Abstract:
An FCC apparatus places a quench chamber above a reactor vessel and a hot stripper below a reactor vessel to provide a progressively decreasing temperature profile up the structure of the FCC arrangement and equipment for sequential reaction control. A riser contains the primary catalytic reactions of the hydrocarbon vapor and delivers the reacted vapors to the reactor structure. Starting from the bottom of the structure the hot stripper has the highest temperature and desorbs or displaces hydrocarbons from the catalyst to terminate long residence time catalytic reactions. Above the hot stripper bulk separation equipment divides the main vapor and catalyst stream to limit residence time of major catalytic reactions. At a yet higher elevation and lower internal temperature quench equipment arrests thermal reactions of the vapor stream. This structure arrangement permits reliable control of reaction time to obtain desired products and enhances mechanical reliability of the structure.

Description:
CROSS-REFERENCE TO RELATED APPLICATIONS 
     This application is a continuation of U.S. Ser. No. 08/101,204 filed Aug. 3, 1993, now abandoned, which is a divisional of U.S. Ser. No. 07/766,498 filed Sep. 26, 1991 and issued as U.S. Pat. No. 5,234,578, which is a continuation-in-part of U.S. Ser. No 07/236,817, filed Aug. 26, 1988, now abandoned. 
    
    
     FIELD OF THE INVENTION 
     This invention relates generally to apparatus for the fluidized catalytic cracking of heavy hydrocarbon streams such as vacuum gas oil and reduced crudes. This invention relates more specifically to an apparatus for reacting hydrocarbons in an FCC reactor and separating reaction products from the catalyst used therein. 
     BACKGROUND OF THE INVENTION 
     The fluidized catalytic cracking of hydrocarbons is the main stay process for the production of gasoline and light hydrocarbon products from heavy hydrocarbon charge stocks such as vacuum gas oils. Large hydrocarbon molecules, associated with the heavy hydrocarbon feed, are cracked to break the large hydrocarbon chains thereby producing lighter hydrocarbons. These lighter hydrocarbons are recovered as product and can be used directly or further processed to raise the octane barrel yield relative to the heavy hydrocarbon feed. 
     The basic equipment or apparatus for the fluidized catalytic cracking (hereinafter FCC) of hydrocarbons has been in existence since the early 1940&#39;s. The basic components of the FCC process include a reactor, a regenerator and a catalyst stripper. The reactor includes a contact zone where the hydrocarbon feed is contacted with a particulate catalyst and a separation zone where product vapors from the cracking reaction are separated from the catalyst. Further product separation takes place in a catalyst stripper that receives catalyst from the separation zone and removes entrained hydrocarbons from the catalyst by counter-current contact with steam or another stripping medium. The FCC process is carried out by contacting the starting material whether it be vacuum gas oil, reduced crude, or another source of relatively high boiling hydrocarbons with a catalyst made up of a finely divided or particulate solid material. The catalyst is transported like a fluid by passing gas or vapor through it at sufficient velocity to produce a desired regime of fluid transport. Contact of the oil with the fluidized material catalyzes the cracking reaction. During the cracking reaction, coke will be deposited on the catalyst. Coke is comprised of hydrogen and carbon and can include other materials in trace quantities such as sulfur and metals that enter the process with the starting material. Coke interferes with the catalytic activity of the catalyst by blocking active sites on the catalyst surface where the cracking reactions take place. Catalyst is transferred from the stripper to a regenerator for purposes of removing the coke by oxidation with an oxygen-containing gas. An inventory of catalyst having a reduced coke content, relative to the catalyst in the stripper, hereinafter referred to as regenerated catalyst, is collected for return to the reaction zone. Oxidizing the coke from the catalyst surface releases a large amount of heat, a portion of which escapes the regenerator with gaseous products of coke oxidation generally referred to as flue gas. The balance of the heat leaves the regenerator with the regenerated catalyst. The fluidized catalyst is continuously circulated from the reaction zone to the regeneration zone and then again to the reaction zone. The fluidized catalyst, as well as providing a catalytic function, acts as a vehicle for the transfer of heat from zone to zone. Catalyst exiting the reaction zone is spoken of as being spent, i.e., partially deactivated by the deposition of coke upon the catalyst. Specific details of the various contact zones, regeneration zones, and stripping zones along with arrangements for conveying the catalyst between the various zones are well known to those skilled in the art. 
     The rate of conversion of the feedstock within the reaction zone is controlled by regulation of the temperature of the catalyst, activity of the catalyst, quantity of the catalyst (i.e., catalyst to oil ratio) and contact time between the catalyst and feedstock. The most common method of regulating the reaction temperature is by regulating the rate of circulation of catalyst from the regeneration zone to the reaction zone which simultaneously produces a variation in the catalyst to oil ratio as the reaction temperatures change. That is, if it is desired to increase the conversion rate an increase in the rate of flow of circulating fluid catalyst from the regenerator to the reactor is effected. Since the catalyst temperature in the regeneration zone is usually held at a relatively constant temperature, significantly higher than the reaction zone temperature, any increase in catalyst flux from the relatively hot regeneration zone to the reaction zone effects an increase in the reaction zone temperature. 
     The hydrocarbon product of the FCC reaction is recovered in vapor form and transferred to product recovery facilities. These facilities normally comprise a main column for cooling the hydrocarbon vapor from the reactor and recovering a series of heavy cracked products which usually include bottom materials, cycle oil, and heavy gasoline. Lighter materials from the main column enter a concentration section for further separation into additional product streams. 
     As the development of FCC units has advanced, temperatures within the reaction zone were gradually raised. It is now commonplace to employ temperatures of about 525° C. (975° F.). At higher temperatures, there is generally a loss of gasoline components as these materials crack to lighter components by both catalytic and thermal mechanisms acting independently. At 525° C., it is typical to have 1% of the potential gasoline components thermally cracked into lighter hydrocarbon gases. As temperatures increase, to say 1025° F. (550° C.), most feedstocks can lose up to 6% or more of the gasoline components to thermal cracking. 
     One improvement to FCC units, that has reduced the product loss by thermal cracking, is the use of riser cracking. In riser cracking, regenerated catalyst and starting materials enter a pipe reactor and are transported upward by the expansion of the gases that result from the vaporization of the hydrocarbons, and other fluidizing mediums if present upon contact with the hot catalyst. Riser cracking provides good initial catalyst and oil contact and also allows the time of contact between the catalyst and oil to be more closely controlled by eliminating turbulence and backmixing that can vary the catalyst residence time. An average riser cracking zone today will have a catalyst to oil contact time of 1 to 5 seconds. A number of riser reaction zones use a lift gas as a further means of providing a uniform catalyst flow. Lift gas is used to accelerate catalyst in a first section of the riser before introduction of the feed and thereby reduces the turbulence which can vary the contact time between the catalyst and hydrocarbons. 
     In most reactor arrangements, catalysts and conversion products still enter a large chamber for the purpose of initially disengaging catalyst and hydrocarbons. The large open volume of the disengaging vessel exposes the hydrocarbon vapors to turbulence and backmixing that continues catalyst contact for varied amounts of time and keeps the hydrocarbon vapors at elevated temperatures for a variable and extended amount of time. Thus, thermal cracking can again be a problem in the disengaging vessel. A final separation of the hydrocarbon vapors from the catalyst is performed by cyclone separators that use centripetal acceleration to disengage the heavier catalyst particles from the lighter vapors which are removed from the reaction zone. 
     In order to minimize thermal cracking in the disengaging vessel, a variety of systems for directly connecting the outlet of the riser reactor to the inlet of a cyclone are suggested in the prior art. Directly connecting the cyclone inlet to the riser outlet in what has been termed a “direct coupled cyclone system” requires a means for relieving pressure surges that can otherwise overload the cyclones and cause catalyst to be carried over into the product stream separation facilities located downstream of the reactor. The development of these systems to handle the overload problem in a variety of ways increases the practicality of directly coupling the riser outlet to the cyclone inlet. Direct coupling of cyclones can greatly reduce thermal cracking of hydrocarbons. 
     It is also known, for purposes of controlling thermal cracking, to lower the temperature of the reaction products upon leaving the cyclone separators by the use of a quench liquid. Quenching the product stream reduces its temperature below that at which thermal cracking can occur and reduces the loss of gasoline products by continued cracking to light ends. 
     DISCLOSURE STATEMENT 
     U.S. Pat. No. 4,624,771, issued to Lane et al. on Nov. 25, 1986, discloses a riser cracking zone that uses fluidizing gas to pre-accelerate the catalyst, a first feed introduction point for injecting the starting material into the flowing catalyst stream, and a second downstream fluid injection point to add a quench medium to the flowing stream of starting material and catalyst. 
     U.S. Pat. No. 4,624,772, issued to Krambeck et al. on Nov. 25, 1986, discloses a closed coupled cyclone system that has vent openings, for relieving pressure surges, that are covered with weighted flapper doors so that the openings are substantially closed during normal operation. 
     U.S. Pat. No. 4,234,411, issued to Thompson on Nov. 18, 1980, discloses a reactor riser disengagement vessel and stripper that receives two independent streams of catalyst from a regeneration zone. 
     U.S. Pat. No. 4,479,870, issued to Hammershaimb et al. on Jun. 30, 1984, and U.S. Pat. No. 4,822,761, issued to Walters et al. on Apr. 18, 1989, teach the use of lift gas having a specific composition in a riser conversion zone at a specific set of flowing conditions with the subsequent introduction of the hydrocarbon feed into the flowing catalyst and lift gas stream. 
     U.S. Pat. No. 3,133,014 shows the use of a spray nozzle in a reactor vapor line to cool high boiling hydrocarbons and prevent the formation of coke deposits on the vapor line wall. 
     U.S. Pat. Nos. 3,290,465; 4,263,128; 4,256,567, and 4,243,514 generally teach the use of quench streams for the purpose of preventing thermal cracking of hydrocarbons in transfer lines. 
     U.S. Pat. Nos. 3,221,076 and 3,238,271 show the direct transfer of vapors from a cyclone separator in a reaction vessel to a contacting vessel for quenching or removing fine catalyst particles that are transported with vapors. 
     BRIEF DESCRIPTION OF THE INVENTION 
     It is an object of this invention to provide an FCC apparatus that improves the control of contact time between catalyst and hydrocarbons. 
     It is a further object of this invention to provide an FCC apparatus that operates with high reaction temperatures and decreases thermal stresses in FCC structure due to temperature gradients. 
     It is a yet further object of this invention to provide an FCC apparatus having reduced times of contact between the catalyst and hydrocarbons, and reduced exposure of the hydrocarbon feeds to elevated temperature exposure. 
     It is another object of this invention to provide an FCC apparatus that will facilitate the separation of catalyst and hydrocarbon vapors. 
     It is a yet another object of this invention to improve the recovery of cracked hydrocarbon products from the disengagement zone and stripper section of the reaction process. 
     These and other objects are achieved by the process of this invention which is an FCC apparatus that converts FCC feed by contact with catalyst in a riser conversion zone, maintains a carefully contact time between the catalyst and hydrocarbon feed, and rapidly quenches hydrocarbon products recovered from the cyclone separators to avoid thermal cracking. This apparatus of this invention places a quench chamber above a reactor vessel and a hot stripper below a reactor vessel to provide a progressively decreasing temperature profile up the structure of the FCC arrangement and equipment for sequential reaction control. A riser contains the primary catalytic reactions of the hydrocarbon vapor and delivers the reacted vapors to the reactor structure. Starting from the bottom of the structure the hot stripper has the highest temperature and desorbs or displaces hydrocarbons from the catalyst to terminate long residence time catalytic reactions. Above the hot stripper bulk separation equipment divides the main vapor and catalyst stream to limit residence time of major catalytic reactions. At a yet higher elevation and lower internal temperature quench equipment arrests thermal reactions of the vapor stream. This structure arrangement permits reliable control of reaction time to obtain desired products and enhances mechanical reliability of the structure. 
     In addition the progressively decreasing temperature gradient lowers thermally induced stresses in the shells of the vessels that make up the structure. In normal operation the stripping vessel will operate at the highest temperature. A reactor vessel housing means for making an initial separation between the catalyst and the hydrocarbon vapors will operate a lower temperature than the stripping vessel. Finally, the quench vessel that cools the product vapors will operate at the lowest temperature. Connecting a reactor vessel on top of a hot stripping vessel and a quench vessel on top of a hot stripping vessel provides a uniformly decreasing temperature profile up the structure of the reactor, stripper and quench vessels. This uniformly changing temperature gradient through lowers thermally induced stresses. 
     Accordingly, in one embodiment this invention is an apparatus for the fluidized catalytic cracking of hydrocarbons. The apparatus includes a riser portion that comprises a substantially vertical riser, means for introducing catalyst into a lower portion of the riser, means for introducing a hydrocarbon feed into the riser and a transfer conduit in communication with the upper end of the riser. The invention also incorporates means for separating catalyst from gases. The means for separating define an inlet in closed communication with the conduit, a catalyst outlet, and a vapor outlet and are at least partially located in the reactor vessel. A stripping vessel located below the reactor vessel communicates with the catalyst outlet and defines a substantial collection volume for receiving catalyst separated by the means for separating catalyst. The stripping vessel also contains means for contacting the catalyst collected therein with a stripping medium and means for heating catalyst in said stripper vessel. A vapor line carries hydrocarbon vapors away from the vapor outlet and into means for quenching the hydrocarbon vapors. The means for quenching have a location above reactor vessel. 
     In an alternate and more limited embodiment of this invention the apparatus of this invention comprises a reactor vessel having a center line and a substantially vertical riser having a center line horizontally offset from the reactor vessel. A catalyst nozzle communicates with a lower part of the riser for introducing catalyst into a lower portion of the riser. A lift gas nozzle in communicates with a lower portion of the riser at a location above the catalyst nozzle for introducing a lift gas into a lower portion of the riser. A feed nozzle in communicates with the riser at a location above the lift gas nozzle for introducing a hydrocarbon feed into the riser. A transfer conduit defines a conduit outlet and a conduit inlet in communication with the upper end of the riser. Means for separating catalyst from gases are located in the reactor vessel. The means for separating define a separation inlet in closed communication with the conduit outlet, and a catalyst outlet and a vapor outlet. A stripping vessel, located below the reactor vessel and in communication with the catalyst outlet, has a substantial collection volume for receiving catalyst separated by the means for separating catalyst, and includes means for contacting the catalyst collected therein with a stripping medium and heating catalyst in the stripper vessel. A gas tube has one end in communication with the stripping vessel and a second end in communication with the transfer conduit. A vapor line is in communication with the vapor outlet for carrying hydrocarbon vapors away from the vapor outlet. A quench vessel is located on top of the reactor vessel for quenching hydrocarbon vapors from the vapor line. 
     In another limited embodiment this invention is an apparatus for the fluidized catalytic cracking of hydrocarbons. The apparatus comprises: a reactor vessel; a substantially vertical riser extending coaxially into the reactor vessel; a catalyst nozzle in communication with the riser for introducing catalyst into a lower portion of the riser; a lift gas nozzle in communication with the riser for introducing a lift gas into a lower portion of the riser; a feed nozzle in communication with the riser and located above the lift gas nozzle for introducing a hydrocarbon feed into an upper portion of the riser; a disengaging vessel surrounding the upper end of the riser for separating catalyst from hydrocarbon vapors; a collector located at the upper end of the riser in the disengaging vessel; a transfer conduit in communication with the collector; a cyclone separator defining an inlet in closed communication with the conduit, a catalyst outlet, and a vapor outlet; a stripping vessel located below the reactor vessel and in communication with the catalyst outlet the stripping vessel having a substantial collection volume for receiving catalyst from the catalyst outlet and means for contacting the catalyst collected therein with a stripping medium and heating catalyst in the stripping vessel; a vapor line in communication with the vapor outlet; and, means for quenching vapors withdrawing from the reactor vessel by the vapor line. 
     Other aspects and embodiments and advantages of this invention are disclosed in the following detailed description of the invention. 
    
    
     BRIEF DESCRIPTION OF THE DRAWINGS 
     FIG. 1 is a schematic elevation showing a cross-section an FCC reactor suitable for the practice of this invention along with an FCC regenerator. 
     FIG. 2 is a cross section of an alternate reactor vessel arrangement suitable for use in this invention. 
    
    
     DETAILED DESCRIPTION OF THE INVENTION 
     The apparatus of this invention will be described with references to the drawings. These references are not meant to limit the process or the apparatus to the particular details of the drawing disclosed in conjunction therewith. Looking first at the operation of the riser conversion zone, a lift gas stream  10  enters an inlet conduit  12  that passes the lift gas into the lower portion of a riser  14 . Hot catalyst from a regenerated standpipe  16  passes through a control valve  18  and is mixed with the lift gas in a junction between the standpipe and lower riser generally referred to as a Y-section and denoted as conduit  20  in FIG.  1  and including a catalyst nozzle. Lift gas carries the catalyst up the riser from lower section  14  to upper riser section  22  and conditions the catalyst by contact therewith. Between the upper and lower riser section, feed nozzles  24  inject hydrocarbon feed into the flowing stream of catalyst and lift gas. Hydrocarbon feed is converted as it travels to the end  26  of the riser. At the top  26 , the riser ends with an abrupt change of direction that directs the mixture of converted feed components and catalyst into transfer conduit  28 . FIG. 1 depicts the use of an external riser where the entire length of the riser is located outside of the reactor vessel. 
     The catalysts which enter the riser and can be used in the process of this invention include those known to the art as fluidizing catalytic cracking catalysts. These compositions include amorphous clay type catalysts which have for the most part been replaced by high activity crystalline alumina silicate or zeolite containing catalysts. Zeolite catalysts are preferred over amorphous type catalysts because of their higher intrinsic activity and their higher resistance to the deactivating effects of high temperature exposure to steam and exposure to the metals contained in most feedstocks. Zeolites are the most commonly used crystalline alumina silicates and are usually dispersed in a porous inorganic carrier material such as silica, aluminum, or zirconium. These catalyst compositions may have a zeolite content of 30% or more. 
     Feeds suitable for processing by this invention, include conventional FCC feedstocks or higher boiling hydrocarbon feeds. The most common of the conventional feedstocks is a vacuum gas oil which is typically a hydrocarbon material having a boiling range of from 343-552° C. and is prepared by vacuum fractionation of atmospheric residue. Such fractions are generally low in coke precursors and heavy metals which can serve to deactivate the catalyst. 
     This invention is also useful for processing heavy or residual charge stocks, i.e., those boiling above 500° C. (930° F.) which frequently have a high metals content and which usually cause a high degree of coke deposition on the catalyst when cracked. Both the metals and coke serve to deactivate the catalyst by blocking active sites on the catalyst. Coke can be removed, to a desired degree, by regeneration and its deactivating effects overcome. Metals, however, accumulate on the catalyst and poison the catalyst by fusing within the catalyst and permanently blocking reaction sites. In addition, the metals promote undesirable cracking thereby interfering with the reaction process. Thus, the presence of metals usually influences the regenerator operation, catalyst selectivity, catalyst activity, and the fresh catalyst make-up required to maintain constant activity. The contaminant metals include nickel, iron and vanadium. In general, these metals affect selectivity in the direction of less gasoline and more coke. Due to these deleterious effects, the use of metal management procedures within or before the reaction zone are anticipated when processing heavy feeds by this invention. Metals passivation can also be achieved to some extent by the use of appropriate lift gas in the upstream portion of the riser. 
     The finely divided regenerated catalyst entering the bottom of a reactor riser leaves the regeneration zone at a high temperature. Where the riser is arranged vertically, the bottom section will be the most upstream portion of the riser. In most cases, the riser will have a vertical arrangement, wherein lift gas and catalyst enter the bottom of the riser and converted feed and catalyst leave the top of the riser. Nevertheless, this invention can be applied to any configuration of riser including curved and inclined risers. The only limitation in the riser design is that it provide a substantially smooth flow path over its length. 
     Where employed, contact of the hot catalyst entering the riser with a lift gas accelerates the catalyst up the riser in a uniform flow regime that will reduce backmixing at the point of feed addition. Reducing backmixing is important because it varies the residence time of hydrocarbons in the riser. Addition of the lift gas at a velocity of at least 1.8 meters per second is necessary to achieve a satisfactory acceleration of the catalyst. The lift gas used in this invention is more effective when it includes not more than 10 mol % of C 3  and heavier olefinic hydrocarbons and is believed to selectively passivate active metal contamination sites on the catalyst to reduce the hydrogen and coke production effects of these sites. Selectively passivating the sites associated with the metals on the catalyst leads to greater selectivity and lower coke and gas yield from a heavy hydrocarbon charge. Some steam may be included with the lift gas and, in addition to hydrocarbons, other reaction species may be present in the lift gas such as H 2 , H 2 S, N 2 , CO, and/or CO 2 . However, to achieve maximum effect from the lift gas, it is important that appropriate contact conditions are maintained in the lower portion of the riser. A residence time of 0.5 seconds or more is preferred in the lift gas section of the riser, however, where such residence time would unduly lengthen the riser, shorter residence times for the lift gas and catalyst may be used. A weight ratio of catalyst to hydrocarbon in the lift gas of more than 80 is also preferred. 
     After the catalyst is accelerated by the lift gas, it enters a downstream portion of the riser which is generally referred to as the upper section. Feed may be injected into the start of the section by nozzles as shown in the FIGS. 1 and 2 or any device that will provide a good distribution of feed over the entire cross-section of the riser. Atomization of the feed, as it enters the riser, promotes good distribution of the feed. A variety of distributor nozzles and devices are known for atomizing feed as it is introduced into the riser. Such nozzles or injectors may use homogenizing liquids or gas which are combined with the feed to facilitate atomization and dispersion. Steam or other non-reactive gases may also be added with the feed, for purposes of establishing a desired superficial velocity up the riser. The superficial velocity must be relatively high in order to produce an average residence time for the hydrocarbons in the riser of less than 5 seconds. Shorter residence times permit the use of higher reaction temperatures and provide additional benefits as discussed below; thus where possible the feed has a residence time of 2 seconds or less. In more limited embodiments of this invention, the residence time may be as low as 0.1 second and in some cases as low as 0.05 seconds. 
     The catalyst and feed mixture has an average temperature of at least 520° C. (970° F.). Higher temperatures for the catalyst and feed mixture are preferred with temperatures of 540° C. (1000° F.) and 550° C. (1025° F.) being particularly preferred. The combination of a short residence time and higher temperatures in the riser shifts the process towards primary reactions. These reactions favor the production of gasoline and tend to reduce the production of coke. Furthermore, the higher temperatures raise gasoline octane. The short catalyst residence time within the riser is also important for maintaining the shift towards primary reactions and removing the hydrocarbons from the presence of the catalyst before secondary reactions that favor coke production have time to occur. The ability to carefully limit residence time also permits the cessation of cracking reactions to produce higher boiling range products where desired. 
     The high velocity stream of catalyst and hydrocarbons is then rapidly separated at the end of the riser. This can be accomplished in a number of ways. FIG. 1 shows one arrangement where the catalyst and hydrocarbons pass directly into a cyclonic separation system or the riser can be configured so as to abruptly change direction before this initial separation. Following separation, the separated vapors begin their path toward the product recovery zone while the separated catalyst is directed toward the stripping zone. 
     The catalyst and hydrocarbon stream carried from the riser by transfer conduit  28  can be diluted by the injection of a suitable diluent through a diluent conduit  30 . The diluent is mixed with the hydrocarbons and catalyst as they progress through conduit  28 . Horizontally arranged transfer conduit  28  carries the hydrocarbons and vapor into a reactor vessel  29 . Slightly farther downstream in conduit  28 , a stream of separated hydrocarbons, as hereinafter described, enters the top of conduit  28  through a tube  32  which is connected to conduit  28  just ahead of the inlet of a first cyclone separator  34 . Hydrocarbon vapor, catalyst, and diluent, when present, pass directly into cyclone separator  34  where separation of catalyst and product vapors occurs. Separator  34  discharges catalyst downwardly through a dip leg  36  and into a hereinafter described stripping zone, while hydrocarbon vapors and small amounts of entrained catalyst are carried from the top of separator  34  through a cross-over conduit  38  and into a second cyclone separator  40 . Cross-over conduit  38  contains an optional weighted flapper door  41  for relieving pressure surges. Cyclone separator  40  performs a more complete separation to recover additional catalyst still entrained in the product vapor. Additional amounts of recovered catalyst are downwardly discharged through a dip leg  42  while hydrocarbon vapors having a very low loading of catalyst particles exit the top of the cyclone through an outlet conduit  44 . 
     The diluent that enters transfer conduit  28  will usually comprise steam. Adding diluent ahead of the separation devices lowers the partial pressure of the hydrocarbons as they enter the cyclones. As the catalyst and hydrocarbons pass into the transfer conduits and through the separation devices, turbulence will vary the residence time of the hydrocarbons in these various devices. Therefore, the addition of diluent at this point, to lower the partial pressure of the hydrocarbons, attenuates the effects of catalytic and thermal cracking. Thus, initial contact with a diluent ahead of the cyclones prevents the loss of product by overcracking. Suppressing cracking reactions by the addition of diluent also allows the reaction time to be controlled. As a result, hydrocarbon reactions occur mainly in the riser and, as previously mentioned, can be limited to a short time. Short reaction times again favor the preferred primary reaction mechanism. Reactions that yield the desired distillate and gasoline products are primary reactions that occur quickly. Coke producing secondary reactions, primarily the polymerization and condensation of polycyclic compounds, over the acid catalyst, are secondary reactions that take longer to occur. The polycyclic compounds that combine in these secondary reactions are first generated by primary reactions such as naphthene cracking and the dealkylation of side chains. It is believed that by careful control, a short reaction time allows the primary reactions to occur while preventing most of the secondary reactions. Therefore, the addition of a diluent can increase the production of distillate and raise the quantity and octane of the gasoline product. 
     However, the addition of diluent through conduit  30  must be limited to avoid condensation of heavier hydrocarbon components in the cyclone separators or transfer conduits and excessive cooling of the catalyst. For this purpose, the temperature of the combined catalyst and hydrocarbon stream should not be reduced below the dew point of the heavier species. 
     Hydrocarbons separated from the catalyst in a manner hereinafter described are returned to the cyclones to remove any entrained catalyst that may accompany it back into the transfer conduit. For this purpose, the lower end of tube  32  is shown in open communication with the interior of reactor vessel  29 . In order to pass hydrocarbons from vessel  29  back into the transfer conduit, a positive pressure must be maintained that will provide the necessary driving force. In order to regulate the pressure drop, these hydrocarbons are transferred back into the transfer conduit through an extended length of gas tube  32 . High gas velocities should be avoided since they can impart momentum to the catalyst that will erode the transfer conduit. Gas tube  32  is arranged to direct catalyst into the top of the transfer conduit. The top has the advantage of placing any gas jet developed by the entry of gas into the transfer conduit across the vertical dimension of the transfer conduit which is usually larger than the width of the conduit. 
     Both tube  32  and diluent conduit  30  also inject gas into the upper surface of the transfer conduit in order to keep catalyst, that tends to flow along the bottom of the conduit, away from the outlets of tubes  32  and conduit  30 . 
     In FIG. 1, transfer conduit  28  communicates the catalyst and hydrocarbons with the cyclones that are located within reactor vessel  29 . The careful control of reaction times requires that catalyst be communicated in as direct a fashion as possible to the separation device. The transfer conduit and cyclone arrangement of the FIG. 1 differs from a number of those commonly used in the prior art by the direct connection of the transfer conduit to the inlet of cyclone  34 . For this reason, transfer conduit  28  can be described as a closed conduit notwithstanding the presence of tube  32  and diluent conduit  30 . It is possible to alter the arrangement of Figure to minimize the volume of the reactor vessel by using cyclone separators that are designed to withstand the internal pressure of the product stream and locating any additional stages of cyclone separators outside of the reactor vessel and discharging separated catalyst from external cyclones back into the stripper vessel. 
     For the most part, cyclones  34  and  40  are of a conventional design but will generally have a larger capacity, at least in separator  34 , for accommodating the larger volume of solids and gases that will enter the cyclones because of the direct coupling of the separator inlet to the transfer conduit. For those units where instabilities in operation, caused by such things as interruption in the flow of catalyst into the riser or the occasional injection of large amounts of water, will cause pressure surges in the riser, provision should be made to prevent these surges from overloading the cyclones. When the cyclone is overloaded, the spiralling effect of the flow through the cyclone that separates particles from fluid, is interrupted and the cyclone begins to act as a simple conduit transferring large amounts of catalyst out of the top of the cyclone with the converted products. Pressure surges, at least in part, can be relieved by venting the cross-over conduit  38  between the two cyclones. 
     A preferred method of venting uses a flapper door  42 . Flapper door  42  covers an opening on the cross-over conduit that is used for venting excessive pressure from the cyclone and preventing overloading of cyclone  40  when cyclone  34  becomes overloaded with catalyst. Door  42  is weighted to minimize leakage during periods of normal operation when it is not opened by internal pressure in the cross-over conduit. The higher operating pressure inside the reactor vessel also tends to keep door  42  closed. Door  42  can be weighted or alternately counter-balanced such that it will open at a predetermined pressure difference between the internal pressure of cross-over conduit  38  and the reactor pressure outside the conduit. In this case, the venting of cross-over conduit  38  will only protect cyclone separator  40 , generally referred to as a secondary cyclone, from overloading. It is expected that during the venting operation the amount of catalyst particles leaving the secondary cyclone through conduit  44  will increase, however, this increase for a short period of time will not impair operation of the downstream separation facilities. A similar type vent can be provided on the portion of the transfer conduit located within vessel  29  to also protect cyclone separator  34  from catalyst overload. Additional details on the direct coupling of a riser to cyclones and for protecting the cyclones against overload can be obtained from the previously mentioned prior art. 
     Dip legs  36  and  42  discharge recovered catalyst into a catalyst stripping section. In the embodiment of the Drawing, dip legs  36  and  42  discharge the catalyst into a relatively dense bed  46  of catalyst particles having an upper bed level  48 . 
     An important element of this invention is the use of a hot catalyst stripping zone. The term “hot catalyst stripping zone” refers to a stripper having a temperature above at least 970° F. Greater advantages are obtained when the stripper is maintained above 1000° F. The high temperature riser operation provides high temperature catalyst that in turn keeps the stripper hot. In many instances, hot catalyst from the separator will have sufficient heat to maintain the necessary stripper temperature. 
     Where a higher stripper temperature than can be obtained from the riser catalyst is desired, any suitable method may be used to heat the catalyst within the stripping zone. Acceptable methods include the use of heat transfer tubes, controlled oxidation of hydrocarbons in the stripper as well as direct and indirect transfer of heat from regenerated catalyst. One form of indirect heat transfer, to raise the temperature of the spent catalyst, can use a catalyst to catalyst heat exchanger within the stripper that circulates hot catalyst from the regenerator through heat exchange tubes and back to the regenerator in a closed system. 
     FIG. 1 shows another approach for heating the catalyst wherein a continuous stream of hot catalyst particles taken from a regenerator  72  by a reheat conduit  50  in an amount regulated by a control valve  52  enters a stripper riser  54 . A lift medium, such as steam, from a conduit  56  lifts hot catalyst from the bottom of riser  54 . Hot regenerated catalyst particles flow out of the upper end of riser  54  and contact a baffle  58  that redirects the catalyst downward into bed  46 . The hot regenerated catalyst heats the spent catalyst particles in bed  46  which are then transferred downward into a stripping vessel  60  having a series of baffles  62  for counter-currently contacting the downward flowing catalyst particles with a stripping medium, such as steam, that enters the stripping zone through a conduit  64 . A distributor  66  distributes the stripping medium over the cross-section of the stripping vessel  60 . Stripped hydrocarbon vapors, as well as stripping medium, rise upwardly through bed  46  and enter the bottom of tube  32  for return to the cyclone separators in the manner previously described. Stripped and fresh catalyst particles are taken from the stripper  60  by a spent catalyst standpipe  68 , in an amount regulated by a control valve  70 , and transferred to regenerator  72  for the oxidative removal of coke from its surface. 
     Catalyst entering the stripper is kept hot to remove additional hydrocarbons from the spent catalyst by vaporizing the higher boiling hydrocarbons from the surface of the catalyst. Since the commonly employed zeolite catalysts can act as an effective adsorbent, a large quantity of hydrocarbons can be absorbed on the surface of the catalyst. Although heating the catalyst will also tend to raise temperatures and again may promote some thermal cracking, any hydrocarbons that remain absorbed on the catalyst are lost by combustion in the regeneration zone. Thus, some small loss to thermal cracking in the stripping zone is preferable to the larger loss of adsorbed product which may be burned in the regenerator. 
     Any catalyst introduced into the stripper for the purpose of heating should be taken from the hottest section of the regenerator in order to minimize the amount of hot catalyst introduced therein. Although the hot clean catalyst is favored as a heating medium due to its high heat capacity and ready availability, the regenerated catalyst can also act as a clean adsorbent which, if introduced in large quantities, can absorb more additional hydrocarbons than the heat released thereby will desorb from the spent catalyst. Therefore, it is preferable to take relatively small amounts of hot regenerated catalyst from the regenerator for the purpose of heating catalyst in the stripper. 
     Spent catalyst taken from stripper  60  through spent catalyst standpipe  68  enters regenerator  72  for the oxidative removal of coke from the surface thereof. A conduit  76  conveys compressed air into a distributor grid  78  that distributes the air over the cross-section of a lower regenerator vessel  80 . Regenerated catalyst is carried by a recirculation conduit  82  into lower regenerator vessel  80  and mixed with air from distributor  78  and spent catalyst from conduit  68 . Combustion of coke deposits begins as oxygen reacts with coke at the elevated temperature of the catalyst and air mixture. Air and combustion gas carry the catalyst and gas mixture upward into regenerator riser  84 . A riser arm  86  having an opening  88  directs the catalyst and gas mixture downward to at least partially disengage gases from the catalyst. The gas mixture plus any entrained catalyst flow upwardly and are collected by cyclone separators  90 . A plenum  92  collects combustion gas from the cyclone separators for removal from the regenerator through a nozzle  94 . Catalyst recovered from the cyclone separators is discharged through conduits  96  where it is collected by a cone  98  along with catalyst that was initially disengaged by discharge through opening  88 . The regenerated catalyst conduit  16  returns regenerated catalyst from cone  98  to riser  14 , as previously described. Hot catalyst for reheat conduit  50  is also withdrawn from standpipe  50 . Other details and variations on the operation of an FCC regenerator are well known by those skilled in the art. 
     Looking again at the reactor, converted hydrocarbons that leave separator  40  through conduit  44  undergo quick quenching to avoid thermal cracking. In order to prevent thermal cracking, these vapors will preferably be quenched to a temperature below about 500° C. Quenching may be accomplished by the injection or contact of the vapor stream with a suitable quench fluid. Quench mediums that can be used include light oil, steam, water or heavy oil. When using light oil, steam or water, care must be taken to avoid condensation of higher boiling compounds on the walls of the piping leading to the product separation facilities. These lighter compounds are either used in or easily converted to the gas phase as these light quench materials rapidly cool the higher boiling components of the product stream. The resulting large concentration of gas in the quench stream may not adequately flush coke condensible compounds from the transfer piping. Heavy quench liquids are preferred since they prevent coke accumulation by providing a large volume of liquid wash. 
     Quench liquid may be injected into the converted hydrocarbons using spray nozzles, showered head injection or staged injection of two or more quench mediums. The quench may be added directly to the cyclone outlets or to a manifold or plenum chamber that collects the hydrocarbon vapors from several cyclone outlets. Thus, the quench vessel can comprise a section of piping or a conduit through which the quench and product vapors pass. 
     FIG. 1 shows an alternate form of incorporating the quench medium that uses a liquid contacting zone. Substantial advantages are achieved in the quench operation when it employs a liquid contacting zone as shown in FIG.  1 . In this type of quench apparatus the quench conduit  44  carries product vapor from each cyclone separator  40  directly into a quench chamber  100 . Quench chamber  100  is separated from the reactor by a partition  111 . Product vapors entering quench chamber  100  will normally have a temperature in the range of from 480-565° C. (900-1050° F.). These vapors leave the end of conduit  44  and travel around an end cover  102 . The purpose of end cover  102  is to prevent the quench liquid, as hereinafter described, from spilling back into the conduit  44 . In a first series of contacting trays comprising heat removal trays  104 , the rising hydrocarbon vapors are contacted by the quench liquid. Heat removal trays  104  are preferably disc and donut trays. At the top of the heat removal trays, a quench liquid is introduced by an extended distributor  106 . The quench is preferably a heavy hydrocarbon having a boiling point range of 290-600° C. (550-1100° F.). A portion of the liquid quench may also be introduced through nozzle  108  below a liquid level  110  at the bottom of the quench chamber to independently control the temperature of the collected liquid. By the addition of quench liquid, the temperature of the collected liquid may be kept below 400° C. (750° F.) or preferably below 370° C. (700° F.). Maintaining the quench liquid below 400° C. prevents the small degree of hydrocarbon cracking which might otherwise occur at higher temperatures and adversely affect the flash point of the bottoms product. This quench material is generally described as a main column bottoms stream which is obtained from the separation facilities for the product stream and will normally include a slurry of catalyst particles. In new FCC units that use high efficiency cyclones, the main column bottoms typically carries about 0.01 to 0.05 wt. % catalyst and other insolubles, but can have solids concentrations as high as 0.15 to 0.2. Older FCC units using a slurry settler will have a much higher wt. % of particulates averaging about 1 to 2%. This quench will usually enter the quench chamber at a temperature in the range of 230-345° C. (450-650° F.). A nozzle  112  withdraws liquid quench from the bottom of chamber  100 . The nozzle  112  has a location well below the top discharge conduits  44  and should be located as low as possible in the quench chamber in order to keep the full volume of quench liquid in circulation. For this reason, it is also preferable to have several withdrawal nozzles spaced about the circumference of the quench chamber. Temperature of the liquid quench as it is withdrawn through nozzle  112  will be between 315-400° C. (600-750° F.). After removal, the quench is normally passed through heat exchange equipment to lower its temperature and pumped back to distributor  106  for return to the top of heat removal trays and to the bottoms quench nozzle  108 . The product vapors will also contain a certain amount of heavy material having a boiling point above the entering temperature of the quench medium which will collect and increase the total volume of the quench liquid. Therefore, a portion of the circulating quench medium is withdrawn continuously as heavy oil product to keep the liquid level  110  below the top of conduit  44 . 
     The quench chamber may contain additional contacting trays which receive the lighter product vapors that have risen above trays  104  and are contacted by a hydrocarbon reflux stream that is relatively lighter than the quench medium passed over trays  104 . In its preferred form, a second series of contacting trays comprising fractionation trays  116  receive the ascending product vapors while an extended distributor  118  delivers a hydrocarbon reflux stream to the top of the fractionation trays that flows counter-currently to the rising vapors. It is preferred that the reflux stream be a heavy cycle oil having a boiling range of 230-400° C. (450-750° F.). As the product vapor enters the fractionation trays, it will usually have a temperature between 275-400° C. (525-750° F.). In the case of heavy cycle oil addition, this will usually enter the fractionation trays at a temperature in the range of 260-320° C. (500-600° F.). The relatively cool vapors are collected at the top of quench chamber  100  and withdrawn through a nozzle  120 . The vapors are carried overhead via line  122  to additional separation facilities for further separation into the various components of the product slate. 
     Quench chamber  100  and the cyclones are supported from the top of the reactor vessel. In this type of arrangement proper design of partition  111  and discharge conduit  44  is important to the operation of the apparatus of this invention. Partition  111  is designed to withstand a liquid loading on its upper side and a pressure loading on its lower side. The pressure loading results from the higher pressure employed in the reactor vessel relative to the quench chamber provides a driving force for transferring vapors to the quench chamber. The hemispherical shape of partition  111 , as shown in the drawing, serves two objectives, one is to withstand the pressure loading on its bottom side when it is greater than the liquid loading on the top side of the partition and to facilitate removal of the bottoms liquid by forming a channel towards the outer periphery of the dome shaped partition. although any shape of partition can be used, it is preferable to avoid a partition that is concave to the quench chamber since this will form a stagnant area of hydrocarbon vapors in upper reactor portion. 
     Contact of partition  111  with the relatively cool quench liquid on its upper side cools the partition. If the product vapors are allowed to come in contact with the cooled surface, this will promote condensation of the relatively heavy hydrocarbons and the accumulation of coke on the lower surface of the partition. For this reason, a layer of an insulating ceramic material is usually used to cover the entire lower surface of partition  111 . This insulating material is composed of an insulating refractory lining having a thickness ranging from 2 to 5 inches depending on the insulating properties of the material. The design and use of such materials is well known to those skilled in the art. Condensation of high boiling product vapors into coke deposits is a similar concern for the discharge conduits  44 . The outer surface of conduit  44  is in contact with liquid from the quench and is cooled thereby. An insulating type refractory lining usually covers the inside of discharge conduit  44 . In the case of conduit  44 , this lining will have a thickness that can vary between 1 to 5 inches depending on the insulating properties of the material. The lining should have a thickness which will keep the surface of the lining that is in contact with the hydrocarbon vapors at a temperature within 9° C. of the vapor temperature in contact therewith. 
     When the quench chamber is incorporated into the top of the reactor, it can replace a portion of the main column that is generally used separating the recovered vapor products from the reactor. A main column will ordinarily contain a quench section. The incorporation of this invention will allow at least the quench system to be removed from the main column. The embodiment of this invention shown in the Drawing also includes the addition of fractionation trays for the rectification of the vapor leaving the heat removal section. Additional fractionation trays, pump around circuits, and withdrawal points may be added to obtain additional product cuts from the quench chamber. 
     Again FIG. 1 demonstrates the use of cyclones for the initial separation of catalyst from hydrocarbon products. Other arrangements for the initial separation of catalyst from hydrocarbons can be used in this invention. One such arrangement is shown in U.S. Pat. No. 5,182,085, the contents of which are hereby incorporated by reference. FIG. 2 demonstrates another embodiment of this invention that does not use cyclones for the initial separation of the catalyst from the product vapors and a reactor riser having an upper end extending inside the reactor vessel. 
     Referring to FIG. 2, regenerated catalyst from a regenerator (not shown) is transferred by a conduit  214 , at a rate regulated by a control valve  216 , to a Y-section  218 . Lift gas injected into the bottom of Y-section  218 , by a conduit  220 , carries the catalyst upward through a lower riser section  222 . Feed is injected into the riser above lower riser section  222  by feed injection nozzles  224 . 
     The mixture of feed, catalyst and lift gas travels up an intermediate section of the riser  226  and into an upper internal riser section  228  that terminates in an upwardly directed outlet end  230  that is located in a dilute phase region  232  of a reactor vessel  234 . The gas and catalyst are separated in dilute phase section  232 . Vapor lines  236  collect gas from the dilute phase section through transfer conduits  237  and transfer it to a collection chamber  238 . From collection chamber  238 , a T-type piping arrangement  240  distributes the gas which still contains a small amount of catalyst particles to a pair of cyclone separators  242 . The T-type piping arrangement includes a single conduit  241  that serves as quench chamber and into which one or more quench lines  243  inject a quench fluid. Cooled and relatively clean product vapors are recovered from the outlets of cyclones  242  by a manifold  244  and withdrawn from the process through an outlet  246 . 
     Catalyst separated by cyclone separators  242  is carried back to reactor vessel  234  by dip pipe conduits  248 . Spent catalyst from dilute phase section  232  and the dip pipe conduits form a dense catalyst bed  250  in a lower portion of the reactor vessel  234 . The dense catalyst bed extends downward into a stripping vessel  252  that operates as a stripping zone. Stripping fluid enters a lower portion of the stripping vessel  252  through a distributor  254  and travels upward through the stripping vessel and reactor vessel in countercurrent flow to the downward moving catalyst. As the catalyst moves downward, it passes over reactor stripping baffles  256  and  258  and stripper baffles  260  and  262  and is transferred into the regenerator by a conduit  264 . 
     The reactor riser of this embodiment of the invention is laid out to perform an initial separation between the catalyst and gaseous components in the riser. In this type of arrangement the end of the riser  230  must terminate with one or more upwardly directed openings that discharge the catalyst and gaseous mixture in an upward direction into a dilute phase section of the reactor vessel. The open end of the riser can be of an ordinary vented riser design as described in the prior art patents of this application or of any other configuration that provides a substantial separation of catalyst from gaseous material in the dilute phase section of the reactor vessel. It is believed to be important that the catalyst is discharged in an upward direction in order to minimize the distance between the outlet end of the riser and the top of the dense phase catalyst bed in the reactor vessel. The flow regime within the riser will influence the separation at the end of the riser. Typically, the catalyst circulation rate through the riser and the input of feed and any lift gas that enters the riser will produce a flowing density of between 3 lbs/ft 3  to 30 lbs/ft 3 , more typically 3 lbs/ft 3  to 20 lbs/ft 3 , and an average vel about 10-100 ft/sec. for the catalyst and gaseous mixture. 
     The manner in which the gaseous vapors are withdrawn from the dilute phase volume of the reactor vessel will influence the initial separation and the degree of re-entrainment that is obtained in the reactor vessel. In order to improve this disengagement and avoid re-entrainment, FIG. 2 shows the use of an annular collector  292  that surrounds the end  230  of the riser. Collector  292  is supported from the top of the reactor vessel  234  by withdrawal conduits  236 . Withdrawal conduits  236  are symmetrically spaced around the annular collector and communicate with the annular collector through a number of symmetrically spaced openings to obtain a balanced withdrawal of gaseous components around the entire circumference of the reactor riser. All of the stripping gas and gaseous components from the reactor riser are withdrawn by annular collector  292 . 
     FIG. 2 shows an arrangement for transferring gases from the conduits  236  to the cyclones that avoids a mal-distribution of the catalyst and gas mixture to the different cyclones. The simplest way to connect the gas conduits with the cyclones is to directly couple one conduit to a corresponding cyclone. This arrangement would also have the advantage of minimizing the flow path between the annular collector of the riser and the cyclones where the final separation of catalyst and gas is performed. However, for reasons related to the complex hydrodynamics in the dilute phase region  232 , it has been found that mixtures of catalyst and gas that are taken from the reactor through a series of conduits may preferentially flow to one conduit. The resulting heavier loading of catalyst and gas can overload the cyclone to which it is directed. For this reason, the Figure shows the use of a chamber  238  that commonly collects the gas from all cyclone conduits  36  and redistributes the gas to the individual cyclones. Although providing chamber  238  and T-section  240  increases the residence time for the catalyst and gas mixture as it flows from the reactor vessel to the cyclone inlets, this minor increase in residence time will not have a substantial impact on the quality of the product recovered from the cyclones. The avoidance of mal-distribution may also be accomplished by the use of a catalyst and gas separation device other than cyclones. 
     A quench fluid contacts vapor products passing from withdrawal conduits  236  to cyclones  242 . Any lowering of the reactor vapor stream temperature will decrease product losses. Accordingly contacting the reactor vapors with the quench at any point downstream of the riser will produce some benefit. Contacting reactor vapors after substantial removal of the catalyst particles minimizes the volume of quench needed to achieve a desired degree of cooling and the amount of quench lost by adsorption on the catalyst. The quick separation arrangement of this invention provides a particularly advantageous arrangement for use of a quench. The ballistic separation of the riser effluent provides faster separation of the catalyst from the vapor than normally attained by the use of cyclones. The rapidly separated vapors from the ballistic separation section exit with only minor catalyst particle loading, typically on the order of 0.1-1.0 lb/ft 3 . Rapid separation and efficient separation minimizes thermal cracking as well as volumetric requirements of quench fluid. 
     The quench fluid can contact the product vapors at any point between the inlets for withdrawal vapor lines  236  and the cyclones  242 . Mixing of the quench fluid with the product vapors downstream of cyclones  242  can add from 0.5 to 5 seconds of high temperature exposure to the product vapors. Secondary cyclones, such as cyclones  242  typically have a high volume which exacerbates the problem of extended residence time. The most rapid quenching is obtained by contacting the quench stream immediately downstream of the ballistic separation. In the preferred form of this invention the quench enters single conduit  241 . Addition of quench to single conduit  241  has the advantage of providing a location external to the reactor vessel for the addition of quench as well as offering a relatively small cross-sectional area for immediate and complete mixing of the quench fluid with the vapors. 
     Catalyst that is initially separated from the gaseous components as it enters the reactor vessel, passes downward through the vessels as previously described. As this catalyst progresses through the vessel, it preferably contacts a series of baffles that improve the contact of the catalyst with a stripping gas that passes upwardly through the vessel. In the embodiment of the invention shown in the FIG. 2, the catalyst passes through a stripping section in the upper portion of the vessel referred to as a disengaging vessel and a separate stripping vessel located therebelow. The Figure shows the baffles  256  and  262  located on the exterior of the vessel walls and baffles  258  and  260  located down the length of the riser through the lower portion of the reactor vessel and the stripping vessel. These stripping baffles function in the usual manner to cascade catalyst from side to side as it passes through the vessel and increase the contact of the catalyst particles with the stripping steam as it passes upward in countercurrent contact with the catalyst. 
     The stripping vessel of FIG. 2 also provides hot stripping using catalyst from the regeneration zone to supply heat to the stripping section. A suitable lift system can be used to transport the catalyst upward from the regeneration zone into a stripping zone at a desired elevation. 
     With the cyclones removed from the reactor vessel, the diameter of the reactor vessel can be kept low enough such that the average residence time in the dilute phase of the reactor vessel is less than three seconds. Nevertheless, this embodiment of the invention also applies to an arrangement where the secondary separation device, such as cyclones  242 , are located within the reactor vessel and the only locations for quench contacting are inside the reactor vessel. In such an arrangement the a separate disengaging vessel is at least partially contained with the reactor vessel to minimize the volume into which the catalyst and hydrocarbons are initially discharged. In the embodiment shown in FIG. 2 the reactor vessel also provides the disengaging vessel. 
     In another alternate arrangement of this invention it is possible to use the vented riser in a manner to eliminate the disengaging vessel altogether. Such an arrangement withdraws catalyst and vapors from an extended riser through ports on the sides of the riser. The ports have a location below an open top of the riser and transfer the catalyst and hydrocarbon vapors to cyclones or other separation devices. The end of the riser extends upwardly by a distance sufficient to form a suspended layer of catalyst that seals the end of the riser. Under normal circumstances this type of riser arrangement operates in much the same manner as the riser and cyclone arrangement shown in FIG.  1  and does not permit catalyst or vapor to exit the top of the riser. However, the open end of the riser relieves pressure surges during upset conditions by venting vapors and catalyst into the open volume of the reactor vessel. Additional details of this arrangement are shown in pending application U.S. Ser. No. 790,924. 
     The unexpected advantages of the FCC arrangement of this invention are demonstrated by the following examples of FCC operations. These examples compare the operation of a conventional FCC operation with the operation of an FCC unit that operates in accordance with this invention. The data for both of these operations are presented in the following case studies which are calculated yield estimates based on simulations that have been developed from pilot plant data and operating data from commercial FCC units. 
     EXAMPLE 1 
     In a base case, a feed having a composition as set forth in the Table 1 was charged to a riser and contacted with a low rare earth catalyst having less than 1 wt. % rare earth exchange, a dealuminated zeolite content of about 30 wt. % in an active matrix component and a MAT activity of  68 . The catalyst was passed from the regenerator to the riser at a temperature of about 1321° F. The feed and catalyst mixture passed through the riser at an average temperature of 970° F. for an average time of three seconds and was discharged directly into a reactor vessel. Separated catalyst from the cyclone was discharged into a subadjacent stripping zone and contacted with a stripping steam at conditions that maintained an average stripping zone temperature of 970° F. Vapors removed from the catalyst in the stripping zone were vented into the reactor vessel and withdrawn through a first cyclone that operates in closed communication with the second cyclone to recover product vapors from the reactor vessel. Additional amounts of catalyst particles separated from the product vapors by the cyclones were discharged into the stripping zone. A vapor line carried all of the product vapors from the second stage cyclone to a main column fractionator. The cooled vapors had the composition set forth in Table 2. 
     EXAMPLE 2 
     In a first light olefin case, a feed again having the composition as set forth in the Table 1 was charged to a riser and contacted with a low rare earth catalyst having less than 1 wt. % rare earth exchange, a dealuminated zeolite content of about 30 wt. % in an active matrix component and a MAT activity of  68 . The catalyst was passed from the regenerator at a temperature of 1350° F. The feed and catalyst mixture passed through the riser for an average riser residence time of three seconds and was discharged from the riser outlet at an average temperature of 1025° F. directly into the first stage of a cyclone separator. Separated catalyst from the first stage cyclone dropped into a subadjacent stripping zone and into contact with a stripping steam at conditions that maintained an average stripping zone temperature of 1100° F. Vapors removed from the catalyst in the stripping zone were vented into a second stage of the cyclone separator that also received, in closed communication, vapors recovered from the first cyclone. Additional amounts of catalyst particles were separated from the product and stripping gases by the second cyclone stage and discharged into the stripping zone. All of the vapor from the second stage cyclone was discharged directly into a quench zone. The quench zone contacted the vapors from the second stage cyclone with cycle oil from the main column fractionator that cooled the product vapors to a temperature of 800° F. The cooled vapors had the composition set forth in Table 2. 
     EXAMPLE 3 
     In a second light olefin case, a feed again having the composition as set forth in the Table 1 was charged to a riser and contacted with a low rare earth catalyst having less than 1 wt. % rare earth exchange, a dealuminated zeolite content of about 40 wt. % in an active matrix component and a MAT activity of  72 . The catalyst was passed from the regenerator at a temperature of 1351° F. The feed and catalyst mixture passed through the riser for an average riser residence time of three seconds and was discharged from the riser outlet at an average temperature of 1025° F. directly into the first stage of a cyclone separator. Separated catalyst from the first stage cyclone dropped into a subadjacent stripping zone and into contact with a stripping steam at conditions that maintained an average stripping zone temperature of 1100° F. Vapors removed from the catalyst in the stripping zone were vented into a second stage of the cyclone separator that also received, in closed communication, vapors recovered from the first cyclone. Additional amounts of catalyst particles were separated from the product and stripping gases by the second cyclone stage and discharged into the stripping zone. All of the vapor from the second stage cyclone was discharged directly into a quench zone. The quench zone contacted the vapors from the second stage cyclone with cycle oil from the main column fractionator that cooled the product vapors to a temperature of 800° F. The cooled. vapors had the composition set forth in Table 2. 
     As compared to the base case, the data demonstrates that the high temperature operation, direct discharge of the riser effluent into the cyclone system, the hot stripping operation, and the immediate quenching of the reactor products after discharge from the cyclones provide significant yield advantages for the first light olefin case both in terms of conversion, olefin production and gasoline octane. The conversion, olefin and gasoline octane advantages more than offset the slightly higher coke and light gas production obtained by the process of this invention as compared to the prior art process. 
     Further improvements in conversion, olefin product and gasoline octane were obtained by the use of a slightly more active catalyst. The rapid quenching and quick quench of this invention permits the beneficial use of a more active catalyst. 
     
       
         
               
               
               
             
           
               
                   
                 TABLE 1 
               
               
                   
                   
               
             
             
               
                   
                 API 
                 23.41 
               
               
                   
                 UOP MOLECULAR K 
                 11.73 
               
               
                   
                 WT. 
                 361.5 
               
               
                   
                 NICKEL, PPM 
                 0.55 
               
               
                   
                 VANADIUM, PPM 
                 0.60 
               
               
                   
                 SULFUR, WT. % 
                 2.38 
               
               
                   
                 RAMMSBOTTOM CARBON, WT. % 
                 0.70 
               
               
                   
                 PERCENT BOILING AT 650 ° F. 
                 0.0 
               
               
                   
                   
               
             
          
         
       
     
     
       
         
               
               
               
               
             
               
               
               
               
             
           
               
                   
                 TABLE 2 
               
               
                   
                   
               
               
                   
                   
                 Example 2 
                 Example 3 
               
               
                   
                 Example 1 
                 Light 
                 Light 
               
               
                   
                 Base 
                 Olefin 
                 Olefin 
               
               
                   
                 Case 
                 Case #1 
                 Case #2 
               
               
                   
                   
               
             
             
               
                   
               
             
          
           
               
                 Conversion, LV % 
                 75.9 
                 80.4 
                 83.0 
               
               
                 YIELDS, LV % on FEED 
               
               
                 C 3  = 
                 7.8 
                 10.5 
                 12.5 
               
               
                 C 3   
                 2.8 
                 3.1 
                 3.5 
               
               
                 C 3  =/C 3   
                 0.74 
                 .77 
                 0.78 
               
               
                 C 4  = 
                 8.5 
                 12.2 
                 13.9 
               
               
                 C 4   
                 6.0 
                 7.1 
                 6.5 
               
               
                 C 4  =/C 4   
                 0.58 
                 0.63 
                 0.68 
               
               
                 C 5  = 
                 6.6 
                 7.1 
                 7.8 
               
               
                 C 5   
                 5.0 
                 4.3 
                 4.3 
               
               
                 C 5  =/C 5   
                 0.57 
                 .62 
                 0.64 
               
               
                 C 5   +  Gasoline 
                 58.1 
                 55.6 
                 54.9 
               
               
                 LCO + MCB 
                 24.5 
                 19.6 
                 17.0 
               
               
                 Coke, wt. % 
                 5.1 
                 6.02 
                 6.4 
               
               
                 C 2  minus, wt. % 
                 3.6 
                 4.43 
                 4.65 
               
               
                 C 5   +  Gasoline 
               
               
                 RON 
                 92.6 
                 94.0 
                 94.8 
               
               
                 MON 
                 80.0 
                 81.8 
                 82.1