FCC DUAL ELEVATION RISER FEED DISTRIBUTORS FOR GASOLINE AND LIGHT OLEFIN MODES OF OPERATION

A fluid catalytic cracking process includes feeding hydrocarbon into a riser in the presence of a catalyst, cracking the hydrocarbon in the riser in the presence of the catalyst to form a cracked stream, and separating the catalyst from the cracked stream. When in a gasoline mode, the hydrocarbon is fed through a first distributor into the riser, and when in a light olefin mode, the hydrocarbon is fed through a second distributor into the riser. The second distributor is positioned at a higher elevation than the first distributor.

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Description

This application is a continuation-in-part of application number Ser. No. 11/365,715, filed on Mar. 1, 2006.

BACKGROUND OF THE INVENTION

This invention relates generally to a process for catalytic cracking of hydrocarbons.

DESCRIPTION OF THE PRIOR ART

Fluid catalytic cracking (FCC) is a catalytic conversion process for cracking, heavy hydrocarbons into lighter hydrocarbons accomplished by contacting the heavy hydrocarbons in a fluidized reaction zone with a catalyst composed of finely divided particulate material. Most FCC units use zeolite-containing catalyst having high activity and selectivity. As the cracking reaction proceeds, substantial amounts of highly carbonaceous material referred to as coke are deposited on the catalyst, forming spent catalyst. High temperature regeneration burns coke from the spent catalyst. The regenerated catalyst may be cooled before being returned to the reaction zone. Spent catalyst is continually removed from the reaction zone and replaced by essentially coke-free catalyst from the regeneration zone.

The basic components of the FCC process include a riser (internal or external), a reactor vessel for disengaging spent catalyst from product vapors, a regenerator and a catalyst stripper. In the riser, feed distributor nozzles input the hydrocarbon feed which contacts the catalyst and is cracked into a product stream containing lighter hydrocarbons. Regenerated catalyst and the hydrocarbon feed are transported upwardly in the riser by the expansion of the lift gases that result from the vaporization of the hydrocarbons, and other fluidizing mediums, upon contact with the hot catalyst. Steam or an inert gas may be used to accelerate catalysis in a first section of the riser prior to or during introduction of the feed.

Riser residence time is one of the leading factors that determine how effectively the heavy hydrocarbon feed is converted to lighter, more valuable products. Increasing riser residence time increases the percentage of heavy hydrocarbon feed that is converted to lighter products. An average riser cracking zone has a catalyst to oil (feed) contact time of 1 to 5 seconds. Generally, a single elevation riser feed distributor is used for the process to yield a single product slate. The elevation of an FCC feed distributor may comprise one or a plurality of distributor nozzles which determine the point of initial contact of feed with the regenerated catalyst. The elevation of the feed distributor along the riser determines the residence time in the riser. Varying residence time determines the products yielded in the process.

A problem presented by the prior art is that changing operation of an FCC unit, such as from olefin production mode to gasoline production mode, by substantially changing riser residence time may require the unit be shut down for equipment changes to be made. Shutting down an FCC process stops the production of the valuable products and disrupts the desired steady state of operations. Changing equipment prolongs the down time. Therefore, changing modes of operation in this way may be undesirable for several reasons, including effects on productivity, efficiency and cost.

SUMMARY OF THE INVENTION

An FCC process which may include a dual elevation riser feed distributor and may vary the residence time requirement for multiple modes of operation. One aspect of the invention may be the ability to shift operation for the unit to different modes of operation without shutting down the unit. In accordance with the invention, an FCC process may utilize a dual elevation riser feed distributor rather than a single elevation feed distributor. The dual elevation riser feed distributor may enable the refiner to operate the unit in one mode of operation, such as a gasoline mode, with a first elevation feed distributor or in a second mode of operation, such as a light olefin mode, with a second, higher, elevation feed distributor. In a preferred embodiment, both sets of feed distributors may be generally identical except for their positions vertically with respect to each other within the riser, and they may have generally the same number, size, and arrangement of nozzles, or the equivalents thereof, in order to provide the refiner with additional flexibility and options for running the new and improved FCC process.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is an elevational diagram showing an FCC unit.

FIG. 2 is a cross section taken along segment A-A in FIG. 1.

FIG. 3 is a cross section taken along segment B-B in FIG. 1.

FIG. 4 is an elevational diagram showing a feed distributor nozzle.

DETAILED DESCRIPTION OF THE INVENTION

This invention relates generally to an improved FCC process. Specifically, this invention may relate to an improved elevation riser distributor and may be useful for FCC operators to more efficiently shift modes to yield different products as dictated by the market. The process and apparatus aspects of this invention can be used in the design of new FCC units or to modify the operation of existing FCC units.

As shown in FIG. 1, an FCC system 10 may be utilized in the FCC process, which may include feeding hydrocarbon into a riser 20 in the presence of a catalyst. Hydrocarbon may be cracked in the riser 20 in the presence of catalyst to form a cracked stream. A reactor vessel 30, with a separation chamber 32, separates spent catalyst particles from the cracked stream. A stripping zone 40 removes residual adsorbed hydrocarbon from the surface of the catalyst optionally as the catalyst travels over baffles 42. Spent catalyst from the stripping zone 40 is regenerated in a regenerator 46 having one or more stages of regeneration. Regenerated catalyst from the regenerator 46 re-enters the riser 20 to continue the process.

FCC feedstocks, suitable for processing by the method of this invention, include conventional FCC feeds and higher boiling or residual feeds. The most common of the conventional feeds is a vacuum gas oil which is typically a hydrocarbon material having a boiling range of from 343° to 552° C. (650° to 1025° F.) and is prepared by vacuum fractionation of atmospheric residue. Heavy or residual feeds, i.e., boiling above 499° C. (930° F.), are also suitable. The FCC process of the present invention is suited best for feed stocks that are heavier than naphtha range hydrocarbons boiling above about 177° C. (350° F.). Hydrocarbon feed may be modified to other feeds with appropriate modifications such as understood by those of ordinary skill in the art.

Looking then at FIG. 1, the riser 20 provides a conversion zone for cracking of the feed hydrocarbons. The vertical riser 20 may have a smaller diameter than a mixing chamber 90, so that catalyst accelerates as it passes out of the mixing chamber 90 into the riser 20. The riser typically operates with dilute phase conditions above the point of feed injection wherein the density is usually less than 320 kg/m3 (20 lb/ft3) and, more typically, less than 160 kg/m3 (10 lb/ft3). Feed is introduced into the riser 20 by either a first feed distributor 52 at a relatively lower elevation or a second feed distributor 54 at a relatively higher elevation between a riser inlet 28 and a riser outlet 24. Volumetric expansion resulting from the rapid vaporization of the feed as it enters the riser further decreases the density of the catalyst within the riser to typically less than 160 kg/m3 (10 lb/ft3). Before contacting the catalyst, the feed will ordinarily have a temperature in a range of from 150° to 370° C. (302° to 698° F.). Additional amounts of feed may be added downstream of the initial feed point.

The blended catalyst and reacted feed vapors are then discharged from the top of the riser 20 through the riser outlet 24 and separated into a cracked product vapor stream and a collection of catalyst particles covered with substantial quantities of coke and generally referred to as “coked catalyst.” This invention can use any arrangement of separators to remove coked catalyst from the product stream quickly. In particular, a swirl arm arrangement 29, provided at the end of the riser 20 can further enhance initial catalyst and cracked hydrocarbon separation by imparting a tangential velocity to the exiting catalyst and cracked product vapor stream mixture. The swirl arm arrangement 29 is located in an upper portion of the separation chamber 32, and the stripping zone 40 is situated in the lower portion of the separation chamber 32. Catalyst separated by the swirl arm arrangement 29 drops down into the stripping zone 40. The cracked product vapor stream comprising cracked hydrocarbons including gasoline and light olefins and some catalyst exit the separation chamber 32 via a gas conduit 36 in communication with cyclones 33. The cyclones 33 remove remaining catalyst particles from the product vapor stream to reduce particle concentrations to very low levels. The product vapor stream then exits the top of the reactor vessel 30 through a product outlet 35. Catalyst separated by the cyclones 33 return to the reactor vessel 30 through diplegs into a dense bed 39 where it will enter the stripping zone 40 through openings 37. The stripping zone 40 removes adsorbed hydrocarbons from the surface of the catalyst by counter-current contact with steam over the optional baffles 42. Steam enters the stripping zone 40 through a line 41.

A first portion of the coked catalyst is returned to the riser 20 without first undergoing regeneration while a second portion of the coked catalyst is regenerated in the regenerator 46 before it is delivered to the riser 20. The first and second portions of the catalyst may be blended in the mixing chamber 90 before introduction to the riser 20. The recycled catalyst portion may be withdrawn from the stripping zone 40 for transfer to the mixing chamber 90. A recycle conduit 22 transfers the first portion of the coked catalyst stripped of hydrocarbon vapors exiting the stripping zone 40 back to the mixing chamber 90 as the recycled catalyst portion at a rate regulated by a control valve. The second portion of the coked, stripped catalyst is transported to the regenerator 46 through a coked catalyst conduit 23 at a rate regulated by a control valve for the removal of coke.

On the regeneration side of the process, coked catalyst transferred to the regenerator 46 via the conduit 23 undergoes the typical combustion of coke from the surface of the catalyst particles by contact with an oxygen-containing gas. The oxygen-containing gas enters the bottom of the regenerator 46 via a distributor 48 and passes through a dense fluid zing bed of catalyst. Flue gas consisting primarily of N2, H2O, O2, CO2 and perhaps containing CO passes upwardly from the dense bed into a dilute phase of the regenerator 46. A separator, such as cyclones 49 or other means, remove entrained catalyst particles from the rising flue gas before the flue gas exits the vessel through an outlet 50. Combustion of coke from the catalyst particles raises the temperatures of the catalyst which is withdrawn by a regenerator standpipe 18.

The regenerator standpipe 18 passes regenerated catalyst from the regenerator 46 into the mixing chamber 90 at a rate regulated by a control valve where it is blended with recycled catalyst from the stripping zone 40 via the recycle conduit 22. Fluidizing gas such as steam passed into the mixing chamber 90 by a distributor 26 contacts the catalyst and maintains the catalyst in a fluidized state to blend the recycled and regenerated catalyst. The regenerated catalyst which is relatively hot is cooled by the unregenerated, coked catalyst which is relatively cool to reduce the temperature of the regenerated catalyst by 28° C. to 83° C. (50° to 150° F.) depending upon the regenerator temperature and the coked catalyst recycle rate.

A dual elevation riser feed distributor may provide more efficient operation of an FCC process, as shown in FIG. 1. The time duration for the stream of hydrocarbon to pass through the riser 20, also known as residence time, affects the selectivity of the conversion of hydrocarbon feed to lighter olefins. We have found that selectivity to lighter olefins is increased when riser residence time is decreased relative to the residence time typical for FCC operation which has greater selectivity to gasoline. The location of the feed distributor along the riser 20 affects the residence time in the riser. Varying residence time may affect the products yielded in the process.

In one embodiment of the invention, when an FCC process is in a first mode, such as a gasoline operation mode, hydrocarbon feed is fed through a first feed distributor 52. When the FCC process is in a second mode, such as a light olefin mode, hydrocarbon feed is fed through the second feed distributor 54 which is higher in elevation than first feed distributor 52. The first feed distributor 52 is located along the riser 20 in a position to input feed upstream of the second feed distributor 54. Hydrocarbon fed through the first feed distributor 52 thereby has a longer residence time in the riser 20 than hydrocarbon fed through the second feed distributor 54. The difference in residence time for hydrocarbon fed through the first feed distributor 52 from hydrocarbon fed through the second feed distributor 54 results in an FCC process that may yield different products. For example, an FCC process designed to maximize gasoline selectivity has a longer residence time than an FCC process designed to maximize light olefin yield.

In an embodiment of the invention, hydrocarbon feed may be fed through the first feed distributor 52 at suitable position along the riser 20 to produce a substantial yield of gasoline in preference to other products. Preferably, residence time for hydrocarbon fed through the first feed distributor 52 may have a residence time between about 2 seconds and about 5 seconds to maximize gasoline yield. Residence time is the time it takes for the hydrocarbon to travel from the distributor 52, 54 to the riser outlet 24. Hydrocarbon may be fed through the second feed distributor 54 at suitable position along the riser 20 to produce a substantial yield of light olefins with greater selectivity to light olefins than in the gasoline mode in which feed is distributed through the first feed distributor 52. Preferably, residence time for hydrocarbon fed through the second feed distributor 54 may be between about 0.5 seconds and about 2 seconds to maximize the yield of light olefins. More preferably, residence time for hydrocarbon fed through the second feed distributor 54 may be about 1 to about 2 seconds.

As shown in FIGS. 2 and 3, each feed distributor 52, 54 may comprise one or more individual feed distributor nozzles 60. Preferably, a plurality of the feed distributor nozzles 60 may be utilized for each feed distributor 52, 54. In a preferred embodiment of the invention, two, three, four or more feed distributor nozzles 60 may be arranged generally uniformly around the riser 20. In a still more preferred embodiment, as shown in FIG. 2, four feed distributor nozzles 60 may be arranged radially around the riser 20. Each feed distributor 52, 54 may have different numbers of the nozzles 60.

When in gasoline operation mode, the process produces a cracked stream having gasoline vol-% between about 50 and about 70. Preferably, producing a cracked stream having gasoline vol-% between about 55 and about 65. When in light olefin operation node, the process may produce a cracked stream having light olefin vol-% between about 20 and about 50. Preferably, producing a cracked stream having light olefin vol-% between about 25 and about 40.

Regenerated catalyst from the regenerator standpipe 18 will usually have a temperature in a range from 677° to 760° C. (1250° to 1400° F.) and, more typically, in a range of from 699° to 760° C. (1290° to 1400° F.). In an embodiment, stripped catalyst from the stripping zone 40 may be recycled to the riser 20 without undergoing regeneration. The mixing chamber 90 may be used to blend spent and regenerated catalyst for sufficient time to achieve substantially thermal equilibrium. The temperature of the recycled catalyst portion will usually be in a range of from 510° to 621° C. (950° to 1150° F.). The relative proportions of the recycled and regenerated catalyst will determine the temperature of the blended catalyst mixture that enters the riser. The blended catalyst mixture will usually range from about 593° to 704° C. (1100° to 1300° F.).

In an embodiment, the temperature in the reactor vessel 30 in light olefin operational mode may be between about 490° and about 630° C. at the riser outlet 24. More preferably, the reactor temperature may be between about 550° and about 590° C. The reactor temperature is typically less in gasoline mode than in light olefin mode. The temperature in the reactor vessel 30 in gasoline operation mode may be between about 460° and about 600° C. More preferably, reactor temperature may be between about 500° and about 550° C.

The recycle of spent catalyst to the riser bypassing regeneration enables the FCC unit to be run at higher ratios of catalyst to feed without impacting the heat balance of the unit. The light olefin mode can be run at higher catalyst to feed ratios than the gasoline mode to produce more light olefins. The catalyst to feed ratio in the riser 20 when in light olefin mode may be 15 or greater. Whereas, the catalyst to feed ratio in the riser may be less than 15 while operating in gasoline mode.

Reactor pressure in kPa may be between about 93 and about 113. Lower hydrocarbon partial pressure in the riser 20 operates to favor the production of light olefins. Low hydrocarbon partial pressure can be achieved by using steam or other inert gas as a diluent. In light olefin mode, 5 to 25 wt-% steam relative to the hydrocarbon feed may be added to reduce hydrocarbon partial pressure. Typically in gasoline mode, FCC units are operated with 0.5 to 5 wt-% steam to disperse the feed and purge stagnant zones. Hence, light olefin mode is operated with higher steam rates than gasoline mode. Only the fluidizing gas distributor 26 is shown in the drawings. However, other steam distributors may be provided along the riser and elsewhere in the FCC unit.

The zeolitic molecular sieves used in typical FCC gasoline mode operation have a large average pore size. Typically, molecular sieves with a large pore size have pores with openings of greater than 0.7 nm in effective diameter defined by greater than 10 and typically 12 membered rings. Pore Size Indices of large pores are above about 31. Suitable large pore molecular sieves include synthetic zeolites such as X-type and Y-type zeolites, mordenite and faujasite. Y zeolites with low rare earth content are preferred. Low rare earth content denotes less than or equal to about 1.0 wt-% rare earth oxide on the zeolitic portion of the catalyst.

Catalyst additive may be added to the catalyst composition when operating in the light olefin mode. Catalyst additive includes a medium or smaller pore zeolite catalyst exemplified by ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38. ZSM-48, and other similar materials. U.S. Pat. No. 3,702,886 describes ZSM-5. Other suitable medium or smaller pore zeolites include ferrierite, erionite, and ST-5, developed by Petroleos de Venezuela, S.A. The medium or smaller pore zeolite may be dispersed on a matrix comprising a binder material such as silica or alumina and an inert filer material such as kaolin. The catalyst additive may also comprise some other active material such as beta zeolite. These catalyst additives have a crystalline zeolite content of 10 to 40 wt-% or more and a matrix material content of 60 to 90 wt-% or less. Catalysts containing 25 wt-% crystalline zeolite are typical. Medium and smaller pore zeolites are characterized by having an effective pore opening diameter of less than or equal to 0.7 nm, rings of 10 or fewer members and a Pore Size Index of less than 31.

When in light olefin operation mode, the cracking step may be conducted in the presence of an effective amount of medium or smaller pore zeolitic catalyst such as ZSM-5 shape-selective catalyst. Preferably, in light olefin mode, the cracking step may be conducted in the presence of 2 to 20 wt-% of medium or smaller pore zeolite. More preferably, the cracking step may be conducted in the presence of an additional 3 to 10 wt-% medium or smaller pore zeolite. Fresh catalyst is added periodically to the FCC unit to make up for catalyst loss in the product outlet 35 and flue gas from the regenerator outlet 50. When switching to gasoline mode from light olefin mode, no small or medium pore zeolite should be added to allow the inventory to purge itself of this additive. Eventually, the catalyst inventory in the FCC unit will have less than 0.5 wt-% additive. When switching to light olefin mode from gasoline mode a large amount of medium or smaller pore zeolite may be added all at once or in increments to bring the additive level in the catalyst inventory up to that desired for light olefin mode.

In the process of the present invention, as shown in FIG. 4, hydrocarbon and steam may be introduced through the same feed distributor nozzle. Other types of feed distributor nozzles may be suitable. A distributor barrel 62 for each distributor nozzle 60 receives steam from a steam inlet pipe 68. A barrel body flange 64 secures the distributor barrel 62 to a riser nozzle 70 in the reactor riser 20 by bolts and may be oriented such that the bolt holes straddle a radial centerline of the riser 20. An oil inlet pipe 76 delivers hydrocarbon feed to an internal oil pipe 78. An oil inlet barrel flange 66 secures the oil inlet pipe 76 to the distributor barrel 62 by bolts. Vanes 83 in the internal oil pipe 82 cause the oil to swirl in the oil pipe before exiting. The internal oil pipe 78 distributes swirling oil to the distributor barrel 62 where it mixes with steam and is injected from orifices 80 in the face of a distributor tip 74 extending into the riser 20. Distributor metallurgy, except for the internal oil pipe 78 and the distributor tip 74 may be a high chromium steel alloy, preferably 9 Cr-1 Mo. Distributor metallurgy for the internal oil pipe 78 and the distributor tip 74 may be a cobalt-based alloy, preferably Cobalt Alloy 6 (AMS 5387).

In one embodiment of this invention, a hydrocarbon stream is introduced through a third feed distributor 56, as shown in FIG. 1. Preferably, the third, alternate, feed distributor 56 may introduce feed into the mixing chamber 90. The FCC unit may be run in a recracking mode in which feed is introduced through the third feed distributor 56 lower than and upstream of the feed distributors 52, 54 solely without or in conjunction with feed being distributed through one or both of the feed distributors 52, 54. In one embodiment, a C8-hydrocarbon feed stream comprising molecules predominantly having less than eight carbons may be introduced through the third, alternate feed distributor 56. In another embodiment, a C6+ hydrocarbon feed stream is introduced through the third, alternate feed distributor 56. The feed may be obtained from an outside source but is preferably derived from products recovered from the product outlet 35 and further processed to yield a cut of intermediate product that can be further upgraded. Upwards of about 30 wt-% of C8-hydrocarbons may be cracked in the mixing chamber 90 to more valuable products. Steam introduced in the mixing chamber 90 to fluidize the chamber and facilitate mixing of spent and regenerated catalyst may be utilized to adjust superficial velocity and catalyst density to achieve a fast fluidized flow regime. Weight hourly space velocity of the hydrocarbons and residence time may also be influenced. Additional steam also may be introduced in the riser 20 below the feed distributor nozzles 60 to adjust the velocity in the riser 20 prior to the introduction of the conventional FCC feed.

In one embodiment of this invention, the FCC process may include injecting a fluid such as light cycle oil (LCO) through an injector 34 to quench cracked product vapors in the gas conduit 36 from the riser outlet 24. The quench from the injector 34 may reduce the outlet temperature. Reducing the temperature helps to minimize post riser thermal cracking and to maintain the yields and selectivity of desired products from the riser 20. Furthermore, the refiner thereby possesses greater flexibility in controlling the reactor outlet temperature. The quench process of this embodiment of the invention may reduce temperatures by about 28° C. to about 56° C. (50° F. to about 100° F.). The quench fluid may be derived from products exiting product outlet 35.

EXAMPLE

In an example of one embodiment of this invention, an FCC process of the present invention was simulated to run in gasoline mode under the conditions given for the present invention which produced a gasoline selectivity of 60 vol-% and a propylene selectivity of 8 vol-%. When the present invention was simulated to run in light olefin mode under the conditions given for the present invention it produced a gasoline selectivity of 33 vol-% and a propylene selectivity of 32 vol-%.

While the foregoing written description of the invention enables one of ordinary skill to make and use what is considered presently to be the best mode thereof, those of ordinary skill will understand and appreciate the existence of variations, combinations, and equivalents of the specific exemplary embodiments thereof. The invention is therefore to be limited not by the exemplary embodiments herein, but by all embodiments within the scope and spirit of the appended claims.

Claims

1. A fluid catalytic cracking process, comprising: wherein, when in a gasoline mode, said hydrocarbon is fed through a first distributor into said riser, and when in a light olefin mode, said hydrocarbon is fed through a second distributor into said riser, wherein said second distributor is positioned at a higher elevation than said first distributor.

feeding hydrocarbon into a riser in the presence of a catalyst;
cracking said hydrocarbon in said riser in the presence of said catalyst to form a cracked stream, and
separating said catalyst from said cracked stream,

2. The process of claim 1, wherein each one of said distributors comprises a plurality of nozzles having substantially the same elevation and being spaced radially around said riser.

3. The process of claim 1, wherein said separation of catalyst from said cracked stream is performed in a swirl arm arrangement.

4. The process of claim 1, wherein the light olefin mode has residence time of between about 0.5 second and about 2 seconds in said riser.

5. The process of claim 1, wherein propylene-preferential mode has residence time of about 1 to about 2 seconds in said riser.

6. The process of claim 1, wherein gasoline-making mode has residence time of between about 2 seconds and about 5 seconds.

7. The process of claim 1, wherein said FCC process is changeable from one mode to another without shutting the process down.

8. The process of claim 1, wherein said hydrocarbon and steam are fed through same distributor nozzle.

9. The process of claim 1, wherein when in said gasoline mode, said process produces a cracked stream having gasoline vol-% between about 50 and about 70.

10. The process of claim 1, wherein, when in said light olefin mode, said process produces a cracked stream having vol-% of propylene between about 20 and about 40.

11. The process of claim 1, wherein said temperature at the outlet of said riser is between about 490° and about 630° C.

12. The process of claim 1, wherein said catalyst is about 2 to about 20 wt-% medium or smaller pore zeolite while operating in said light olefin mode and less than 0.5 wt-% medium or smaller pore zeolite while operating in said gasoline mode.

13. The process of claim 1, wherein said temperature in said reactor in said light olefin mode is between about 550° and about 590° C.

14. The process of claim 1, wherein said temperature in said reactor in said gasoline-preferential mode is between about 500° and about 550° C.

15. The process of claim 1, wherein more steam is added to the riser when operating in the light olefin mode than when operating in the gasoline mode.

16. The process of claim 1, wherein a greater catalyst to feed ratio is used when operating in the light olefin mode than when operating in the gasoline mode.

17. The process of claim 1 further comprising the step of quenching said cracked stream with a liquid after exiting said riser.

18. A fluid catalytic cracking process, comprising:

feeding hydrocarbon into a first region having a first predetermined elevation in a riser in the presence of a catalyst, wherein said thirst region and said first predetermined elevation are selected to provide a greater yield of light olefins in a light olefin mode;
cracking said hydrocarbon in said riser in the presence of said catalyst to form a cracked stream;
separating said catalyst from said cracked stream; and
changing from said light olefin mode to a gasoline mode by changing the location of the feed of said hydrocarbon to a second region having a second predetermined elevation lower than that of said first predetermined elevation to produce a smaller yield of light olefins than when operated in the light olefin mode.

19. A fluid catalytic cracking process, comprising: wherein, when in a recracking mode, said hydrocarbon is fed through a first distributor into a mixing chamber below said riser.

feeding hydrocarbon into a riser in the presence of a catalyst;
cracking said hydrocarbon in said riser in the presence of said catalyst to form a cracked stream; and
separating said catalyst from said cracked stream,

20. The process of claim 19 further comprising the step of introducing C8-feed into said mixing chamber below said riser.

Patent History

Publication number: 20070205139
Type: Application
Filed: Dec 29, 2006
Publication Date: Sep 6, 2007
Inventors: SATHIT KULPRATHIPANJA (Des Plaines, IL), Daniel N. Myers (Des Plaines, IL), Paolo Palmas (Des Plaines, IL), Mark W. Schnaith (Des Plaines, IL), Ismail B. Cetinkaya (Des Plaines, IL), Peter J. Van Opdorp (Des Plaines, IL), Charles L. Hemler (Des Plaines, IL)
Application Number: 11/617,894

Classifications

Current U.S. Class: Catalytic (208/113)
International Classification: C10G 11/00 (20060101);