MXPA00010668A - Process for selectively producing light olefins in a fluid catalytic cracking process from a naphtha/steam feed - Google Patents

Abstract

A process for selectively producing C2-C4 olefins from a catalytically cracked or thermally cracked naphtha stream. A mixture of the naphtha stream and a stream of steam is fed into a reaction zone where it is contacted 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 from about 500 to 650°C and a hydrocarbon partial pressure from about 10 to 40 psia.

Description

PROCESS FOR SELECTIVELY PRODUCING LIGHT OLEFINS IN A CATALYTIC FRACTIONATION PROCESS FLUID FROM A NAFTA / VAPOR FEED FIELD OF THE INVENTION The present invention relates to a process for selectively producing C2-C4 olefins from a stream of catalytically fractionated naphtha or thermally fractionated. A mixture of naphtha stream and vapor stream is fed into a reaction zone where it comes into contact 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 of about 500-650 ° C and under a partial hydrocarbon pressure of about 7,000 to 28,000 kg / m2 absolute pressure (10 to 40 psia). BACKGROUND OF THE INVENTION The need for fuel with low emissions has created an increased demand for light olefins for 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 feed for polyolefin, especially for the production of polypropylene. Fixed bed procedures for the dehydrogenation of light paraffin have recently attracted renewed interest to increase olefin production. However, these types of procedures typically require relatively large capital investments and present high operating costs. Therefore, it is advantageous to increase olefin yield using procedures that require relatively small capital investments. It would be particularly advantageous to increase the yield of olefins in catalytic fractionation processes. 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 riser tubes into which a different feed stream is introduced. The operation of the risers is designed in such a way that a suitable catalyst acts to convert a heavy gas oil into a rising tube and another suitable catalyst acts to fractionate a lighter olefin / naphtha feed into the other riser tube. 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. Also, U.S. Patent No. 5,026,936 to Arco presents a process for the preparation of propylene 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, U.S. Patent Nos. 5,026,935; 5,171,921 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 wherein the olefins are produced by reacting said feed in the presence of a 2SM-5 catalyst. An inherent problem in the production of olefin products using FCC units is that the process depends on a specific balance of catalyst to optimize the production of light olefins while also achieving a high conversion of the components fed at a temperature of 340 ° C (650 ° + F). further, even if a specific catalyst equilibrium can be maintained to optimize the overall olefin production, the olefin selectivity is generally low due to undesirable side reactions such as extensive fractionation, isomerization, aromatization and hydrogen transfer reactions. Light saturated gas produced from undesirable side reactions result in increased costs to recover desirable light olefins. Accordingly, it is desirable to optimize olefin production in a process that allows a high degree of control over the selectivity of C3 and C olefins. SUMMARY OF THE INVENTION In accordance with the present invention there is provided a process for the selective production of C2 to C4 olefins which comprises feeding a stream of catalytic or thermally fractionated naphtha containing paraffins and olefins and steam in a reaction zone and by the reaction of the naphtha with a catalyst containing from 10 to 50% by weight of a crystalline zeolite having an average pore diameter of less than about 0.7 nm under conditions including a temperature of about 500 ° C to 650 ° C, hydrocarbon partial pressure of 7,000 to 28,000 kg / m2, absolute pressure (from 10 to 40 psia), a hydrocarbon residence time of 1 to 10 seconds and a catalyst to feed ratio of approximately 2 to 10, where it does not become more than about 20 wt% of paraffin in olefins. In a preferred embodiment, a method is provided for selectively producing C2 to C4 olefins in a processing unit comprising a reaction zone, a depletion zone, and a catalyst regeneration zone. The naphtha stream comes into contact in the reaction zone, which contains a catalyst bed, preferably in a fluid state. The catalyst consists of a zeolite having an average pore diameter of less than about 0.7 nm and where the reaction zone is operated at a temperature of about 500 ° C to 650 ° C, under a partial hydrocarbon pressure of 7,000 to 28,000 kg. / m2, absolute pressure (from 10 to 40 psia), a hydrocarbon residence time of 1 to 10 seconds, and a catalyst to feed ratio of approximately 2 to 10, where no more than about 20% by weight of paraffins are It converts into olefins. In a preferred embodiment of the present invention, the crystalline zeolite is selected from the ZSM series.

In another preferred embodiment of the present invention the catalyst is a ZSM-5 type catalyst. In another preferred embodiment of the present invention, the 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 reaction zone operates at a temperature of about 525 ° C to about 600 ° C. DETAILED DESCRIPTION OF THE INVENTION Suitable feed streams to produce the relatively high yields of olefin C2, C3 and C4 are streams boiling in the range of naphtha and containing from about 5% by weight to about 35% by weight, preferably from about 10% by weight to about 30% by weight, and more preferably from about 10 to 25% by weight of paraffin, and from about 15% by weight, preferably from about 20% by weight to about 70% by weight olefin weight. The food can also contain nafteños and aromatics. Naphtha boiling range streams are typically streams having a boiling range of about 18 ° C to about 221 ° C, preferably about 18 ° C to about 149 ° C. Naphtha can be thermally fractionated or catalytically fractionated. Such streams can be derived from any suitable source, for example, they can be derived from the fluid catalytic fractionation (FCC) of gas oils and residues, or they can be derived from the delayed or fluid coking of residues. It is preferred that the naphtha streams employed in the practice of the present invention are derived from the fluid catalytic fractionation of gas oil and waste. Such naphthas are typically rich in olefins and / or diolefins and relatively poor in paraffins. The process of the present invention 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 to the reaction zone as a mixture of naphtha and steam, where it comes into contact with a hot, regenerated catalyst source. The hot catalyst vaporizes and fractionates the feed at a temperature within a range of about 500 ° C to 650 ° C, preferably about 525 ° C to 600 ° C. The fractionation reaction deposits carbonaceous hydrocarbons, or coke, in the catalyst, thus deactivating the catalyst. The fractionated products are separated from the coked catalyst and sent to a fractionator. The coked catalyst is passed through the depletion zone where volatile substances of the catalyst particles are depleted with steam. The depletion can be carried out under conditions of low severity in order to conserve adsorbed hydrocarbons for thermal equilibrium. The spent catalyst is then passed to the regeneration zone where it is regenerated by the 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 from about 650 ° C to about 750 ° C. The hot catalyst is then recycled to the reaction zone to react with a 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 a C fraction, a C4 fraction rich in olefins, and a Cs fraction rich in olefins. The amount of steam co-fed with the naphtha feed stream will typically be within a range of about 10 to 250 mol%, preferably, within a range of about 25 to 150 mol% between steam to naphtha. 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, which receives naphtha coming 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. Catalysts suitable for use in the second step of the present invention are shaped catalysts of a crystalline zeolite having an average pore diameter of less than about 0.7 nanometers (nm), said crystalline zeolite comprising 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.mFI Zeolites with medium pore sizes that can be employed in the practice of the present invention are described in "Atlas of Zeolite Structure Types", editions .H. Meier and D.H. Olson, Butterworth-Heineman, third edition, 1992, which is incorporated herein by reference. Zeolites of medium pore sizes generally have a pore size of about 5 A to about 7 A and include, for example, zeolites of MFS, MEL, MTW, EUO, MTT, HEU, FER, and TON type structures (Nomenclature). of the IUPAC Zeolite Commission). Non-limiting examples of such mid-pore size zeolites include ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-34, ZSM-35, ZSM-38, ZSM-48, ZSM-50, silicalite, and silicalite 2. The most preferred zeolite is ZSM-5, which is described in U.S. Patent Nos. 3,702,886 and 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 is 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 silicoaluminophosphates (SAPO), such as SAPO-4 and SAPO-11, which is described in US Pat. No. 4,440,871; cro osilicates; gallium silicates, iron silicates, aluminum phosphates (ALPO); as for example ALPO-11 that is described in the North American Patent No. 4,310,440; titanium aluminosilicates (TASO), such as for example TASO-45 which is described in EP-A No. 229,295; borosilicate, which is described in U.S. Patent No. 4,254,297; titanium aluminophosphate (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 mixtures" which are considered to be the result of faults occurring within the crystal or area crystalline during the synthesis of zeolites. Examples of crystalline mixtures of ZSM-5 and ZSM-11 are presented in U.S. Patent No. 4,229,424 which is incorporated herein by reference. The crystalline mixtures are themselves zeolites of medium pore sizes and should not be confused with physical mixtures of zeolites in which crystals other than crystallites of different zeolites are physically present in the same catalyst compound or in 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, boehmite, diaspore, and transition aluminas such as α-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 phosphorous or aluminum phosphate. Preferred 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 weight of the catalyst is the total weight of the catalyst compound. It is also preferred that the steam is introduced concurrently with the stream of naphtha 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, 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 will be converted into products and C4- and less than about 25% by weight, preferably less than about 20% by weight of the paraffins are converted to C-products, and that 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 C2- products greater than about 3.5. It is also preferred that the ethylene comprises 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 as much. in motor octane and research in relation to the supply of naphtha. Within the scope of the present invention, the catalysts can be coked prior to the introduction of the feed in order to improve the selectivity for propylene. Within the scope 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 may be from an external source such as for example reforming processing unit or may consist of heavy naphtha recycling products of the present process. The following examples are presented for illustrative purposes only and should not be construed as limiting the present invention in any way. Examples 1-12 The following examples illustrate the critical aspect of the processing operation conditions for maintaining a purity of chemical grade propylene with samples of naphtha catheter fractionated in ZCAT-40 (a catalyst containing ZSM-5) that has been vaporized at a temperature of 815 ° C (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 the propylene yield, but sacrifices the propylene purity. The comparison of examples 3 and 4 and 5 and 6 shows that the reduction of the partial pressure of oil greatly improves the purity of the propylene without compromising propylene production. The comparison of examples 7 and 8 and 9 and 10 shows that a rising temperature improves both the production and the purity of the propylene. The comparison of examples 11 and 12 shows that a shorter cat residence time improves the yield and purity of the propylene. Example 13 shows an example in which both a high propylene yield and a high purity are obtained at a reactor temperature and with a cat / oil ratio 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 Example (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 Example Res Time,% by weight% by weight Purity of Cat, sec C3"C3" Propylene,% 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% 5 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%% eenn ppeeso Relation 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"= CH4 + C2H4 + C2H6 The above examples (1,2,7 and 8) show that C3 = / C2 => 4 and 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 naphtha streams (e.g., 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 for propane varies depending on the operating conditions of the process and depending on the catalyst used. 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. Table 2 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 olefins. Even when the yield of propylene 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 ((PPeettrróólleeoo pps: ia) Oil Kg / m2 Oil Res Time, 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% wt% wt Purity of Cat, sec Propylene Propane Propylene,% 14 8.0 11.7 0.3 97.5% 15 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 (3)

1. CLAIMS A process for the selective production of C2 to C4 olefins comprising the feeding of a stream of catalytic or thermally fractionated naphtha containing paraffins and olefins and steam to a reaction zone and the reaction of the naphtha with a catalyst containing to 50% by weight of a crystalline zeolite having an average pore diameter of less than about 0.7 nm in conditions including a temperature of about 500 ° C to 650 ° C, a hydrocarbon partial pressure of 7,000 to 28,000 kg / m2, absolute pressure (from 10 to 40 psia), a hydrocarbon residence time of 1 to 10 seconds, and a catalyst to feed ratio of approximately 2 to 10, wherein no more than about 20% by weight of paraffins are converted to olefins . The procedure in accordance with the claim 1 wherein the amount of steam that is fed into the reaction zone with the naphtha feed is about 1 to 50 mol%. The procedure in accordance with the claim 1 wherein the crystalline zeolite is selected from the ZSM series. The process according to claim 3 wherein the crystalline zeolite is ZSM-5. The process according to claim 3 wherein the naphtha feed contains from about 10 to 30% by weight of paraffins and from about 15 to 70% by weight of olefins. The process according to claim 5 wherein the reaction temperature is from about 500 ° C to about 600 ° C. The process according to claim 6 wherein at least about 60% by weight of the C5 + olefins in the feed stream are converted into C4- products and less than about 25% by weight of the paraffins are converted into C4- products. The procedure in accordance with the claim
2. Wherein the propylene comprises at least about 90 mole% of the total C3 products. The procedure in accordance with the claim 8 where the weight ratio between propylene and total C? -products is greater than about
3. 3. 5. The procedure in accordance with the claim 9 wherein the amount of steam feed in the reaction zone with the naphtha feed is from about 2 to 20 mol%.