, wherein said conversion products are then further converted in a cracking zone by contact with tiuidized catalytic solids maintained at a cracking temperature with deposition of coke on said catalytic solids, wherein coke-containing catalytic solids are withdrawn from said cracking zone and regenerated as a iluid bed by contact with a free oxygen-containing gas at a -8 combustion temperature 4in a regeneration zone, and Wherein regenerated catalytic solids are circulated from said regeneration zone to said cracking zone to maintain said cracking temperature, the improvement which comprises maintaining said combustion temperature at least 100 F.

above said Icoking temperature, and maintaining said coking temperature by circulating said inert solids and said regenerated catalytic solids through a heat exchange zone in indirect heat exchange relation, the solids in said heat exchange zone being in a iluidized condition.

3. The residual oil conversion process of claim 2 wherein said heat exchange zone is situated within said regeneration zone.

' 4. A process according to claim 2 wherein said conversion products are withdrawn from said coking vzone and fractionally separated to obtain a gas oil fraction which is then passed to said cracking zone, the end boiling point of said gas oil fraction being regulated to control the amount of coke deposited on said uidized catalytic solids whereby a suiicient amount of coke is deposited on said uidized catalytic solids to meet the heat requirements of said coking and cracking zones when the coke on the catalyst is consumed in said regeneration zone.

References Cited in the fde of this patent UNITED STATES PATENTS 2,348,009 Johnson et al. May 2, 1944 2,388,055 Hemminger Oct. 30, 1945 2,436,486 Scheinenian Feb. 24, 1948 l2,447,505 Johnson Aug. 24, 1948 2,676,668 Lindsay Apr. 27, 1954

Claims

1. A PROCESS FOR CONVERTING A RESIDUAL PETROLEUM STOCK WHICH COMPRISES THE FOLLOWING STEPS: CONTACTING A RESIDUAL STOCK WITH COKE PARTICLES MAINTAINED AS A FLUID BED AT A COKING TEMPERATURE IN A COKING ZONE TO FORM ADDITIONAL AMOUNTS OF COKE AND HYDROCARBON VAPORS, WITHDRAWING NET COKE PRODUCT FORM SAID COKING ZONE, PASSING SAID HYDROCARBON VAPORS AFTER REMOVING ENTRAINED SOLIDS THEREFROM DIRECTLY TO THE LOWER PORTION OF A CRACKING ZONE, CONTACTING SAID HYDROCARBON VAPORS THEREIN WITH A FLUID BED OF PARTICULATE CRACKING CATALYST AT A CRACKING TEMPERATURE, WITHDRAWING CRACKED PRODUCT VAPORS FROM SAID CRACKING ZONE, WITHDRAWING AND MIXING A PORTION OF SAID CRACKING CATALYST WITH A FREE OXYGEN-CONTAINING GAS IN A COMBUSTION ZONE, FORMING A FLUID BED OF THE WITHDRAWN CATALYTIC SOLIDS THEREIN AT A REGENERATION TEMPERATURE AT LEAST 100* F. ABOVE SAID COKING TEMPERATURE, WITHDRAWING FLUE GASES FROM SAID COMBUSTION ZONE, CIRCULATING REGENERATED CATALYST FROM SAID REGENERATION ZONE TO SAID CRACKING ZONE, ALSO CIRCULATING REGENERATED CATALYST FROM SAID REGENERATION ZONE AND COKE PARTICLES FROM SAID COKING ZONE TO A HEAT EXCHANGE ZONE, MAINTAINING THE REGENERATED CATALYST AND COKE PARTICLES IN SAID HEAT EXCHANGE ZONE IN A FLUIDIZED CONDITION IN SEPARATE ADJACENT CHAMBERS AND IN INDIRECT HEAT EXCHANGE RELATION, WITHDRAWING AND RETURNING COOLED REGENERATED CATALYST FROM SAID HEAT EXCHANGE ZONE TO SAID PROCESS, AND WITHDRAWING AND RETURNING HEATED COKE PARTICLES FROM SAID HEAT EXCHANGE ZONE TO SAID COKING ZONE TO MAINTAIN SAID COKING TEMPERATURE.