JP2012519747A - Method for preventing coking by metal catalyst - Google Patents

Abstract

Disclosed is a method and apparatus in which a sulfiding agent is added to a catalytic conversion reactor to prevent metal catalyzed coking. The catalytic reactor is C ₁₀ downstream of catalytic reactor _- may be downstream of the first fluid catalytic cracking reactor for supplying hydrocarbon as the feed.
[Selection] Figure 1

Description

  The present invention relates generally to an apparatus and method for producing desired products such as light olefins including propylene.

  Fluid catalytic cracking (FCC) is a hydrocarbon catalytic conversion process carried out by contacting a relatively heavy hydrocarbon in a fluidization reaction zone with a catalyst particle material. Unlike the hydrogenolysis, the reaction in the catalytic cracking is carried out without substantial hydrogen addition or hydrogen consumption. As the cracking reaction proceeds, a large amount of high carbonaceous material called coke is deposited on the catalyst to produce a coking catalyst, that is, a used catalyst. Volatile light products are separated from the spent catalyst in the reaction vessel. The used catalyst may be volatilized with an inert gas such as water vapor to volatilize the hydrocarbon gas taken into the used catalyst. Coke is burned from spent catalyst that has been stripped by high temperature regeneration with oxygen in the regeneration zone operation. Such a process produces a variety of products including gasoline products and / or light products such as propylene and / or ethylene.

  In such a process, a single reactor or a dual reactor can be used. Additional investment costs are incurred by using a dual reactor unit, but one of the dual vessels is adjusted to maximize products such as light olefins including propylene and / or ethylene Can be operated with.

  Maximizing the yield of product in one of the reactors is often advantageous. Further, there may be a desire to maximize production of product from one reactor that can be recycled to the other reactor producing the desired product, such as propylene.

The focus of FCC technology development over the past few years has been mainly to maximize propylene selectivity. This has led many FCC technology licensors to promote dual riser FCC technology grants. In this technique, a primary feed, typically VGO, is fed to a first riser and a recirculated stream of C ₁₀ − or a portion thereof is recycled to a second riser. In this way, the first riser and the second riser can be operated in different ways, maximizing overall net selective yield. In typical operation, the operation of the first riser is not as severe as the second riser. The second riser is operated much more rigorously, preferably at relatively high temperatures, typically in the range of 538 ° C. to 593 ° C. (1000 ° F. to 1100 ° F.), and less than 138 kPa (absolute) (20 psia) Promotes the formation of light olefins such as butylene, propylene, and ethylene at low hydrocarbon partial pressures. Feed to the second riser C ₁₀ from FCC recycle or other processing device _- may be a material.

  Those who commercialized dual riser technology that recirculates naphtha to the second riser have been plagued by excessive coke formation in the second riser, limiting the operational capacity of these processes. In the known case, operation was limited to weeks instead of months, and the equipment had to be shut down to remove coke. Thus, it is necessary to provide a catalytic cracking twin reactor apparatus that prevents excessive coke from forming in the second riser.

  The inventors have discovered that excessive coking in the second reactor is due to metal catalyzed coking (MCC). MCC is suppressed in conventional FCC devices because there are enough sulfur species in the hydrocarbons supplied to the FCC device to decompose in the FCC riser to form hydrogen sulfide. The hydrogen sulfide subsequently passivates the active metal in the FCC unit. We propose a method and apparatus for adding a sulfiding agent to an FCC riser or other reactor when hydrogen sulfide is not present enough to inhibit MCC. The sulfur species of the sulfiding agent may be provided as hydrogen sulfide or provide a source of hydrogen sulfide by decomposition, liberation, or other chemical reaction to subsequently form a metal sulfide on the inner metal surface inside the reactor. . The metal sulfide layer isolates the gas phase coke precursor from the active metal sites on the inner surface and prevents coking.

FIG. 1 is a schematic diagram of the present invention.

  MCC is characterized by the deposition of carbonaceous solids on the hot metal surface and is manifested in the process at temperatures higher than 400 ° C., with the highest rate of filamentous carbon formation in the range of 550 ° C. to 600 ° C. MCC is a pyrolysis or catalytic reaction with active metals that has considerable impact on many industrial processes including processes involving steam catalytic reforming of methane, steam cracking of paraffin feedstocks, and carbon monoxide disproportionation. Effect. It is well known that certain metals can increase the overall MCC deposition rate by catalyzing the growth of filamentous and graphitic deposits. The highest catalysis for carbon deposition is demonstrated by iron, cobalt, and nickel, and alloys containing these metals. The overall catalytic reaction pathway for MCC is generally considered to be the adsorption of ethylene, propylene, or butylene onto the metal surface. The adsorbed light olefin is then subjected to further dehydrogenation conversion to become aromatic or alkylaromatic, which are further condensed to form coke.

  A typical FCC reaction is operated in the range of 500 ° C. to 600 ° C. to meet the maximum reaction rate of filamentous carbon formation. The metals that are known to most actively promote MCC are present in FCC equipment. Active hydrocarbon species that promote filamentous carbon formation are ethylene, propylene, and butylene, which are target products of FCC technology that produces propylene at a high rate. Therefore, the present inventors consider that the problem of coking in the second FCC riser process is caused by MCC.

  MCC has never been observed in FCC operations. Many FCC units provide a feed containing a substantial amount of sulfur, typically 0.1% to 1.0% by weight. Sulfur present in the FCC feed decomposes into hydrogen sulfide and is adsorbed on the metal surface to form a metal sulfide layer, isolating the gas phase coke precursor from the active metal sites on the internal surface of the FCC reactor, thereby Reduce coke formation. The inventors have found that in the recycle stream, the hydrogen sulfide generated by the decomposition of the first FCC feed is not normally present in the naphtha feed that is recycled to the second FCC riser. Organic sulfur in the main FCC product is preferentially allocated to hydrogen sulfide and coke in the reaction product, followed by priority to heavier products, to naphtha and liquefied petroleum gas (LPG). Only leaves a minimal amount of sulfur. In the second riser treating naphtha, the naphtha is greatly deficient in mixed sulfur, resulting in insufficient formation of a sulfide layer on the metal in the second riser to prevent MCC. Even if sulfur is present in the naphtha, it does not form a layer that passivates the active metal that contributes to MCC unless it is thermally decomposed to form hydrogen sulfide.

We propose adding a sulfurizing agent to the catalytic reactor to prevent MCC from causing coke problems that are resident in the second reactor. The sulfiding agent may be hydrogen sulfide or an organic sulfur compound that decomposes into hydrogen sulfide in a catalytic conversion environment and particularly in a fluid catalytic cracking environment. The hydrogen sulfide can be provided to the dry gas supplied to the second reactor prior to the amine treatment. Hydrogen sulfide can also be supplied by adding a SO _x scavenging additive such as magnesium aluminum oxide having a commercially available spinel structure to the circulating catalyst inventory. The additive adsorbs SO _x in the oxidizing environment of the regenerator, the reactor riser of a reducing environment to desorb the hydrogen sulfide. However, the technical ability to provide sufficient hydrogen sulfide capacity in the second reactor using the SO _x additive is highly dependent on the sulfur content of the feed to the first reactor. Preferred organic sulfur sources include methyl sulfides, mercaptans, and polysulfides such as dimethyl sulfide (DMS) or dimethyl disulfide (DMDS) conventionally used in industrial operations as sulfiding agents for hydrotreaters and dry distillation furnaces. And a commercially available sulfurizing agent. These organic sulfur decomposes to hydrogen sulfide in fluid catalytic cracking and other reaction environments. Sulfur-containing oils in FCC products such as LCO, HCO, and CSO are not preferred sulfiding agents. This is because they cannot be expected to generate the amount of hydrogen sulfide necessary to effectively pyrolyze and passivate the active metal. However, these heavy FCC products may be effective under certain conditions. Lighter FCC products such as naphtha and LPG can also be effective sulfiding agents under certain conditions where sulfide compounds are not removed from them.

  A dry gas comprising hydrogen sulfide is preferably added to the fluidizing gas distributor or added to the riser reactor distributor as an atomizing dispersion medium. The organosulfur sulfiding agent may be added to the fluidizing gas distributor or preferably to the supply system at the upstream point of the supply distributor. The maximum sulfur proportion is not limited, but a range of 20 wppm to 2000 wppm is suitable for the fluid present in the reactor, preferably in the range of 50 wppm to 500 wppm. The sulfiding agent should be added continuously. This is because coking begins very quickly and sulfides are continuously adsorbed and desorbed on active metals.

  The present invention can be described with reference to four components: the first reactor 10, the regeneration vessel 60, the product fractionation section 90, and the second reactor 170. Although the present invention can have many configurations, a specific embodiment is presented as an example in this specification. All other possible embodiments for carrying out the invention are considered to fall within the scope of the invention. For example, if the first and second reactors 10 and 170 are not FCC reactors, one or both of the regeneration vessel 60 and the product separation unit 90 may not be essential. Furthermore, the present invention can be embodied in a single FCC reactor 170.

  Although the figure shows a first reactor 10, the reactor 10 may be an FCC reactor including a first reactor riser 12 and a first reaction vessel 20. The regenerator catalyst tube 14 communicates with the first reactor riser 12 upstream so that material flows from the regenerator catalyst tube 14 to the first reactor riser 12. Communication means that material flow is possible between the listed areas. The regenerator catalyst tube 14 adjusts the speed of the regenerated catalyst supplied from the regenerative vessel 60 by the control valve 16 and then supplies the regenerated catalyst to the reactor riser 12 from the regenerated catalyst inlet. A fluid medium such as water vapor from the distributor 18 drives the flow of regenerated catalyst upward in the first reactor riser 12 at a relatively high density. A plurality of feed distributors 22 communicate with the first reactor riser 12 upstream to inject the first hydrocarbon feed 8 with the inert atomizing gas, preferably steam, across the flow of catalyst particles. The hydrocarbon feed to the first reactor riser 12. When the hydrocarbon feed comes into contact with the catalyst in the first reactor riser 12, heavy hydrocarbons decompose to produce a light gaseous first cracked product, while converted coke and mixed coke precursors. Deposits on catalyst particles to form a coking catalyst.

  Conventional FCC feed and high boiling hydrocarbon feed are suitable first feeds 8 to the first FCC reactor. The most common of such conventional feedstocks is “vacuum gas oil” (VGO), typically a hydrocarbon material having a boiling range of 343 ° C. to 552 ° C. (650 ° F. to 1025 ° F.). And is prepared by vacuum fractionation of atmospheric residue. Such fractions are generally less contaminated with heavy metals that can contaminate the coke precursor and catalyst. Heavy hydrocarbon feedstock to which the present invention can be applied includes heavy bottom from crude oil, heavy bituminous crude oil, shale oil, tar sand extract, defoaming residue, coal liquefaction product, atmospheric distillation and vacuum distillation Of residual oil. The heavy feed in the present invention also includes a mixture of the above hydrocarbons, and the above list is not exhaustive. Usually the first feed 8 has a temperature of 140 ° C to 320 ° C. In addition, additional quantities of feed can be introduced downstream of the initial feed point.

  The first reaction vessel 20 communicates with the first reactor riser 12 downstream so that material flows from the reactor riser 12 to the first reaction vessel 20. The resulting gaseous product hydrocarbon and spent catalyst mixture travels upward through the first reactor riser 12 and is received by the first reactor vessel 20 where the spent catalyst and gaseous product are received. Are separated. A pair of release arms 24 is a release vessel that partially separates the gas from the catalyst from the top of the first reactor riser 12 via one or more outlets 26 (only one shown). 28 is discharged tangentially and horizontally. The transport conduit 30 carries hydrocarbon vapors including volatilized hydrocarbons, volatilization media, and entrained catalyst to one or more cyclones 32 in the first reaction vessel 20 that separates the spent catalyst from the hydrocarbon gas product stream. The release vessel 28 is partially disposed within the first reaction vessel 20 and is considered part of the first reaction vessel 20. The gas conduit 34 feeds the separated hydrocarbon gas stream from the cyclone 32 to the recovery plenum 36 in the first reaction vessel 20, which stream passes through the outlet nozzle 38 and through the product line 88 and finally product recovery. Into the product fractionation section 90 for The dipleg 40 discharges the catalyst from the cyclone 32 to the low bed 42 in the first reaction vessel 20. The catalyst with adsorbed or mixed hydrocarbons can finally pass through the opening 46 defined by the wall of the release vessel 28 from the low bed 42 into the volatilization portion 44 provided as needed. The catalyst separated in the separation container 28 can be directly passed through the bed 48 and into the volatilization part 44 provided as necessary. The fluidizing distributor 50 supplies an inert fluidizing gas, usually water vapor, to the volatilization section 44. Volatilization section 44 includes baffle 52 or other equipment that facilitates contact between stripping gas and the catalyst. The spent catalyst after the volatilization treatment is separated from the volatilization portion 44 of the separation container 28 of the first reaction vessel 20 together with a lower concentration of admixture or adsorbed hydrocarbon than when the volatilization treatment is entered or not subjected to the volatilization treatment. The spent catalyst is preferably volatilized and leaves the detachment vessel 28 of the first reaction vessel 20, passes through the spent catalyst conduit 54, and proceeds into the regeneration vessel 60 at a rate regulated by a slide valve 56.

  The first reactor riser 12 can be operated at any suitable temperature and is typically operated at a temperature of 150 ° C. to 580 ° C., preferably 520 ° C. to 580 ° C. at the riser outlet 24. In a typical embodiment, a relatively high riser temperature is desirable, for example, higher than 565 ° C. at riser outlet 24, pressure is 69 kPa to 517 kPa (gauge) (10 psig to 75 psig), but typically 275 kPa (gauge) (40 psig). Is less than. The ratio of catalyst to oil is in the range of up to 30: 1 based on the weight of catalyst and feed hydrocarbon entering the bottom of the riser, but typically between 4: 1 and 10: 1, 7: 1 and 25 Range between 1 and 1. Normally, hydrogen is not added to the riser. Steam can pass through the first reactor riser 12 and the first reaction vessel 20 in an amount equivalent to 2% to 35% by weight of the feed. However, typically the proportion of water vapor is between 2% and 7% by weight of the maximum gasoline production and between 10% and 15% by weight of the maximum light olefin production. The average residence time of the catalyst in the riser may be less than 5 seconds.

  The catalyst in the first reactor 10 may be a single catalyst or a mixture of different catalysts. Typically, the catalyst comprises two components, a catalyst, a first component catalyst and a second component catalyst. Such a catalyst mixture is disclosed, for example, in US Pat. No. 7,312,370B2. In general, the first component can be any well-known catalyst used in FCC technology, such as an active amorphous clay type catalyst and / or a highly active crystalline molecular sieve. Zeolites can be used as molecular sieves in the FCC process. Preferably, the first component includes a large pore size zeolite such as Y-type zeolite, an activated alumina material, a binder material containing either silica or alumina, and an inert filler such as kaolin.

  Typically, a zeolite molecular sieve suitable for the first component has a large average pore size. Usually, a molecular sieve having a large pore size has pores having openings having an effective diameter defined by a ring larger than a 10-membered ring, typically a 12-membered ring, greater than 0.7 nm. The pore size index of large pores exceeds 31. Suitable large pore zeolite components include synthetic zeolites such as X and Y zeolites, mordenite, and faujasite. A portion of the first component, such as zeolite, can have any suitable amount of rare earth metal or rare earth metal oxide.

  The second component may include a medium or small pore zeolite catalyst, such as ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48. MFI zeolite exemplified by at least one of the above and other similar materials. Other suitable medium or small pore zeolites include ferrierite and erionite. Preferably the second component has a medium or small pore size zeolite dispersed in a matrix comprising a binder material such as silica or alumina and an inert filler material such as kaolin. The second component may also include other active materials such as beta zeolite. These compositions may have a crystalline zeolite content of 10 wt% to 50 wt% or more and a matrix content of 50 wt% to 90 wt%. A component containing 40% by weight crystalline zeolitic material is preferred, and components containing more crystalline zeolitic material may be used. In general, medium and small pore zeolites are characterized by having an effective pore opening diameter of 0.7 nm or less, a ring of 10 members or less, and a pore size index of less than 31, wherein the second catalyst component is MFI zeolite with a silicon to aluminum ratio of greater than 15, preferably greater than 75. In an exemplary embodiment, the silicon to aluminum ratio may be 15: 1 to 35: 1.

  The total mixture in the first reactor 10 may contain 1% to 25% by weight of the second component, ie, medium to small pore crystalline zeolite, the second component being 1.75% by weight. % Or more is preferable. When the second component is 40% by weight crystalline zeolite and the balance is a binder, an inert filler such as kaolin, and optionally an activated alumina component, the mixture is from 4% to 40% by weight of the second. The preferred content is at least 7% by weight. The first component may include the remainder of the catalyst composition. In certain preferred embodiments, the relative proportions of the first and second components in the mixture may not vary substantially throughout the first reactor 10. A high concentration of medium or small pore zeolite as the second component of the catalyst mixture improves selectivity to light olefins. In a typical embodiment, the second component is a ZSM-5 zeolite and the mixture is 4 wt% to 10 wt% ZSM-5 zeolite and does not contain other components such as binders and / or fillers. .

The regeneration container 60 communicates with the first reaction container 20 downstream. In the regeneration container 60, a part of the used catalyst coke supplied to the regeneration container 60 is brought into contact with an oxygen-containing gas such as air and burned to generate a regeneration catalyst. The regenerator 60 may be a combustor type regenerator as shown in the figure, and the mixed turbulent bed type high-speed fluidization condition is used in the high-efficiency regenerator 60 for completely regenerating the spent catalyst. Can do. However, other regeneration containers and other flow conditions are also suitable for the present invention. The spent catalyst conduit 54 supplies spent catalyst to the first or lower chamber 62 defined by the outer wall via the spent catalyst inlet. The spent catalyst from the first reaction vessel 20 usually contains carbon in an amount of 0.2 wt% to 2 wt%, which carbon is present in the form of coke. Coke is primarily composed of carbon, but may contain sulfur and other materials in addition to 3-12% hydrogen by weight. The oxygen-containing combustion gas, usually air, enters the lower chamber 62 of the regeneration vessel 60 through a conduit and is distributed by a distributor 64. As the combustion gas enters the lower chamber 62, it contacts the spent catalyst entering from the spent catalyst conduit 54 and causes the catalyst to enter the lower chamber 62, possibly at least 1.1 m / s (3.5 ft / s), under fast fluidization flow conditions. Lift the catalyst at the superficial velocity of the combustion gas. In one aspect, the lower chamber 62 has a catalyst density of 48 kg / m ^{3 to} 320 kg / m ³ ( ³ lb / ft ^{3 to} 20 lb / ft ³ ) and 1.1 m / s to 2.2 m / s (3.5 ft / s to 7 ft / s) superficial gas velocity. Oxygen in the combustion gas contacts the spent catalyst and burns the carbonaceous deposits of the catalyst, at least partially regenerating the catalyst and generating combustion exhaust gas.

The mixture of the catalyst and the combustion gas in the lower chamber 62 passes through the conical transition 66 and rises to the transfer riser portion 68 of the lower chamber 62. The riser portion 68 is preferably cylindrical and is partitioned by a tube extending preferably from the lower chamber 62 upward. The catalyst and gas mixture moves at a higher superficial gas velocity than in the lower chamber 62. The gas velocity is increased because the cross-sectional area of the riser portion 68 is smaller than the cross-sectional area of the lower chamber 62 below the transition portion 66. Therefore, the superficial gas velocity is usually higher than 2.2 m / s (7 ft / s). The riser portion 68 has a catalyst density smaller than 80 kg / m ³ (5 lb / ft ³ ).

The regeneration container 60 also includes an upper chamber or second chamber 70. A mixture of the catalyst particles and the combustion exhaust gas is discharged from the upper portion of the riser portion 68 into the upper chamber 70. Although substantially fully regenerated catalyst can exit from the top of the transfer riser portion 68, a configuration in which partially regenerated catalyst exits the lower chamber 62 is also contemplated. Discharge is through a detachment device 72 that separates most of the regenerated catalyst from the flue gas. In one embodiment, the catalyst and gas flowing upward through the riser section 68 collide with the elliptical lid on the top of the riser section 68 and back flow. The catalyst and gas then exit through the downward outlet of the detachment device 72. Sudden loss of momentum and downward backflow causes most of the heavy catalyst to fall into the high density catalyst bed 74 and causes a small portion of the catalyst still mixed with light flue gas to rise upward in the upper chamber 70. Cyclones 75, 76 further separate the catalyst from the rising gas and deposit the catalyst on the dense catalyst bed 74 through dipregs 77, 78. The flue gas exits the cyclones 75, 76, passes through the gas conduit, collects in the plenum 82, and enters the outlet nozzle 84 of the regeneration vessel 60 and possibly the flue gas or power recovery system (not shown). The catalyst density of the dense catalyst bed 74 is typically maintained within the range of 640 kg / m ^{3 to} 960 kg / m ³ (40 lb / ft ^{3 to} 60 lb / ft ³ ). The fluidizing conduit feeds fluidizing gas, typically air, through fluidizing distributor 86 to the dense catalyst bed 74. In one aspect, the heat regeneration catalyst from the dense catalyst bed 74 in the upper chamber 70 can be recirculated to the lower chamber 62 via the recirculation conduit 80 to facilitate coke combustion in the lower chamber 62. .

  The regeneration vessel 60 typically requires 14 kg of air per kg of removed coke for complete regeneration. A larger amount of feed may be processed in the first reactor 10 when more catalyst is regenerated. The regeneration container 60 typically has a temperature of 594 ° C. to 704 ° C. (1100 ° F. to 1300 ° F.) in the lower chamber 62 and a temperature of 649 ° C. to 760 ° C. (1200 ° F. to 1400 ° F.) in the upper chamber 70. Have. The regeneration catalyst tube 14 communicates with the regeneration container 60 downstream. The regenerated catalyst from the high density catalyst bed 74 is transferred through the regenerated catalyst tube 14 from the regeneration vessel 60 through the control valve 16 and back to the first reactor riser 12, where it is fed again when the FCC process continues. Contact objects.

  Further, in one desirable embodiment, the first reactor 10 can be operated at a low hydrocarbon partial pressure. Generally, when the hydrocarbon partial pressure is low, the production of light olefins is promoted. Therefore, the pressure in the first reactor riser 12 may be 170 kPa to 250 kPa, and the hydrocarbon partial pressure may be 35 kPa to 180 kPa, preferably 70 kPa to 140 kPa. A relatively low hydrocarbon partial pressure can be achieved using water vapor in the amount of 10% to 55%, preferably up to 15% by weight of the feed as diluent. Equivalent hydrocarbon partial pressures can also be achieved using other diluents such as dry gas.

  The first cracked product in line 88 from the first reactor 10 is relatively free of catalyst particles, contains volatilizing fluid, and exits the first reaction vessel 20 through the outlet nozzle 38. The first cracked product stream in line 88 may be further processed to further treat the stream prior to removing or separating fine catalyst particles. Line 88 moves the first cracked product stream to product fractionation section 90, which may include main column 100 and gas concentrating section 114 in one aspect. Various products are recovered from the main column 100. In this case, the main column 100 collects an overhead stream of light products containing unstable gasoline and light gases in the overhead line 102. The overhead stream in the overhead line 102 is condensed in the condenser 104 and cooled in the cooler 106 before entering the receiver 108. Line 110 collects a light off-gas stream from receiver 108. Off-gas includes LPG and dry gas. The dry gas contains hydrogen sulfide that functions as a sulfiding agent. The light liquid bottom liquid stream leaves receiver 108 via line 112. Both lines 110 and 112 can be supplied to the gas condensing section 114. In the gas condensing unit 114, a large number of streams are separated by separation or the like to produce a light olefin line 116, a light naphtha line 118, and a dry gas line 120. The dry gas stream can be primarily condensed into a hydrogen sulfide stream or can be part of a more comprehensive stream, but is represented by the dry gas stream 120. At least a portion of the dry gas stream is directed to a recirculating dry gas sulfidant line 122 and supplied to a dry gas mixed sulfidant line 124 and / or a dedicated dry gas sulfidant line 184. The main column 100 also feeds a heavy naphtha stream, a light circulating oil (LCO) stream, and a heavy circulating oil (HCO) stream through lines 126, 128, and 130, respectively. Some of the flow in lines 126, 128, and 130 circulate all through heat exchangers 132, 134, and 136 and reflux loops 138, 140, and 142, respectively, to remove heat from main column 100. Heavy naphtha, LCO, and HCO streams are transferred from the main column 100 through lines 144, 146, and 148, respectively. The cracked residual oil (CO) fraction can be recovered from the bottom of the main column 100 via line 150. A portion of the CO fraction is recirculated via a heat exchanger 152 and returned to the main column 100 via line 154. The CO stream is removed from main column 100 via line 156.

The light naphtha fraction preferably has an initial boiling point (IBP) of less than 127 ° C. (260 ° F.) in the C ₅ range, ie 35 ° C. (95 ° F.), and a temperature of 127 ° C. (260 ° F.) or more. Has an end point (EP). The boiling points of these fractions are measured using a procedure known as ASTM D86-82. A portion of the light naphtha stream in light naphtha line 118 is withdrawn in line 156 for further processing or storage, and the other portion of the feed in feed line 158 regulated by a control valve is the second reactor. The feed to 170 is sent to the recycle feed line 166. The heavy naphtha fraction has an IBP greater than 127 ° C (260 ° F) and a temperature higher than 200 ° C (392 ° F), preferably between 204 ° C and 221 ° C (400 ° F and 430 ° F), especially 216 ° C. It has an EP of (420 ° F). A portion of the heavy naphtha stream in line 144 is recovered in line 160 for further processing or storage, and another portion in line 162 regulated by a control valve is fed to the second reactor 170. To the recycle feed line 166 for recirculation. The LCO stream has IBP at the EP temperature of heavy naphtha, the EP is in the range of 260 ° C. to 371 ° C. (500 ° F. to 700 ° F.), preferably 288 ° C. (550 ° F.). The HCO stream has an IBP at the EP temperature of the LCO stream and the EP is in the range of 371 ° C. to 427 ° C. (700 ° F. to 800 ° F.), preferably 399 ° C. (750 ° F.). The CO stream has an IBP at the EP temperature of the HCO stream and includes everything with a higher temperature boiling point.

  Further, the product recovery unit 90 performs rough separation of the dry gas from the LPG and / or naphtha stream, and the dry gas containing hydrogen sulfide is included in the LPG and / or naphtha stream instead of being transferred via another sulfidizing agent line. It is also conceivable to add to the second reactor 170 in the hydrocarbon feed line.

The second reactor 170 may be a second FCC reactor. Although the second reactor 170 is depicted as a second FCC reactor, it will be appreciated that any suitable reactor such as a fixed bed or fluidized bed can be used. The second hydrocarbon feed passes to the second FCC reactor in the recycle feed line 166 via the feed distributor line 168 and / or the fluidized feed line 172 and the fluidized distributor feed line 174. Can be supplied. The second feed at least partially comprises C ₁₀ -hydrocarbons and preferably C ₄ to C ₁₀ olefins. Preferably, the second hydrocarbon feed mainly comprises hydrocarbons having 10 or fewer carbon atoms. Mainly means more than 50% by weight, preferably more than 80% by weight. The second feed consists of pyrolysis oil from the pyrolysis reactor, Fischer-Tropsch wax from the Fischer-Tropsch reactor, reformate from the catalytic reforming reactor, straight run naphtha from the crude oil column, and appropriate It can be any hydrocarbon with low sulfur compounds that decompose into hydrogen sulfide, such as animal fats and vegetable oils from a simple reactor or feedstock. The second feed is preferably part of the first cracked product produced in the first reactor 10 and is fractionated in the main column 100 of the product fractionation section 90 via the recycle line 166. Feed to the second reactor 170. In one embodiment, the second reactor communicates downstream with the product fractionator 90 and / or the first reactor 10 that communicates with the product fractionator 90 upstream. The second reactor 170 can include a second reactor riser 180. The second hydrocarbon feed is contacted with the catalyst sent to the second reactor 170 by a catalyst return tube 176 communicating upstream with the second reactor riser 180 to produce a high quality cracked product.

  The present invention contemplates adding a sulfiding agent to the second reactor 170 to suppress coking with the metal catalyst therein. The recycle dry gas sulfiding agent line 122 is a dedicated source of sulfiding agent that communicates upstream with the second reactor riser 180. In other words, dry gas and hydrogen sulfide do not convert to the desired hydrocarbon product and must be removed from the high quality cracked product exiting the second reactor 170, other than to prevent coking with the metal catalyst. Is not fed to the second reactor 170. The hydrocarbon feed and sulfiding agent can be introduced into the second reactor 170 in several ways as shown.

  In the first embodiment, the second hydrocarbon feed is in upstream communication with the second reactor riser 180 and in downstream communication with a feed distributor line 168 that is in communication with the recycle feed line 166 downstream. A feed distributor 178 can be injected into the second reactor riser 180. Feed distributor line 168 may capture some or all of the recycle feed stream from recycle feed line 166. The recycle feed line 166 communicates downstream with the top line 102 of the main column 100 that communicates with the first reactor 10 downstream. The feed rate in the feed distributor line 168 may be adjusted by a control valve. Feed distributor 178 may be located above fluidizing distributor 182 that communicates upstream with second reactor riser 180. The fluidizing distributor 182 supplies a flowing gas such as water vapor and / or light hydrocarbons to the second reactor riser 180 to fluidize the catalyst. In such an embodiment, the dry gas from the recirculating dry gas sulfidant line 122 passes downstream through a dedicated dry gas sulfidant line 184 that communicates with the recirculated dry gas sulfidant line 122 to a fluidized sulfidant line. 188 and the atomizing dry gas sulfiding agent line 186 in the fluidizing distributor feed line 174 can be bypassed and added independently to the fluidizing distributor 182 at the bottom of the second reactor riser 180. The dry gas thus serves as a fluidizing gas and as a sulfiding agent added to the second reactor riser 180 of the second reactor 170. Recycled dry gas sulfidant line 122, dedicated dry gas sulfidant line 184, and fluidized sulfidant line 188 are dedicated sulfidant sources that communicate upstream with fluidized distributor 182 and second reactor 170. . The dry gas with hydrogen sulfide in the recirculating dry gas sulfiding agent line 122, the dedicated drying gas sulfiding agent line 184, and the fluidized sulfiding agent line 188 is also an inert flow to the rest of the second reactor 170. It can be used as a chemical gas. In this embodiment, the control valves in the feed lines 158 and / or 162 and 168 and in the sulfidant lines 122, 184 and 188 are opened and the control valves in the feed line 172 and the sulfidant lines 124 and 186 are closed. It may be done.

  In a second embodiment, when the second feed is liquid, a dry gas containing hydrogen sulfide is added to the liquid second feed in feed distributor 178 to provide a second feed of liquid hydrocarbons. Atomize and passivate the metal in the second reactor. The recirculating dry gas sulfiding agent line 122 is a dedicated source of sulfiding agent that communicates upstream with the feed distributor 178 via the atomized drying gas sulfiding agent line 186. An atomized dry gas sulfide line 186 that communicates downstream with the dedicated dry gas sulfide line 184 supplies dry gas to the gas inlet of the feed distributor 178. According to this embodiment, the sulfiding agent can be added to the second reactor instead of or in addition to the method of adding the sulfiding agent in the first embodiment, ie, addition via fluidizing distributor 182. As a result, the control valve in line 186 is opened in addition to control valves that are opened and closed in other manners, allowing operation according to this second manner. Thus, at least the control valves in the sulfidant lines 122, 184, and 186 must be opened and operated in this manner.

  In a third embodiment, essentially all of the second hydrocarbon feed in the recycle feed line 166, i.e. at least 90 mol%, is in the gas phase. In general, the temperature of the second hydrocarbon feed may be between 120 ° C. and 600 ° C., preferably at least above the boiling point of the components, when entering the second reactor riser 180. In this embodiment, the second hydrocarbon feed can be fed directly to the fluidizer distributor 182 at the bottom of the second riser and the catalyst is fluidized and fed to the second reactor riser 180. In this embodiment, as shown in the figure, the sulfidant lines 122 and 124 and one or all of the control valves in the feed lines 158 and / or 162 and 172 are opened and the recycle dry gas sulfidant line 122 and Recycle feed line 166, fluidized supply of hydrogen sulfide-containing dry gas in dry gas mixture sulfidant line 124, and light naphtha in light naphtha line 158 and / or heavy naphtha in heavy naphtha line 162. Recirculate as a second feed in product line 172 and fluidizer distributor line 174 and distribute to risers by fluidizer distributor 182. In this embodiment, the feed line 168 and the valves in the sulfidant lines 184, 186, and 188 can normally be closed. The dry gas must contain sufficient hydrogen sulfide to passivate the metals that catalyze the coking in the second reactor riser 180 of the second reactor 170. A heat exchanger 190 may be necessary to volatilize the second feed that is recycled in the fluidized feed line 172. In this embodiment, fluidizer distributor feed line 174 serves as a feed line and fluidizer distributor 182 serves as a feed distributor.

  Hydrogen sulfide, or organic sulfur additives such as methyl sulfide, mercaptans and polysulfides, whether or not included in the dry gas, are suitable sulfiding agent additives to be added to the second reactor 170. The sulfiding agent additive can be added to the second feed in the feed lines 158, 162, 166, 168, 172, or 174 or somewhere upstream of the second reactor 170. For example, the sulfiding agent additive line 192 can add sulfiding agent directly to the fluidized feed line 172. The sulfiding agent may also be added directly to the second reactor riser 180, upstream of the fluidizing gas in the fluidizing distributor 182, or even to the catalyst entering the riser in the catalyst return line 176. When the SOx scavenging additive is added to the catalyst, the hydrogen sulfide adsorbed on the additive can be sent to the second reactor 170 via the pipe 216 and the catalyst return pipe 176, and the catalyst return pipe 176 and the pipe 216. One or both of these can be sulfiding agent lines. The sulfiding agent stream in the sulfiding agent line preferably has a hydrogen sulfide concentration or a compound concentration that can be converted to hydrogen sulfide in the reactor environment of at least 1000 wppm. The sulfur concentration for the fluid in the second reactor 170 should be maintained at least 20 wppm, preferably 50 wppm. In the riser reactor, the sulfur concentration should be maintained at least 20 wppm, preferably 50 wppm relative to the hydrocarbons and inert gases in the reactor. In one aspect, the sulfur concentration for the fluid in the second reactor should be maintained at 2000 wppm or less, preferably 500 wppm or less. In riser reactors, the sulfur concentration should be maintained at 2000 wppm or less, preferably 500 wppm relative to the hydrocarbons and inert gases in the reactor.

  Sulfidant lines 122, 124, 176, 184, 186, 188 and 192 are distinguished from feed lines 158 and 162. When the control valve in line 124 is closed, lines 166, 168, and 172 are also feed lines that are distinct from sulfiding agent lines 122, 184, 186, and 188. When the control valves in lines 124 and 172 are closed, fluidized feed line 172 no longer carries feed and fluidized distributor feed line 174 becomes the sulfiding agent line from which feed lines 158, 162. 166 and 168 are distinguished. The streams in the sulfiding agent line and the feed line may be mixed at a downstream location, but these streams are separated from each other at least at an upstream location. Thus, the sulfiding agent line supplies sulfiding agent separated from the second hydrocarbon feed upstream of the second reactor 170.

In general, the second reactor 170 can be operated under conditions to convert the hydrocarbon feed to smaller hydrocarbon products. C ₁₀ -olefins decompose into one or more light olefins such as ethylene and / or propylene. The second reaction vessel 194 communicates downstream with the second reactor riser 180 to receive the reformed product and catalyst from the second reactor riser. The mixture of gaseous reformed hydrocarbon product and catalyst travels up second reactor riser 180 and is received in second reaction vessel 194, where the reformed product is catalyst and gaseous hydrocarbons. Are separated. A pair of release arms 196 second isolates the gas from the catalyst from the top of the second reactor riser 180 via one or more outlets 198 (only one shown). Into the reaction vessel 194 tangentially and horizontally. The catalyst can fall into the dense catalyst bed 200 in the second reaction vessel 194. The reformed hydrocarbon product can then be separated from the catalyst and removed from the second reactor 170 through the reformed product line 206 and through an outlet 204 that communicates downstream with the second reactor 170. The reformed product in the reformed product line 206 can be directed to one or more cyclones 32 in the first reaction vessel 20 of the first reactor 10. These cyclones 32 may be dedicated to the reformed product from the second reactor 170 having a dedicated line (not shown) to the product fractionation section 90 or in particular the gas concentrating section 114, or the first reactor. The product from the riser 12 may be mixed and transferred to the product fractionation unit 90 in the line 88 together. Alternatively, the second reaction vessel 194 has one or more cyclones to further separate the gaseous reformed product from the catalyst and to the gas concentrating unit 114 of the product fractionating unit 90 via the reformed product line 206. You may migrate. The reformed product line 206 alternatively sends the reformed product to line 88 for transfer to the main column 100 of the product fractionator 90.

  In some embodiments, the second reactor 170 may include a mixture of first and second catalyst components as described above. In a preferred embodiment, the second reactor 170 can comprise less than 20% by weight, preferably 5% by weight of the first component, and at least 20% by weight of the second component. In another preferred embodiment, the second reactor 170 may contain only the second component as catalyst, preferably ZSM-5 zeolite.

  The separated catalyst may be regulated by control valve 210 and recycled from second reactor 194 via recycle catalyst tube 208 and returned to second riser reactor 180 to contact the second feed. Although not required, the catalyst is fed from the volatilization section 44 of the first FCC reactor to the second reactor 170 via the regeneration vessel 60 via pipe 214 and / or pipe 216, both regulated by control valves. . Both tubes 214 and 216 may be in communication with the recirculation catalyst tube 208 upstream. The catalyst return pipe 176 may be a part of the recirculation catalyst pipe 208. In one embodiment, the catalyst from the second reaction vessel 194 is sent by tube 202 to the first reactor, preferably the volatilization section 44, preferably after the volatilization process, through the spent catalyst conduit 54 to the regeneration vessel 60 for regeneration. Sent. The regenerated catalyst may be returned to the bottom of the second reactor riser 180 by way of the pipe 216 via the catalyst return pipe 176. In this embodiment, the catalysts in the first and second reactors 10 and 170 may be mixed and of a homogeneous composition in both reactors.

  In another embodiment, the second reactor 170 is isolated from the regeneration vessel 60 so that the regenerated catalyst is returned only to the first reactor 10 and the second reactor 170 sends the catalyst to the regeneration vessel 60. And no regenerated catalyst is received from it. In this embodiment, the second catalyst component is not exposed to repeated regeneration and therefore maintains high activity. Instead, the second catalyst component is added to the second reactor 170 and the catalyst in the second reaction vessel 194 is periodically or continuously adjusted to the volatilization portion 44 of the first reactor 10 by a control valve. Dispensed through pipe 202. The dispensed catalyst combines with the catalyst in the first reactor 10 where it provides additional catalytic activity. The fresh catalyst can replace the dispense catalyst to maintain activity in the second reactor 170.

  The second reactor riser 180 has a temperature of 425 ° C. to 705 ° C., preferably a temperature of 550 ° C. to 600 ° C., and a pressure of 40 kPa to 700 kPa, preferably a pressure of 40 kPa to 400 kPa, optimally a pressure of 200 kPa to 250 kPa. Can be operated under appropriate conditions such as Typically, the residence time of the second reactor riser 180 may be less than 5 seconds, preferably between 2 seconds and 3 seconds. Typical risers and / or operating conditions are disclosed, for example, in US 2008/0035527 A1 and US 7,261,807 B2.

  Without further elaboration, it is believed that one skilled in the art can fully utilize the present invention based on the foregoing description. The preferred embodiments described above should therefore be construed as merely illustrative and not in any way limiting on the remainder of the disclosure.

  In the above, all temperatures are given in degrees Celsius and parts and percentages are by weight unless otherwise specified.

  From the above description, those skilled in the art can easily confirm the essential characteristics of the present invention, and various modifications and changes can be made to the various uses and conditions without departing from the spirit and scope of the present invention. Can be adapted.

Claims (10)

Supplying a hydrocarbon feed to the reactor;
Sending the catalyst to the reactor;
Contacting the hydrocarbon feed with the catalyst;
Supplying a sulfiding agent separated from the feed upstream of the reactor;
Adding the sulfiding agent to the reactor;
A fluid catalytic cracking process comprising cracking the hydrocarbon feed into smaller hydrocarbon products and separating the hydrocarbon products from the catalyst.   The fluid catalytic cracking method of claim 1, further comprising distributing fluidized gas to the reactor and fluidizing the catalyst in the reactor.   3. The fluid catalytic cracking method of claim 2, further comprising adding the sulfiding agent to the hydrocarbon feed or the fluidizing gas or supplying a fluidizing gas containing the sulfiding agent.   A first hydrocarbon feed in a first fluid catalytic cracking reactor is cracked to provide a first cracked product and a portion of the first cracked product is fed to the reactor in the hydrocarbon feed. The fluid catalytic cracking method according to claim 1, further comprising supplying as a product.   The fluid catalytic cracking process of claim 4, further comprising fractionating the first cracked product and feeding the hydrocarbon feed to the reactor.   6. The fluid catalytic cracking method of claim 5, further comprising fractionating the first cracked product to provide a dry gas stream comprising hydrogen sulfide, wherein the dry gas stream further supplies the sulfiding agent.   The fluid catalytic cracking method according to claim 1, wherein the sulfurizing agent includes methyl sulfide, hydrogen sulfide, mercaptan, and polysulfide.   The fluid catalytic cracking method of claim 1, further comprising maintaining the sulfur concentration of the fluid in the reactor between 20 wppm and 2000 wppm.   The fluid catalytic cracking process of claim 1, wherein the hydrocarbon feed mainly comprises hydrocarbons having 10 or fewer carbon atoms. A riser for contacting the hydrocarbon feed with the catalyst to produce a product;
A catalyst tube in communication with the riser for sending catalyst to the riser;
A feed line for carrying hydrocarbon feeds;
A feed distributor in communication with the feed line and the riser and for sending the hydrocarbon feed to the riser;
A reaction vessel in communication with the riser for receiving product and catalyst from the riser;
A fluid catalytic cracker comprising: a sulfiding agent line that is distinct from the feed line and communicates with the riser.

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