MXPA00010667A - Two stage fluid catalytic cracking process for selectively producing c2 - Google Patents

Abstract

A two stage process for selectively producing C2 to C4 olefins from a gas oil or resid. The gas oil or resid is reacted in a first stage comprised of a fluid catalytic cracking unit wherein it is converted in the presence of conventional large pore zeolitic catalyst to reaction products, including a naphtha boiling range stream. The naphtha boiling range stream is introduced into a second stage comprised of a process unit containing a reaction zone, a stripping zone, a catalyst regeneration zone, and a fractionation zone. The naphtha feedstream is contacted in the reaction zone with a catalyst containing from about 10 to 50 wt.%of a crystalline zeolite having an average pore diameter less than about 0.7 nanometers at reaction conditions which include temperatures ranging from about 500 to 650°C anda hydrocarbon partial pressure from about 10 to 40 psia. Vapor products are collected overhead and the catalyst particles are passed through the stripping zone on the way to the catalyst regeneration zone. Volatiles are stripped with steam in the stripping zone and the catalyst particles are sent to the catalyst regeneration zone where coke is burned from the catalyst, which is then recycled to the reaction zone.

Description

CATALYTIC FRACTIONING PROCEDURE TWO STAGES FLUID TO SELECTIVELY PRODUCE OLEFINS C2 TO C4 FIELD OF THE INVENTION The present invention relates to a process in do stages to selectively produce C2 to C4 olefins from diesel or waste. The gas oil or residue reacts in a first stage comprising a fluid catalytic fractionation unit where it is converted into the presence of conventional large pore zeolitic catalyst and reaction products, including a boiling range stream of naphtha. The boiling range stream of naphtha is introduced into a second stage consisting of a processing unit containing a reaction zone, a depletion zone, a catalyst regeneration zone, and a fractionation zone. The naphtha feed stream comes into contact in the reaction zone with a catalyst containing from about 10 to 50% by weight of a crystalline zeolite having an average pore diameter of less than about 0.7 nanometers under the reaction conditions including temperatures that fall within a range of approximately 500 to 650 ° C and a partial hydrocarbon pressure of approximately 7,000 to 28,000 kg / m2, absolute pressure (approximately 10 to 40 ). The products in the vapor state are collected in the upper part and the catalyst particles pass through the exhaustion zone towards the catalyst regeneration zone. The volatile parts are exhausted with steam in the exhaustion zone and the catalyst particles they are sent to the catalyst regeneration zone where the catalyst is burned and coked, which is then recycled to the reaction zone. BACKGROUND OF THE INVENTION The need for fuel with low emissions has created an increased demand for light olefins for its use in alkylation, oligomerization, MTBE and ETBE synthesis processes. In addition, a low-cost supply of light olefins, particularly propylene, is still demanded to serve as a feed for polyolefin, especially for the production of polypropylene. Fixed-bed procedures for light paraffin dehydrogenation have recently attracted renewed interest in increasing olefin production. However, these types of procedures typically require relatively large capital investments and present high operating costs. It is therefore advantageous to increase the yield of olefin by employing procedures which require relatively small capital investments. It would be particularly advantageous to increase the yield of olefins in catalytic fractionation processes.

Catalytic fractionation is a established and widely used method in the oil refining industry to convert relatively high boiling point petroleum oils into lower valued, more valuable boiling products, including gasoline as well as intermediate distillate products such as, For example, kerosene, fuel for airplanes and oil for heating. The dominant catalytic fractionation process now in use is the fluid catalytic fractionation (FCC) process in which a preheated feed comes in contact with a thermofraction catalyst d in the form of a fine powder, typically a powder having a particle size of approximately 10 to 300 microns, usually approximately 60 to 7 microns, so that the desired reactions d fractionation occurs. During fractionation, coke and hydrocarbon material are deposited in the catalyst particles. This results in a loss of catalyst activity and a loss of its selectivity. The particles of catalyst with coke, and associated hydrocarbon material, are subjected to a depletion process, usually with steam, to remove as much technical and economically feasible amount of hydrocarbon material as possible. Depleted particles that contain coke that can not be depleted, they are removed from the depletion column and sent to a regenerator where the catalyst particles with coke are regenerated by their contact with the air, or a mixture of air and oxygen, at an elevated temperature. This results in the combustion of the coke which is a strongly exothermic reaction which, in addition to removing the coke, serves to heat the catalyst to the temperatures appropriate for the endothermic fractionation reaction. The process is carried out in an integrated unit comprising the fractionation reactor, the depletion column, the regenerator, and the appropriate auxiliary equipment. The catalyst flows continuously from the reactor or reaction zone, to the depletion column and then to the regenerator and back to the reactor. The flow rate is typically adjusted in relation to the oil feed rate to maintain a thermally balanced operation in which the heat produced in the regenerator is sufficient to maintain the fractionation reaction with the circulating regenerated catalyst used as the transfer medium of heat. Typical fluid catalytic fractionation procedures are described in the monograph "Fluid Catalytic Cracking with Zeolite Catalysts", (Fluid Catalytic Fractionation with Zeolite Catalysts), Venuto, P.B. and Aviv, E.T., Marcel Dekker Inc. N.Y. 1979, which is incorporated herein by reference. As described in this monograph, the conventionally used catalysts are based on zeolites, especially the large pore synthetic faujasites, X and Y zeolites. Typical feeds to a catalytic fractionator can generally be characterized as a relatively high boiling point oil or residue. , either only well mixed with other fractions, usually also with a relatively high boiling point. The most common feeds are diesel, that is, high-boiling non-residual oil with an initial boiling point generally above about 230 ° C, more commonly above about 350 ° C, with end points up to about 620 ° C. . Typical diesel oils include direct (atmospheric) gas oil, vacuum gas oil, and coke gas oils. While such conventional fluid catalytic fractionation processes are suitable for the production of conventional transportation fuels, such fuels generally can not meet the most demanding requirements of low emission fuels and chemical feedstock production. To increase the volume of low emission fuels, and it is desirable to increase the amounts of light olefins, such as, for example, propylene, isobutylenes and normal butylenes, as well as isoamylene. Propylene, isobutylene, isoamylene can be reacted with methanol to form methyl propyl ethers, methyl tertiary butyl ether (MTBE), methyl tertiary amyl ether (TAME). They are mixed components of high octane that can be added to gasoline to meet the oxygen requirements established by law. In addition to increasing the volume and octane number of gasoline, they also reduce emissions. It is particularly desirable to increase the yield of ethylene and propylene which are valuable as chemical raw materials. The conventional fluid catalytic fractionation n produces sufficiently large quantities of these light olefins, particularly ethylene. Accordingly, there is a need in the art to have methods for producing higher amounts of ethylene and propylene for raw material for chemical agents, as well as other light olefins for low emission fuels such as, for example, gasoline and distillates. . U.S. Patent No. 4,830,728 discloses a fluid catalytic fractionation unit (FCC) that operates to optimize olefin production. The FCC unit has two separate risers in which a different feed stream is introduced. The operation of the riser is designed in such a way that a suitable catalyst acts to convert a heavy gas oil into a riser pipe and another suitable catalyst acts to split a lighter olefin / naphtha feed into the other riser pipe. Conditions inside the riser tube of heavy gas oil can be modified to optimize the production of either gasoline or olefin. The primary means to optimize the production of the desired product is through the use of a specific catalyst. Likewise, US Patent No. 5,026,936 to Arco presents a process for the preparation of propylene a * from C4 supplies or higher by a combination of fractionation and metathesis, where the higher hydrocarbon is fractionated to form ethylene and propylene and at least part of the ethylene is metatized to propylene. See, also, US Patent Nos. 5,026,935 and 5,043,522. U.S. Patent No. 5,069,776 discloses a process for the conversion of a hydrocarbon feed by contacting the feed with a mobile bed of a zeolitic catalyst comprising a zeolite with a pore diameter of 0.3 to 0.7 nm, at a temperature above about 500 ° C, and with residence time less than about 10 seconds. Olefins are produced with relatively small formation of saturated gaseous hydrocarbons. Likewise, US Patent No. 3,928,172 to Mobil presents a process for converting hydrocarbon feeds in which the olefins are produced by reacting said feed in the presence of a ZSM-5 catalyst. A problem inherent in the production of olefin products using FCC units is that the process depends on a specific balance of catalyst to optimize production. Furthermore, even if a specific catalyst equilibrium can be maintained to optimize the overall olefin production, the selectivity of the olefin is generally low due to undesirable side reactions such as, for example, extensive fractionation, isomerization, aromatization, and transfer reactions. hydrogen. Accordingly, it is desirable to optimize olefin production in a process that allows a high degree of control over the selectivity of C2, C3 and C4 olefins. SUMMARY OF THE INVENTION In accordance with the present invention, there is provided a two step process for selectively producing C2 to C olefins from a gas oil or residue. The gas oil or residue reacts in a first stage comprising a fluid catalytic fractionation unit where it is converted in the presence of conventional large pore zeolitic catalyst into reaction products, including a stream of boiling range of naphtha. The boiling range stream of naphtha is introduced in a second stage consisting of a process unit comprising a reaction zone, a depletion zone, a catalyst regeneration zone, and a fractionation zone. The naphtha feed stream comes into contact in the reaction zone with a catalyst containing from about 10 to 50% by weight of a crystalline zeolite having an average pore diameter of less than about 0.7 nanometers under reaction conditions including temperatures within a range of approximately 500 to 650 ° C and a partial hydrocarbon pressure of approximately 7,000 to 28,000 Kg / m2, absolute pressure (approximately 10 to 40 psia). Products are collected in the vapor state above and the catalyst particles are passed through the depletion zone towards the catalyst regeneration zone. The volatile parts are depleted with steam in the depletion zone and the catalyst particles are sent to the catalyst regeneration zone where the coke of the catalyst is burned, which is then recycled to the reaction zone. In another preferred embodiment of the present invention, the second stage catalyst is a ZSM-5 type catalyst. In another preferred embodiment of the present invention, the second stage feed contains from about 10 to 30% by weight of paraffins, and from about 20 to 70% by weight of olefins.

In another preferred embodiment of the present invention, the second stage reaction zone operates at a temperature from about 525 ° C to about 600 ° C. DETAILED DESCRIPTION OF THE INVENTION The feed stream of the first stage of the present invention is preferably a hydrocarbon fraction having an initial ASTM boiling point of about 315 ° C (600 ° F). Such hydrocarbon fractions include gas oil (including vacuum gas oil), thermal oils, residual oils, crude primary primary distillation, oils from tar sands, shale oils, synthetic fuels, fractions of heavy hydrocarbons derived from the destructive hydrogenation of coal, tar, fish, asphalts, hydrotreated feeds derived from any of the previous The feed reacts (is converted) in a first stage, preferably in a fluid catalytic fractionation reactor vessel where it comes into contact with a catalytic fractionation catalyst continuously recycled. The feed may be mixed with a stream or an inert gas under conditions such that it forms a highly atomized stream of a hydrocarbon-catalyst vapor suspension that is subjected to reaction. Preferably, this reaction suspension flows through a riser tube in the reactor vessel. The container of the reaction zone preferably operates at a temperature of about 425 to 650 ° C (about 800 to 1200 ° F) and at a pressure of about 0-70,000 kg / m2 above atmospheric pressure (from about 0 to 100). psig). The catalytic fractionation reaction is essentially turned off by separating the catalyst from the vapor. The separated vapor comprises the fractionated hydrocarbon product, and the separated catalyst contains a carbonaceous material (i.e., coke) as a result of the catalytic fractionation reaction. The coked catalyst is preferably recycled to come into contact with additional hydrocarbon feed after the removal of the coke material. Preferably, the coke is removed from the catalyst in a regenerator vessel by combustion of the catalyst coke. Preferably, the coke is burned at a temperature of about 480-760 ° C (900-1400 ° F) at a pressure of about 0-70,000 kg / m2 above atmospheric pressure (0-100 psig). After the combustion step, the regenerated catalyst is recycled to the riser for contact with additional hydrocarbon feed. The catalyst that is employed in the first step of this invention can be any catalyst typically 12 used to catalytically fractionate hydrocarbon feeds. It is preferred that the catalytic fractionation catalyst comprises an oxide component of crystalline tetrahedral structure. This component is used to catalyze the decomposition of the primary products from the catalytic fractionation reaction in clean products such as petrol naphtha and olefins for chemical feeds. Preferably, the crystalline tetrahydric structure oxide component is selected from the group consisting of zeolites, tectosilicates, tetrahedral aluminophosphate (ALPOs) as well as tetrahedral silicoaluminophosphate (SAPOs). More preferably, the oxide component of the crystal structure is a zeolite. Zeolites that can be employed in the first stage catalysts of the present invention include both natural and synthetic zeolites with average pore diameters greater than about 0.7 nm. These zeolites include gmelinite, chabazite, daquiardite, clinoptilotite, faujasite, heulandite, analyst, levinite, heroinite, sodalite, cancrinite, nepheline, lazurite, scolecite, natrolite, ofendite, mesolite, mordenite, brusterite, and ferrierite. Among the synthetic zeolites are the zeolites X, Y, A, L, ZK-4, ZK-5, B, E, F, H, J, M, Q, T, W, Z, alpha, bet and omega as well as zeolites USY. USY zeolites are preferred.

In general, aluminosilicate zeolites are effectively employed in this invention. However, the aluminum as well as the silicon component may be substituted by other structure components. For example, the aluminum portion may be replaced by boron, gallium, titanium or trivalent metal compositions heavier than aluminum. Germanium may be used instead of the silicon portion. The catalytic fractionation catalyst employed in the first step of this invention may further comprise an active porous inorganic oxide catalyst structure component and an inert catalyst structure component. Preferably, each component of the catalyst is agglomerated by the use of an inorganic oxide matrix component. The active porous inorganic oxide catalyst structure component catalyzes the formation of primary products by fractionating molecules that are too large to fit within the oxide component of tetrahedral structure. The active porous inorganic oxide catalyst structure component of this invention is preferably a porous inorganic oxide that fractionates a relatively large amount of hydrocarbons into hydrocarbons of lower molecular weights as compared to an acceptable thermal target. A silica with a low surface area (eg, quartz) is an acceptable type of thermal target. The extent of the fractionation can be measured in any of several ASTM tests such as MAT (microactivity test, ASTM # D3907-8). Compounds such as those disclosed in Greensfeider, B.S., et al. , Industrial and Engineering Chemistry, pages 2573-83, November 1949, are desirable. Alumina, silica-alumina as well as silica-alumina-zirconia compounds are preferred. The inert catalyst structure component densifies, reinforces and acts as a protective thermal sink. The inert catalyst structure component employed in this invention preferably has a fractionating activity that is not significantly greater than the acceptable thermal target. Kaolin and other clays as well as α-alumina, titania, zirconia, quartz and silica are examples of preferred inert components. The inorganic oxide matrix component binds the catalyst components in such a way that the catalyst product is sufficiently hard to survive collisions between particles and against the walls of the reactor. The inorganic oxide matrix can be made from a dried inorganic oxide sol or gel to "glue" the catalyst components together. Preferably, the inorganic oxide matrix can be formed of silicon and aluminum oxides. It is also preferred that separate alumina phases are incorporated into the inorganic oxide matrix. Aluminum oxyhydroxide-g-alumina species, diasporous boehmite, as well as transition aluminas such as alumina, b-alumina, g-alumina, d-alumina , e-alumina, k alumina, and r-alumina can be used. Preferably, the alumina species is an aluminum trihydroxide such as, for example, gibbsite, bayerite, nordstrandite, or doyelite. The matrix material may also contain phosphorus or aluminum phosphate. A fraction of the boiling range of naphtha from the product stream from the fluid catalytic fractionation unit is used as the feed stream to a second reaction step to selectively produce C2 to C4 olefins. This feed stream for the second reaction stage is preferably a suitable stage for the production of relatively high olefin yields C2, C3 and C4. Such streams are those which boil in the range of naphtha and contain from about 5% by weight to about 35% by weight, preferably from about 10% by weight to about 30% by weight, more preferably from about 10% by weight to 25% by weight of paraffins and of about 15% by weight, preferably from about 20% by weight to about 70% by weight of olefins. The food can also contain nafteños and aromatics. Naphtha boiling range streams are typically those having a boiling range of about 18 ° C (65 ° F) to about 220 ° C (430 ° F), preferably about 18 ° C (65 ° F) at approximately 150 ° C (300 ° F). Naphtha streams from other sources in the refinery can be mixed with the previously mixed feed stream and fed to this second reaction stage. The second stage is carried out in a processing unit comprising a reaction zone, a depletion zone, a catalyst regeneration zone and a fractionation zone. The naphtha feed stream is fed into the reaction zone where it comes in contact with a source of regenerated, hot catalyst. The hot catalyst vaporizes and fractionates the feed at a temperature of about 500 ° C to 650 ° C, preferably about 500 ° C to 600 ° C. The fractionation reaction deposits carbonaceous hydrocarbons, or coke, on the catalyst, thus deactivating the catalyst. The fractionated products are separated from the catalyst with coke and sent to a fractionator. The coked catalyst is passed through the depletion zone where the volatile portions of the catalyst particles are depleted with steam. Depletion may be carried out under looser conditions in order to preserve adsorbed hydrocarbons for thermal equilibrium. The spent catalyst is then passed to the regeneration zone where it is regenerated by combustion of the coke in the catalyst in the presence of a gas containing oxygen, preferably air. The removal of the coke restores the catalyst activity and simultaneously heats the catalyst at a temperature of, for example, 650 ° C to 750 ° C. The hot catalyst is then recycled to the reaction zone to react with the fresh naphtha feed. The combustion gas formed by the combustion of the coke in the regenerator can be treated for the removal of particles and for the conversion of carbon monoxide, after which the combustion gas is normally discharged into the atmosphere. The fractionated products from the reaction zone are sent to a fractionation zone where several products are recovered, particularly fractions C2, C3 and C4. While attempts were made to increase the yields of light olefins in the FCC processing unit itself, the practice of the present invention employs its own separate processing unit, as previously described, that receives naphtha that comes from a suitable source in the refinery. The reaction zone is operated under process conditions that optimize the selectivity for C2 to C4 olefins, particularly propylene, with a relatively high conversion of C5 + olefins. Suitable catalysts for use in the second step of the present invention are the catalysts formed of a crystalline zeolite having an average pore diameter of less than about 0.7 nanometers (nm), said crystalline zeolite comprises from about 10 wt.% To about 50 % by weight of the total fluidized catalyst composition. It is preferred that the crystalline zeolite be selected from the family of crystalline aluminosilicates of medium pore sizes (<0.7 nm) also known as zeolites. Of particular interest are zeolites with medium pore sizes with a molar ratio between silica and alumina less than about 75: 1, preferably less than about 50: 1, and more preferably less than about 40: 1. The pore diameter (sometimes also known as the effective pore diameter) can be measured using standard adsorption techniques and hydrocarbon compounds of known minimum kinetic diameters. See, Breck, Zeoli te Molecular Sieves (Molecular sieves of zeolite), 1974 and Anderson et al., J. Catalysis 58, 114 (1979) both being incorporated herein by reference. Zeolites with medium pore sizes that can be employed in the practice of the present invention are described in "Atlas of Zeolite Structure Types" (Atlas of Zeolite Structure Types), W.H. Meier and D.H. Olson, Butterworth-Heineman, third edition, 1992, which was incorporated here by reference. Zeolites of medium pore sizes generally have a pore size of about 5A to about 7A and include, for example, zeolites of MFI, MFS, MEL, MTW, EUO, MTT, HEU, FER, and TON structure types (Nomenclature). of the IUPAC Zeolite Commission). Non-limiting examples of such pore-size mid-zeolites include ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-34, ZSM-35, ZSM-38, ZSM-48, ZSM-50, silicalite, silicalite 2. The most preferred zeolite is ZSM-5, which is described in U.S. Patent Nos. 3,702,886 3,770,614. ZSM-11 is described in U.S. Patent No. 3,709,979; ZSM-12 is described in U.S. Patent No. 3,832,449; ZSM-21 and ZSM-38 are described in U.S. Patent No. 3,948,758; ZSM-23 is described in U.S. Patent No. 4,076,842; and ZSM-35 s described in U.S. Patent No. 4,016,245. All of the above patents are incorporated herein by reference. Other zeolites of suitable medium pore sizes include the silicoaluminophosphates (SAPO), such as for example SAPO-4 and SAPO-ll which is described in US Pat. No. 4,440,871; chromosilicates; gallium silicates, iron silicate, aluminum phosphates (ALPO); as for example ALPO-11 which is described in the North American Patent No. 4,310,440; titanium aluminosilicates (TASO), as for example TASO-45 which is described in EP-A No. 229,295; borosilicate, which is described in the North American Patent No. 4,254,297; titanium aluminophosphates (TAPO), such as for example TAPO-11 which is described in US Pat. No. 4,500,651; and iron aluminosilicates. In one embodiment of the present invention, the Si / Al ratio of said zeolites is greater than about 40. Zeolites of medium pore sizes may include "crystal blends" which are considered to be the result of faults occurring within the crystal or area cristalin during, the synthesis of zeolites. Examples of crystalline mixtures of ZSM-5 and ZSM-11 are presented in the North American Patent No. 4,229,424 which is incorporated herein by reference. The crystalline mixtures are in themselves zeolites of medium pore sizes and should not be confused with physical mixtures of zeolites in which different crystallite crystals of different zeolites are physically present in the same catalyst compound or mixtures of hydrothermal reactions. The catalysts of the second stage of the present invention are linked with an inorganic oxide matrix component. The inorganic oxide matrix component binds the catalyst components together in such a way that the catalyst product has a hardness sufficient to survive collisions between particles and against the walls of the reactor. The inorganic oxide matrix can be made from a dried inorganic oxide sol or gel to "glue" the catalyst components together. Preferably, the inorganic oxide matrix is not catalytically active and comprises oxides of silicon and aluminum. It is also preferred that separate alumina phases are incorporated into the inorganic oxide matrix. Species of aluminum-g-alumina oxides, boehmite, diaspore, and transition aluminas such as α-alumina, b-alumina, g-alumina, d-alumina, e-alumina, k-alumina, and r-alumina can be employed. Preferably, the alumina species is an aluminum trihydroxide such as, for example, gibbsite, bayerite, nordstrandite or doyelite. Preferred second stage processing conditions include temperatures from about 500 ° C to about 650 ° C, preferably from about 525 ° C to about 600 ° C; partial hydrocarbon pressures of about 7,000 to 28,000 kg / m2, absolute pressure (of about 10 to 40 psia), preferably of about 14,000 to 24,500 kg / m2, absolute pressure (of about 20 to 35 psia); and the ratio between catalyst and naphtha (w / w) is from about 3 to 12, preferably from about 4 to 10, where the catalyst weight is the total weight of the catalyst compound. It is also preferred that the steam be introduced concurrently with the catalyst. the naphtha stream in the reaction zone, the vapor comprising up to about 50% by weight of the hydrocarbon feed. Also, it is preferred that the residence time of naphtha in the reaction zone be less than about 10 seconds, for example, from about 1 to 10 seconds. The above conditions will be such that at least about 60% by weight of the C5 + olefins in the naphtha stream become products and C4- and less than about 25 by weight, preferably less than about 20% by weight of the products. The paraffins are converted into C- products, and the propylene comprises at least about 90 mol% preferably more than about 95 mol% of the total C3 reaction products with the weight ratio between propylene and the C2-products greater than approximately 3.5. It is also preferred that the ethylene comprise at least about 90 mole% of the C2 products, with the weight ratio between propylene and ethylene greater than about 4, and that the "total range of the C5 + naphtha product be increased in both octane d motor and research in relation to the naphtha feed Within the scope of the present invention the catalysts of this second stage can be coked before the introduction of the feed in order to improve the selectivity for propylene. of the present invention is the feeding of an effective amount of single ring aromatics to the reaction zone of said second step to also improve the selectivity of propylene versus ethylene The aromatics can be from an external source such as for example reforming process or may consist of heavy naphtha recycling products of the present process The combustion gases of the first stage and second stage regenerator are combined in one embodiment of this invention, and the recovery section of light end products or products can also be shared with suitable modifications to the equipment. The high selectivity for the light olefin products desired in the second stage decreases the immersion required to adapt the installations for existing light end products for an additional recovery of light olefins. The catalyst composition of the first stage is typically selected to optimize hydrogen transfer. In this way, the second stage naphtha feed can be optimized for optimum yields of C2, C3, and C4 olefins with relatively high selectivity using the preferred second stage et operating conditions and catalyst. The total high-value light olefin products from the two combined stages included the products generated with a relatively low yield in the second stage plus the products generated with relatively high yield in the second stage. The following examples are presented for illustrative purposes only and should not be considered as limiting the present invention in any way. Examples 1-12 The following examples illustrate the critical aspect of the processing operation conditions to maintain a 'Purity of chemical grade propylene with samples of fractionated naphtha in ZCAT-40 (a catalyst containing ZSM-5) that has been vaporized at a temperature of 815 (1500 ° F) for 16 hours to simulate a commercial equilibrium. The comparison of examples 1 and 2 shows that increasing the cat / oil ratio improves propylene yield, but sacrifices the purity of propylene. The comparison of examples 3 and 4 and 5 and shows that the reduction of the partial pressure of petroleum greatly improves the purity of propylene without compromising propylene production. The comparison of the examples and 8 and 9 and 10 shows that a rising temperature improves both the production and the purity of the propylene.

Comparison of examples 11 and 12 shows that a shorter residence time of cat improves the yield and purity of propylene. Example 13 shows an example in which they obtain both a high propylene yield and a purity at a reactor temperature and with a ratio / oil that can be achieved using a conventional FCC reactor / regenerator design for the second stage. TABLE 1 Example Feeding Temperature Cat / Olefin oil, ° C% by weight 1 38.6 566 4.2 2 38.6 569 8.4 3 22.2 510 8.8 4 22.2 511 9.3 5 38.6 632 16.6 6 38.6 630 16.6 7 22.2 571 5.3 8 22.2 586 5.1 9 22.2 511 9.3 10 22.2 607 9.2 11 22.2 576 18.0 12 22.2 574 18.3 13 38.6 606 8.5 'oil psia) Oil Kg / m2 Res. of Oil, sec 1 36 25,311.6 Kg / m2 0.5 2 32 22,499.2 Kg / m2 0.6 3 18 12,655.8 Kg / m2 1.2 4 38 26,717.8 Kg / m2 1.2 5 20 14,062.0 Kg / m2 1.7 6 13 9,140.3 Kg / m2 1.3 7 27 18,983.7 Kg / m2 0.4 8 27 18,983.7 Kg / m2 0.3 9 38 26,717.8 Kg / m2 1.2 10 37 26,014.7 Kg / m2 1.2 11 32 22,499.2 Kg / m2 1.0 12 32 22,499.2 Kg / m2 1.0 13 22 15,468.2 Kg / m2 1.0; j Time of Res.% In weight% in pe: so Purity of Cat, sec C3 = C3_ Propylene, í 1 4.3 11.4 0..5 95.8% 2 4.7 12.8 0, .8 94.1% 3 8.6 8.2 1. .1 88.2% 4 5.6 6.3 1, .9 76.8% 9.8 16.7 1, .0 94.4% 6 7.5 16.8 0. .6 96.6% 7 0.3 6.0 0, .2 96.8% 8 0.3 7.3 0, .2 97.3% 9 5.6 6.3 1, .9 76.8% 10 6.0 10.4 2.2 82.5% 11 9.0 9.6 4.0 70.6% 12 2.4 10.1 1.9 84.2% 13 7.4 15.0 0.7 95.5% Example% by weight% by weight Relationship Relation% by weight C2"C2" between C3"between C3 * C3" and C2"and C2" 1 2.35 2.73 4.9 4.2 11.4 2 3.02 3.58 4.2 3.6 12.8 3 2.32 2.53 3.5 3.2 8.2 4 2.16 2.46 2.9 2.6 6.3 6.97 9.95 2.4 1.7 16.7 6 6.21 8.71 2.7 1.9 16.8 7 1.03 1.64 5.8 3.7 6.0 8 1.48 2.02 4.9 3.6 7.3 9 2.16 2.46 2.9 2.6 6.3 5.21 6.74 2.0 1.5 10.4 11 4.99 6.67 1.9 1.4 9.6 12 4.43 6.27 2.3 1.6 10.1 13 4.45 5.76 3.3 2.6 15.0 C2"= CH + C2H + C2H6 The above examples (1,2,7 and 8) show that C3" VC2 = > 4 C3_ / C2"> 3.5 can be achieved by selecting suitable conditions for the reactor, Examples 14-17 The catalytic fractionation of olefins and paraffins contained in streams of naphtha (eg, FCC naphtha, coke naphtha) in zeolites with small or medium pores such as, for example, ZSM-5 can produce significant amounts of ethylene and propylene.The selectivity for ethylene or propylene and the selectivity of propylene to propane varies depending on the operating conditions of the process and in function of the used catalyst It has been found that the yield of propylene can be increased by co-feeding a stream together with cat naphtha to the reactor.The catalyst can be ZSM-5 or other small or medium pore zeolites. below illustrates the increase in propylene yield when a stream of 5% by weight is co-fed with an FCC naphtha containing 38.8% by weight of lefines Even though the propylene yield was increased, the propylene purity decreased. Thus, adjustment of other operating conditions may be required to maintain the selectivity of white propylene. TABLE 2 Example Current Temperature Cat / Oil Co-fed ° C 14 No 630 8.7 15 Yes 631 8.8 16 No 631 8.7 17 Yes 632 8.4 Example (Psia oil) Oil Kg / m2 Res Pet time: oil, sec. 14 18 12,655.8 Kg / m2 0., 8 15 22 15,468.2 Kg / m2 1., 2 16 18 12,655.8 Kg / m2 0., 8 17 22 15,468.2 Kg / m2 1., 1 Example Res Time.% By weight% in Weight Purity of Cat, sec Propylene Propane Propylene,% 14 8.0 11.7 0.3 97.5% 6.0 13.9 0.6 95.9% 16"7.8 13.6 0.4 97.1% 17 6.1 14.6 0.8 94.8%

Claims (7)

1. CLAIMS A two-stage process for selectively producing C2 to C4 olefins from a heavy hydrocarbon feed, said process comprising: (a) Reacting said feed in a first step comprising a fluid catalytic fractionation unit where it is converted in the presence of a conventional large pore zeolitic catalytic fractionation catalyst in low boiling point reaction products; (b) fractionating said low boiling point reaction products into fractions of various boiling points, one of which is a fraction with a boiling range of naphtha; (c) reacting said boiling range fraction of naphtha in a second reaction step consisting of a processing unit formed by a reaction zone, a depletion zone, a catalyst regeneration zone, and a fractionation zone , wherein the boiling range fraction of naphtha is in contact in the reaction zone with a catalyst containing from about 10 to 50% by weight of a crystalline zeolite having an average pore diameter of less than about 0.7 nm under conditions of the reaction including temperatures within a range of about 500 to 650 ° C, and a partial hydrocarbon pressure of about 7,000 to 28,000 kg / m2, absolute pressure (of approximately 10 to 40 psia); 10 (d) passing the resulting vapor products and catalyst particles through the depletion zone where the volatile elements are depleted with steam; (e) passing the particles to an area of 15 regeneration where the catalyst coke is burned; and (f) recycling the hot catalyst particles to the reaction zone. 2. The procedure in accordance with the claim
2. Wherein the crystalline zeolite is selected from the group consisting of ZSM-5 and ZSM-11. 3. The process according to claim 2 wherein the naphtha feed contains from about 10 to 30% by weight of paraffins and of
3. About 15 to 70% by weight of olefins.
4. 4. The process according to claim 3 wherein the reaction temperature is from about 500 ° C to about 600 ° C.
5. 5. The process according to claim 3 wherein at least about 60% by weight of the C5 + olefins in the feed stream are converted to C < - and less than about 25% by weight of the paraffins are converted into C, - products.
6. 6. The process according to claim 5 wherein the propylene comprises at least about 90 mole% of the total C3 products.
7. 7. The process according to claim 6 wherein the weight ratio between propylene and the 15 products C2- is greater than approximately 3.5. twenty SUMMARY OF THE INVENTION A two-stage process for selectively producing C2 to C4 olefins from a gas oil or residue is presented. The gas oil or residue reacts in a first stage comprising a fluid catalytic fractionation unit where it is converted in the presence of a conventional large pore zeolitic catalyst into reaction products, including a stream within a boiling range of naphtha. The stream within a boiling range of naphtha is introduced into a second stage comprising a processing unit containing a reaction zone, a depletion zone, a catalyst regeneration zone, and a fractionation zone. The naphtha feed stream is in contact in the reaction zone with a catalyst containing from about 10 to 50% by weight of a crystalline zeolite having an average pore diameter of less than about 0.7 nanometers under reaction conditions including temperatures within a range of approximately 500 to 650 ° C and a partial hydrocarbon pressure of approximately 7, 000 to 28,000 kg / m2, absolute pressure (approximately 10 to 40 psia). The steam products are collected above and the catalyst particles are passed through the depletion zone towards the catalyst regeneration zone. The volatile elements are depleted with steam in the depletion zone and the catalyst particles are sent to the catalyst regeneration zone where the catalyst coke is burned, which is then recycled to the reaction zone.