Patent Grant
8894840
This invention generally relates to recovering product from catalytic reactors.
Fluid catalytic cracking (FCC) is a catalytic hydrocarbon conversion process accomplished by contacting heavier hydrocarbons in a fluidized reaction zone with a catalytic particulate material. The reaction in catalytic cracking, as opposed to hydrocracking, is carried out in the absence of substantial added hydrogen or the consumption of hydrogen. As the cracking reaction proceeds substantial amounts of highly carbonaceous material referred to as coke are deposited on the catalyst to provide coked or spent catalyst. Vaporous lighter products are separated from spent catalyst in a reactor vessel. Spent catalyst may be subjected to stripping over an inert gas such as steam to strip entrained hydrocarbonaceous gases from the spent catalyst. A high temperature regeneration with oxygen within a regeneration zone operation burns coke from the spent catalyst which may have been stripped. Various products may be produced from such a process, including a naphtha product and/or a light product such as propylene and/or ethylene.
In such processes, a single reactor or a dual reactor can be utilized. Although additional capital costs may be incurred by using a dual reactor apparatus, one of the reactors can be operated to tailor conditions for maximizing products, such as light olefins including propylene and/or ethylene. It can often be advantageous to maximize yield of a product in one of the reactors. Additionally, there may be a desire to maximize the production of a product from one reactor that can be recycled back to the other reactor to produce a desired product, such as propylene.
Normally if two reactors are used, a single product recovery system is utilized for product separation. Separate product recovery systems have also been proposed. Maximizing synergies between two reactor systems is greatly desired.
As used herein, the following terms have the corresponding definitions.
The term “communication” means that material flow is operatively permitted between enumerated components.
The term “downstream communication” means that at least a portion of material flowing to the subject in downstream communication may operatively flow from the object with which it communicates.
The term “upstream communication” means that at least a portion of the material flowing from the subject in upstream communication may operatively flow to the object with which it communicates.
The term “direct communication” means that flow from the upstream component enters the downstream component without undergoing a compositional change due to physical fractionation or chemical conversion.
The term “column” means a distillation column or columns for separating one or more components of different volatilities which may have a reboiler on its bottom and a condenser on its overhead. Unless otherwise indicated, each column includes a condenser on an overhead of the column to condense and reflux a portion of an overhead stream back to the top of the column and a reboiler at a bottom of the column to vaporize and send a portion of a bottoms stream back to the bottom of the column. Feeds to the columns may be preheated. The top pressure is the pressure of the overhead vapor at the outlet of the column. The bottom temperature is the liquid bottom outlet temperature.
The term “Cx−” wherein “x” is an integer means a hydrocarbon stream with hydrocarbons having x and/or less carbon atoms and preferably x and less carbon atoms.
The term “Cx+” wherein “x” is an integer means a hydrocarbon stream with hydrocarbons having x and/or more carbon atoms and preferably x and more carbon atoms.
The term “predominant” means a majority, suitably at least 80 wt-% and preferably at least 90 wt-%.
In a process embodiment, the subject invention involves a catalytic cracking process comprising feeding a first hydrocarbon feed to a wash column and feeding the hydrocarbon feed from the wash column to a first reactor. Catalyst is delivered to the first reactor and contacted with the first hydrocarbon feed to provide first cracked products. A portion of the first cracked products are fed as a second hydrocarbon feed to a second reactor. Catalyst is delivered to the second reactor and contacted with the second hydrocarbon feed to provide second cracked products. The second cracked products are fed to the wash column. In another process embodiment, the subject invention involves vaporizing a portion of the first cracked products to provide the second hydrocarbon feed.
In another process embodiment, the subject invention involves a fluid catalytic cracking process comprising a first hydrocarbon feed in route to a first fluid catalytic cracking reactor that is contacted with a second hydrocarbon product from a second fluid catalytic cracking reactor.
In an apparatus embodiment, the subject invention involves a catalytic cracking apparatus comprising a first catalytic reactor in communication with a wash column. A second catalytic reactor is in communication with the first catalytic reactor, and the wash column is in communication with the second reactor. In an alternative embodiment, a main column is in communication with the first catalytic reactor and a second catalytic reactor is in communication with the main column. In a further alternative embodiment, a debutanizer column is in communication with the first catalytic reactor and a naphtha splitter column is in communication with the debutanizer column. The second catalytic reactor is in communication with the naphtha splitter column.
Commercially there is a demand for FCC technology capable of producing high propylene yields from conventional feedstocks. While it is possible to affect the propylene yield in a conventional FCC unit by adjusting the process conditions and the catalyst composition the extent of propylene production is equilibrium-limited. One means of increasing the propylene yield is to decrease the reactor pressure to decrease olefin partial pressure. However, reducing the reactor pressure leads to a large increase in capital cost and an even larger increase in the utility costs. An alternative solution is feeding light naphtha to the primary reactor riser or to a second reactor riser from a conventional separation section having a main column and gas recovery unit. Both of these options result in an increase in capital costs, but the process economics are much more favorable than simply reducing the reactor pressure. If one recycles light naphtha to a conventional reactor riser to increase propylene yield, the capital costs increase slightly with essentially no increase in utility costs. Propylene yield can be further increased if the recycle is instead fed to a second riser with a common separation system, but obviously the capital and the utility costs increase substantially but less than by simply reducing the reactor pressure.
We have found that propylene yield can be increased to a still greater extent more economically by directing the effluent from the second riser reactor to a segregated separation section. Exploiting a dual riser-dual separation section flow scheme it was possible to increase the propylene yield but with surprisingly significantly less capital and utility costs over that provided by an equivalent dual riser with common separation system.
The present invention is an apparatus and process that may be described with reference to four components shown in
A conventional FCC feedstock and higher boiling hydrocarbon feedstock are a suitable first feed 8 to the first FCC reactor. The most common of such conventional feedstocks is a “vacuum gas oil” (VGO), which is typically a hydrocarbon material having a boiling range of from 343° to 552° C. (650° to 1025° F.) prepared by vacuum fractionation of atmospheric residue. Such a fraction is generally low in coke precursors and heavy metal contamination which can serve to contaminate catalyst. Heavy hydrocarbon feedstocks to which this invention may be applied include heavy bottoms from crude oil, heavy bitumen crude oil, shale oil, tar sand extract, deasphalted residue, products from coal liquefaction, atmospheric and vacuum reduced crudes. Heavy feedstocks for this invention also include mixtures of the above hydrocarbons and the foregoing list is not comprehensive. Moreover, additional amounts of feed may also be introduced downstream of the initial feed point. The first feed in line 8 may be preheated in wash column 30 which will be further discussed hereafter.
The first reactor 10 which may be a catalytic or an FCC reactor that includes a first reactor riser 12 and a first reactor vessel 20. A regenerator catalyst pipe 14 is in upstream communication with the first reactor riser 12. The regenerator catalyst pipe 14 delivers regenerated catalyst from the regenerator vessel 60 at a rate regulated by a control valve to the reactor riser 12 through a regenerated catalyst inlet. A fluidization medium such as steam from a distributor 18 urges a stream of regenerated catalyst upwardly through the first reactor riser 12. At least one feed distributor 22 in upstream communication with the first reactor riser 12 injects the first hydrocarbon feed 8, preferably with an inert atomizing gas such as steam, across the flowing stream of catalyst particles to distribute hydrocarbon feed to the first reactor riser 12. Upon contacting the hydrocarbon feed with catalyst in the first reactor riser 12 the heavier hydrocarbon feed cracks to produce lighter gaseous first cracked products while conversion coke and contaminant coke precursors are deposited on the catalyst particles to produce spent catalyst.
The first reactor vessel 20 is in downstream communication with the first reactor riser 12. The resulting mixture of gaseous product hydrocarbons and spent catalyst continues upwardly through the first reactor riser 12 and are received in the first reactor vessel 20 in which the spent catalyst and gaseous product are separated. A pair of disengaging arms 24 may tangentially and horizontally discharge the mixture of gas and catalyst from a top of the first reactor riser 12 through one or more outlet ports 26 (only one is shown) into a disengaging vessel 28 that effects partial separation of gases from the catalyst. A transport conduit 30 carries the hydrocarbon vapors, including stripped hydrocarbons, stripping media and entrained catalyst to one or more cyclones 32 in the first reactor vessel 20 which separates spent catalyst from the hydrocarbon gaseous product stream. The disengaging vessel 28 is partially disposed in the first reactor vessel 20 and can be considered part of the first reactor vessel 20. Gas conduits deliver separated hydrocarbon gaseous streams from the cyclones 32 to a collection plenum 36 in the first reactor vessel 20 for passage to a product line 88 via an outlet nozzle and eventually into the product fractionation section 90 for product recovery. Diplegs discharge catalyst from the cyclones 32 into a lower bed in the first reactor vessel 20. The catalyst with adsorbed or entrained hydrocarbons may eventually pass from the lower bed into an optional stripping section 44 across ports defined in a wall of the disengaging vessel 28. Catalyst separated in the disengaging vessel 28 may pass directly into the optional stripping section 44 via a bed. A fluidizing distributor 50 delivers inert fluidizing gas, typically steam, to the stripping section 44. The stripping section 44 contains baffles 52 or other equipment to promote contacting between a stripping gas and the catalyst. The stripped spent catalyst leaves the stripping section 44 of the disengaging vessel 28 of the first reactor vessel 20 with a lower concentration of entrained or adsorbed hydrocarbons than it had when it entered or if it had not been subjected to stripping. A first portion of the spent catalyst, preferably stripped, leaves the disengaging vessel 28 of the first reactor vessel 20 through a spent catalyst conduit 54 and passes into the regenerator vessel 60 at a rate regulated by a slide valve. The regenerator 60 is in downstream communication with the first reactor 10. A second portion of the spent catalyst is recirculated in recycle conduit 56 from the disengaging vessel 28 back to a base of the riser 12 at a rate regulated by a slide valve to recontact the feed without undergoing regeneration.
The first reactor riser 12 can operate at any suitable temperature, and typically operates at a temperature of about 150° to about 580° C., preferably about 520° to about 580° C. at the riser outlet 24. In one exemplary embodiment, a higher riser temperature may be desired, such as no less than about 565° C. at the riser outlet port 24 and a pressure of from about 69 to about 517 kPa (gauge) (10 to 75 psig) but typically less than about 275 kPa (gauge) (40 psig). The catalyst-to-oil ratio, based on the weight of catalyst and feed hydrocarbons entering the bottom of the riser, may range up to 30:1 but is typically between about 4:1 and about 10:1 and may range between 7:1 and 25:1. Hydrogen is not normally added to the riser. Steam may be passed into the first reactor riser 12 and first reactor vessel 20 equivalent to about 2-35 wt-% of feed. Typically, however, the steam rate may be between about 2 and about 7 wt-% for maximum gasoline production and about 10 to about 15 wt-% for maximum light olefin production. The average residence time of catalyst in the riser may be less than about 5 seconds.
The catalyst in the first reactor 10 can be a single catalyst or a mixture of different catalysts. Usually, the catalyst includes two components or catalysts, namely a first component or catalyst, and a second component or catalyst. Such a catalyst mixture is disclosed in, e.g., U.S. Pat. No. 7,312,370 B2. Generally, the first component may include any of the well-known catalysts that are used in the art of FCC, such as an active amorphous clay-type catalyst and/or a high activity, crystalline molecular sieve. Zeolites may be used as molecular sieves in FCC processes. Preferably, the first component includes a large pore zeolite, such as a Y-type zeolite, an active alumina material, a binder material, including either silica or alumina, and an inert filler such as kaolin.
Typically, the zeolitic molecular sieves appropriate for the first component have a large average pore size. Usually, molecular sieves with a large pore size have pores with openings of greater than about 0.7 nm in effective diameter defined by greater than about 10, and typically about 12, member rings. Pore Size Indices of large pores can be above about 31. Suitable large pore zeolite components may include synthetic zeolites such as X and Y zeolites, mordenite and faujasite. A portion of the first component, such as the zeolite, can have any suitable amount of a rare earth metal or rare earth metal oxide.
The second component may include a medium or smaller pore zeolite catalyst, such as a MFI zeolite, as exemplified by at least one of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48, and other similar materials. Other suitable medium or smaller pore zeolites include ferrierite, and erionite. Preferably, the second component has the medium or smaller pore zeolite dispersed on a matrix including a binder material such as silica or alumina and an inert filler material such as kaolin. The second component may also include some other active material such as Beta zeolite. These compositions may have a crystalline zeolite content of about 10 to about 50 wt-% or more, and a matrix material content of about 50 to about 90 wt-%. Components containing about 40 wt-% crystalline zeolite material are preferred, and those with greater crystalline zeolite content may be used. Generally, medium and smaller pore zeolites are characterized by having an effective pore opening diameter of less than or equal to about 0.7 nm, rings of about 10 or fewer members, and a Pore Size Index of less than about 31. Preferably, the second catalyst component is an MFI zeolite having a silicon-to-aluminum ratio greater than about 15, preferably greater than about 75. In one exemplary embodiment, the silicon-to-aluminum ratio can be about 15:1 to about 35:1.
The total catalyst mixture in the first reactor 10 may contain about 1 to about 25 wt-% of the second component, including a medium to small pore crystalline zeolite with greater than or equal to about 7 wt-% of the second component being preferred. When the second component contains about 40 wt-% crystalline zeolite with the balance being a binder material, an inert filler, such as kaolin, and optionally an active alumina component, the catalyst mixture may contain about 0.4 to about 10 wt-% of the medium to small pore crystalline zeolite with a preferred content of at least about 2.8 wt-%. The first component may comprise the balance of the catalyst composition. In some preferred embodiments, the relative proportions of the first and second components in the mixture may not substantially vary throughout the first reactor 10. The high concentration of the medium or smaller pore zeolite as the second component of the catalyst mixture can improve selectivity to light olefins. In one exemplary embodiment, the second component can be a ZSM-5 zeolite and the catalyst mixture can include about 0.4 to about 10 wt-% ZSM-5 zeolite excluding any other components, such as binder and/or filler.
The regenerator vessel 60 is in downstream communication with the first reactor vessel 20. In the regenerator vessel 60, coke is combusted from the portion of spent catalyst delivered to the regenerator vessel 60 by contact with an oxygen-containing gas such as air to provide regenerated catalyst. The regenerator vessel 60 may be a combustor type of regenerator as shown in
The mixture of catalyst and combustion gas in the lower chamber 62 ascend through a frustoconical transition section 66 to the transport, riser section 68 of the lower chamber 62. The riser section 68 defines a tube which is preferably cylindrical and extends preferably upwardly from the lower chamber 62. The mixture of catalyst and gas travels at a higher superficial gas velocity than in the lower chamber 62. The increased gas velocity is due to the reduced cross-sectional area of the riser section 68 relative to the cross-sectional area of the lower chamber 62 below the transition section 66. Hence, the superficial gas velocity may usually exceed about 2.2 m/s (7 ft/s). The riser section 68 may have a catalyst density of less than about 80 kg/m3 (5 lb/ft3).
The regenerator vessel 60 also may include an upper or second chamber 70. The mixture of catalyst particles and flue gas is discharged from an upper portion of the riser section 68 into the upper chamber 70. Substantially completely regenerated catalyst may exit the top of the transport, riser section 68, but arrangements in which partially regenerated catalyst exits from the lower chamber 62 are also contemplated. Discharge is effected through a disengaging device 72 that separates a majority of the regenerated catalyst from the flue gas. In an embodiment, catalyst and gas flowing up the riser section 68 impact a top elliptical cap of a disengaging device 72 and reverse flow. The catalyst and gas then exit through downwardly directed discharge outlets of the disengaging device 72. The sudden loss of momentum and downward flow reversal cause a majority of the heavier catalyst to fall to the dense catalyst bed and the lighter flue gas and a minor portion of the catalyst still entrained therein to ascend upwardly in the upper chamber 70. Cyclones 75, 76 further separate catalyst from ascending gas and deposits catalyst through diplegs into dense catalyst bed. Flue gas exits the cyclones 75, 76 through a gas conduit and collects in a plenum 82 for passage to an outlet nozzle of regenerator vessel 60 and perhaps into a flue gas or power recovery system (not shown). Catalyst densities in the dense catalyst bed are typically kept within a range of from about 640 to about 960 kg/m3 (40 to 60 lb/ft3). A fluidizing conduit delivers fluidizing gas, typically air, to the dense catalyst bed 74 through a fluidizing distributor. In an embodiment, to accelerate combustion of the coke in the lower chamber 62, hot regenerated catalyst from a dense catalyst bed in the upper chamber 70 may be recirculated into the lower chamber 62 via recycle conduit (not shown).
The regenerator vessel 60 may typically require 14 kg of air per kg of coke removed to obtain complete regeneration. When more catalyst is regenerated, greater amounts of feed may be processed in the first reactor 10. The regenerator vessel 60 typically has a temperature of about 594° to about 704° C. (1100° to 1300° F.) in the lower chamber 62 and about 649° to about 760° C. (1200° to 1400° F.) in the upper chamber 70. The regenerated catalyst pipe 14 is in downstream communication with the regenerator vessel 60. Regenerated catalyst from dense catalyst bed is transported through regenerated catalyst pipe 14 from the regenerator vessel 60 back to the first reactor riser 12 through the control valve where it again contacts the first feed in line 8 as the FCC process continues.
The first cracked products in the line 88 from the first reactor 10, relatively free of catalyst particles and including the stripping fluid, exit the first reactor vessel 20 through the outlet nozzle. The first cracked products stream in the line 88 may be subjected to additional treatment to remove fine catalyst particles or to further prepare the stream prior to fractionation. The line 88 transfers the first cracked products stream to the product fractionation section 90 that in an embodiment may include a main fractionation column 100 and a gas recovery section 120.
The main column 100 is a fractionation column with trays and/or packing positioned along its height for vapor and liquid to contact and reach equilibrium proportions at tray conditions and a series of pump-arounds to cool the contents of the main column. The main fractionation column is in downstream communication with the first reactor 10 and can be operated with an top pressure of about 35 to about 172 kPa (gauge) (5 to 25 psig) and a bottom temperature of about 343 to about 399° C. (650 to 750° F.). In the product recovery section 90, the gaseous FCC product in line 88 is directed to a lower section of an FCC main fractionation column 100. A variety of products are withdrawn from the main column 100. In this case, the main column 100 recovers an overhead stream of light products comprising unstabilized naphtha and lighter gases in an overhead line 94. The overhead stream in overhead line 94 is condensed in a condenser and perhaps cooled in a cooler both represented by 96 before it enters a receiver 98 in downstream communication with the first reactor 10. A line 102 withdraws a light off-gas stream of liquefied petroleum gas (LPG) and dry gas from the receiver 98. An aqueous stream is removed from a boot in the receiver 98. A bottoms liquid stream of light unstabilized naphtha leaves the receiver 98 via a line 104. A first portion of the bottoms liquid stream is directed back to an upper portion of the main column and a second portion in line 106 may be directed to the gas recovery section 120. Both lines 102 and 106 may be fed to the gas recovery section 120.
Several other fractions may be separated and taken from the main column including an optional heavy naphtha stream in line 108, a light cycle oil (LCO) in line 110, a heavy cycle oil (HCO) stream in line 112, and heavy slurry oil from the bottom in line 114. Portions of any or all of lines 108-114 may be recovered while remaining portions may be cooled and pumped back around to the main column 100 to cool the main column typically at a higher entry location. The light unstabilized naphtha fraction preferably has an initial boiling point (IBP) below in the C5 range; i.e., below about 35° C. (95° F.), and an end point (EP) at a temperature greater than or equal to about 127° C. (260° F.). The boiling points for these fractions are determined using the procedure known as ASTM D86-82. The optional heavy naphtha fraction has an IBP at or above about 127° C. (260° F.) and an EP at a temperature above about 200° C. (392° F.), preferably between about 204° and about 221° C. (400° and 430° F.), particularly at about 216° C. (420° F.). The LCO stream has an IBP at or above 177° C. (350° F.) if no heavy naphtha cut is taken or at about the EP temperature of the heavy naphtha if a heavy naphtha cut is taken and an EP in a range of about 260° to about 371° C. (500° to 700° F.) and preferably about 343° C. (650° F.). The HCO stream has an IBP of the EP temperature of the LCO stream and an EP in a range of about 371° to about 427° C. (700° to 800° F.), and preferably about 399° C. (750° F.). The heavy slurry oil stream has an IBP of the EP temperature of the HCO stream and includes everything boiling at a higher temperature.
The gas recovery section 120 is shown to be an absorption based system, but any vapor recovery system may be used including a cold box system. To obtain sufficient separation of light gas components the gaseous stream in line 102 is compressed in a compressor 122 also known as a wet gas compressor. Any number of compressor stages may be used, but typically a dual stage compression is utilized. In a dual stage compression, compressed fluid from compressor 122 is cooled and enters an interstage compressor receiver 124. Liquid in line 126 from a bottom of the compressor receiver 124 joins the unstabilized naphtha in line 106 and together flow in line 136 to a top section of a primary absorber column 140. Gas in line 128 from a top of the compressor receiver 124 enters a second compressor 130, also known as a wet gas compressor. Compressed effluent from the second compressor 130 in line 131 is joined by streams in lines 138 and 142 and are cooled and fed to a second compressor receiver 132. Compressed gas from a top of the second compressor receiver 132 travels in line 134 to enter a lower section of a primary absorber column 140. A liquid stream from a bottom of the second compressor receiver 132 travels in line 144 to a stripper column 146. The first compression stage compress gaseous fluids to a pressure of about 345 to about 1034 kPa (gauge) (50 to 150 psig) and preferably about 482 to about 690 kPa (gauge) (70 to 100 psig). The second compression stage compresses gaseous fluids to a pressure of about 1241 to about 2068 kPa (gauge) (180 to 300 psig).
The gaseous hydrocarbon stream in line 134 is routed to the primary absorber column 140 in which it is contacted with unstabilized naphtha from the main column receiver 98 in line 106 to effect a separation between C3+ and C2− hydrocarbons by absorption of the heavier hydrocarbons into the naphtha stream by counter-current contact. The primary absorber column 140 utilizes no condenser or reboiler but may have one or more pump-arounds (not shown) to cool the materials in the column. The primary absorber column may be operated at a top pressure of about 1034 to about 2068 kPa (gauge) (150 to 300 psig) and a bottom temperature of about 27 to about 66° C. (80 to 150° F.). A predominantly liquid C3+ stream with a relatively small amount of C2− material in solution in line 142 from the bottom of the primary absorber column 140 is returned to line 131 upstream of the condenser to be cooled and returned to the second compressor receiver 132.
An off-gas stream in line 148 from a top of the primary absorber column 140 is directed to a secondary or sponge absorber column 150. A circulating stream of LCO in line 152 diverted from line 110 to the secondary absorber column 150 absorbs most of the remaining C5+ and some C3-C4 material in the off-gas stream in line 148. LCO from a bottom of the secondary absorber column in line 156 richer in C3+ material is returned in line 156 to the main column 100 via the pump-around for line 110. The secondary absorber column 150 may be operated at a top pressure just below the pressure of the primary absorber column 140 of about 965 to about 2000 kPa (gauge) (140 to 290 psig) and a bottom temperature of about 38 to about 66° C. (100 to 150° F.). The overhead of the secondary absorber column 150 comprising dry gas of predominantly C2− hydrocarbons with hydrogen sulfide, amines and hydrogen is removed in line 158 and may be subjected to further separation to recover ethylene and hydrogen.
Liquid from a bottom of the second compressor receiver 132 in line 144 is sent to the stripper column 146. Most of the C2− is removed in an overhead of the stripper column 146 and returned to line 131 via overhead line 138 without first undergoing condensation. The condenser on line 131 will partially condense the overhead stream in line 138 with the gas compressor discharge in line 131 and with the bottoms stream 142 from the primary absorber column 140 will together undergo vapor-liquid separation in second compressor receiver 132. The stripper may be run at a pressure above the compressor 130 discharge at about 1379 to about 2206 kPa (gauge) (200 to 320 psig) and a temperature of about 38 to about 149° C. (100 to 300° F.).
A liquid bottoms stream comprising C3+ material from the stripper column 146 is sent to a debutanizer column 160 via line 162. The debutanizer column 160 is in downstream communication with the first reactor 10 and the primary absorber column 140 and fractionates a portion of first cracked products from the first reactor 10 to provide a C4− overhead stream and C5+ bottoms stream. The debutanizer column may be operated at a top pressure of about 1034 to about 1724 kPa (gauge) (150 to 250 psig) and a bottom temperature of about 149 to about 204° C. (300 to 400° F.). The pressure should be maintained as low as possible to maintain reboiler temperature as low as possible while still allowing complete condensation with typical cooling utilities without the need for refrigeration. The overhead stream in line 164 from the debutanizer comprises C3-C4 olefinic product which can be sent to an LPG splitter column 170 which is in downstream communication with an overhead of the debutanizer column 160. The bottoms stream in line 166 may be split between line 168 for delivering debutanized naphtha to the primary absorber column 140 to assist in the absorption of C3+ materials and line 172 for delivery to the naphtha splitter column 180.
In the LPG splitter column 170, C3 materials may be forwarded from the overhead in a line 174 to a C3 splitter to recover propylene product. C4 materials from the bottoms in line 176 may be recovered for blending in a gasoline pool as product or further processed. The LPG splitter 170 may be operated with a top pressure of about 69 to about 207 kPa (gauge) (10 to 30 psig) and a bottom temperature of about 38 to about 121° C. (100 to 250° F.).
In an embodiment, the naphtha splitter column 180 may be in downstream communication with a bottom of the debutanizer column 160. In the naphtha splitter column 180, a light naphtha stream, typically a C5-C6 or a C5-C7 stream is recovered from the overhead in line 182 for gasoline blending or further processed. Heavy naphtha from the bottom in line 184 typically comprising C7+ materials may be recovered or further processed. The naphtha splitter column may be operated with a top pressure of about 69 to about 448 kPa (gauge) (10 to 65 psig) and a bottom temperature of about 121 to about 232° C. (250 to 450° F.). The pressure of this column may be adjusted into a different range to facilitate heat integration and minimize utility consumption.
In an embodiment, C4 material in line 176 is vaporized in an evaporator 177 to provide a vaporized C4 stream 178. The light naphtha in line 182 may be vaporized in an evaporator 188 to provide a vaporized light naphtha stream in line 186. The vaporized streams in lines 178 and 186 may be mixed to provide a mixed vaporized light naphtha stream in line 190. The streams in lines 176 and 182 may be vaporized in the same evaporator. The vaporized stream in line 190 may be delivered as a second hydrocarbon feed to a second catalytic reactor 200 which is in downstream communication with an overhead of the main fractionation column 100, a bottoms of the primary absorber 140, a bottoms of the LPG splitter and an overhead of the naphtha splitter 180. In an embodiment, the mixed vaporized light naphtha stream in line 190 may be superheated in a heat exchanger before it is fed to the second catalytic reactor 200 in line 190.
The second catalytic reactor 200 may be a second FCC reactor. Although the second reactor 200 is depicted as a second FCC reactor, it should be understood that any suitable catalytic reactor can be utilized, such as a fixed bed or a fluidized bed reactor. The second hydrocarbon feed may be fed to the secondary FCC reactor 200 in recycle feed line 190 via feed distributor 202. The second feed can at least partially be comprised of C10-hydrocarbons, preferably comprising C4 to C7 olefins. The second hydrocarbon feed predominantly comprises hydrocarbons with 10 or fewer carbon atoms and preferably between 4 and 7 carbon atoms. The second hydrocarbon feed is preferably a portion of the first cracked products produced in the first reactor 10, fractionated in the main column 100 of the product recovery section 90 and provided to the second reactor 200. In an embodiment, the second reactor is in downstream communication with the product fractionation section 90 and/or the first reactor 10 which is in upstream communication with the product fractionation section 90.
The second reactor 200 may include a second riser reactor 212. The second hydrocarbon feed is contacted with catalyst delivered to the second reactor 200 by a catalyst return pipe 204 in upstream communication with the second reactor riser 212 to produce cracked upgraded products. The catalyst may be fluidized by inert gas such as steam from distributor 206. Generally, the second reactor 200 may operate under conditions to convert the light naphtha feed to smaller hydrocarbon products. C4-C7 olefins crack into one or more light olefins, such as ethylene and/or propylene. A second reactor vessel 220 is in downstream communication with the second reactor riser 212 for receiving upgraded products and catalyst from the second reactor riser. The mixture of gaseous, upgraded product hydrocarbons and catalyst continues upwardly through the second reactor riser 212 and is received in the second reactor vessel 220 in which the catalyst and gaseous hydrocarbon, upgraded products are separated. A pair of disengaging arms 208 may tangentially and horizontally discharge the mixture of gas and catalyst from a top of the second reactor riser 212 through one or more outlet ports 210 (only one is shown) into the second reactor vessel 220 that effects partial separation of gases from the catalyst. The catalyst can drop to a dense catalyst bed within the second reactor vessel 220. Cyclones 224 in the second reactor vessel 220 may further separate catalyst from second cracked products. Afterwards, the second cracked hydrocarbon products can be removed from the second reactor 200 through an outlet 226 in downstream communication with the second reactor riser 212 through a second cracked products line 228. Separated catalyst may be recycled via a recycle catalyst pipe 204 from the second reactor vessel 220 regulated by a control valve back to the second reactor riser 212 to be contacted with the second hydrocarbon feed.
In some embodiments, the second reactor 200 can contain a mixture of the first and second catalyst components as described above for the first reactor. In one preferred embodiment, the second reactor 200 can contain less than about 20 wt-%, preferably less than about 5 wt-% of the first component and at least 20 wt-% of the second component. In another preferred embodiment, the second reactor 200 can contain only the second component, preferably a ZSM-5 zeolite, as the catalyst.
The second reactor 200 is in downstream communication with the regenerator vessel 60 and receives regenerated catalyst therefrom in line 214. In an embodiment, the first catalytic reactor 10 and the second catalytic reactor 200 both share the same regenerator vessel 60. The same catalyst composition may be used in both reactors 10, 200. However, if a higher proportion of small to medium pore zeolite is desired in the second reactor 200, replacement catalyst added to the second reactor 200 may comprise a high proportion of the second catalyst component. Because the second catalyst component does not lose activity as quickly as the first catalyst component, less of the catalyst inventory need be forwarded to the catalyst regenerator 60 but more catalyst inventory may be recycled to the riser 212 in return conduit 204 without regeneration to maintain the high level of the second catalyst component in the second reactor 200. Line 216 carries spent catalyst from the second reactor vessel 220 with a control valve for restricting the flow rate of catalyst from the second reactor 200 to the regenerator vessel 60. The catalyst regenerator is in downstream communication with the second reactor 200 via line 216. A means for segregating catalyst compositions from respective reactors in the regenerator 60 may also be implemented.
The second reactor riser 212 can operate in any suitable condition, such as a temperature of about 425° to about 705° C., preferably a temperature of about 550° to about 600° C., and a pressure of about 40 to about 700 kPa (gauge), preferably a pressure of about 40 to about 400 kPa (gauge), and optimally a pressure of about 200 to about 250 kPa (gauge). Typically, the residence time of the second reactor riser 212 can be less than about 5 seconds and preferably is between about 2 and about 3 seconds. Exemplary risers and operating conditions are disclosed in, e.g., US 2008/0035527 A1 and U.S. Pat. No. 7,261,807 B2.
One unique feature of the disclosed apparatus and process is the separate recovery processing of the effluent from the first and second reactors 10, 200. We have surprisingly found that the separate processing of the products of the first and second reactors not only results in a higher propylene yield, but also reduces the capital cost and utility cost when compared to a two riser reactor system with co-mingled reactor effluent in the same product recovery section. The separate product recovery sections result in less dilution of the second hydrocarbon feed with paraffins hence providing a feed richer in olefins. With less dilution of the second hydrocarbon feed with paraffins, the second hydrocarbon feed rate is lower to the second catalytic reactor 200 and recirculation of C4+ material in the gas recovery section is limited to the primary absorber lean oil in line 142.
The second products from the second reactor 200 in line 228 are directed to a second product recovery section 230. Another aspect of the apparatus and process is heat recovery from the second products in line 228 from the second reactor 200 in the wash column 30. The wash column 30 is in downstream communication with said second reactor 200 and in upstream communication with the first reactor 10.
The cooled second products exit from the wash column 30 in overhead line 232, are partially condensed and enter into a wash column receiver 234. A liquid potion of the second products are returned to an upper section of the wash column 30 and a vapor portion of the second products is directed to a third compressor 240 which is in downstream communication with the wash column 30 and the second reactor 200. The third compressor 240 may be only a single stage or followed by one compressor 244 or more. In the case of two stages, as shown in
The depropanizer column 250 is in downstream communication with the second reactor 200. In the depropanizer column 250, fractionation of the compressed second product stream occurs to provide a C3− overhead stream and a C4+ bottoms stream. To avoid unnecessarily duplicating equipment the depropanizer column overhead stream carrying a light portion of the second products from the second reactor is processed in the gas recovery section 120. An overhead line 154 carries an overhead stream of C3− materials to join line 134 and enter a lower section of the primary absorber column 140 in the gas recovery section 120. The heavier C3 hydrocarbons from the C3− overhead stream are absorbed into the naphtha stream in the primary absorber column 140. This allows common recovery of propylene and dry gas and eliminates the need for duplicate absorption systems or alternate light olefin separation schemes. The depropanizer column 250 operates with a top pressure of about 1379 to about 2413 kPa (gauge) (200 to 350 psig) and a bottom temperature of about 121 to about 177° C. (250 to 350° F.). A depropanized bottom stream in line 254 exits the bottom of the depropanizer column 250 and enters a second debutanizer column 260 through line 254.
The second debutanizer column 260 is in downstream communication with the second reactor 200. In the second debutanizer column 260, fractionation of a depropanized portion of the compressed second product stream occurs to provide a C4− overhead stream and a C5+ light naphtha bottoms stream. An overhead line 262 carries an overhead stream of predominantly C4 hydrocarbons to undergo further processing or recovery. The second debutanizer column 260 operates with a top pressure of about 276 to about 690 kPa (gauge) (40 to 100 psig) and a bottom temperature of about 93 to about 149° C. (200 to 300° F.). A debutanized bottoms light naphtha stream in line 264 exits the bottom of the second debutanizer column 260 which may be further processed or sent to the gasoline pool.
The apparatus and process has the flexibility of providing recycle material from the second product recovery section 230 with no impact on the gas recovery section 120. If a small recycle flow rate is required to achieve the target propylene yield then, in an embodiment, vaporized C4 hydrocarbons from the overhead line 262 may be diverted in line 266 prior to condensation and carried to line 190 for recycle to the second reactor In this embodiment, the second reactor 200 may be in downstream communication with an overhead of the second debutanizer column 260. C4 hydrocarbon recycle from the debutanizer column 260 may be practiced with any other embodiment herein.
In an alternative embodiment, the first debutanizer column is replaced with a first depropanizer column and the LPG splitter column is eliminated to result in a more energy efficient and lower capital cost design.
The gas recovery system 120′ is different in
In an embodiment, a naphtha splitter column 180 may be in downstream communication with a bottom of the depropanizer column 160′. In the naphtha splitter column 180, a light naphtha stream, typically a C4-C6 stream is recovered from the overhead in line 182′ for gasoline blending or further processing. The overhead stream may be taken before condensation to assure a vapor naphtha stream is taken as the second hydrocarbon feed in line 190′. Heavy naphtha from the bottoms in line 184 typically comprising C7+ materials may be recovered or further processed.
The second product recovery section 230′ is also different in
The embodiment of
The gas recovery system 120″ is different in
In the LPG splitter 170, C3 materials may be forwarded from the overhead in a line 174 to a C3 splitter to recover propylene product. C4 materials from the bottoms in line 176″ may be recovered for blending in a gasoline pool as product or further processed. In this embodiment the bottoms stream in line 176″ is reboiled and split with a portion going back to the column and the other portion of vaporized C4 hydrocarbons for recycle in line 178. The vaporized stream in line 178 is mixed with vaporous heart cut naphtha in line 183 to form a light naphtha stream in line 190″. Alternatively, the bottoms stream in line 176″ may be reboiled in a typical reboiler with the recycle in line 178 being vaporized in a separate evaporator heat exchanger (not shown).
The second product recovery section 230″ is different in
Preferably, a side cut from the bottom of the depropanizer column 250″pulls a vapor side draw from near the bottom of the column in line 258″ to provide C4+ vapor and a bottoms stream in line 254 is forwarded to the second debutanizer column 260. The embodiment of
In this embodiment, it is preferred that all streams making up the second hydrocarbon feed in line 190″ are vaporous, obviating vaporizers.
In an embodiment shown in
The gas recovery system 420 is different in
The naphtha splitter column 480 may split naphtha into a heavy naphtha bottoms, typically C7+, in line 492 which may be recovered in line 184 with control valve thereon open and control valve on line 285 closed or further processed in line 285 with control valve thereon open and control valve on line 184 closed. An overhead stream from the naphtha splitter column 480 may carry light naphtha in line 482, typically a C7− material, to the primary absorber column 140. An overhead stream in line 154 from a depropanizer column 250 may join the compressed gas stream in line 134 to enter the primary absorber column 140 which is in downstream communication with the naphtha splitter column 480. In this location the naphtha splitter column 480 may be operated at a top pressure to keep the overhead in liquid phase, such as about 344 to about 3034 kPa (gauge) (50 to 150 psig) and a temperature of about 135 to about 191° C. (275 to 375° F.).
The gaseous hydrocarbon streams in lines 134 and 154 fed to the primary absorber column 140 are contacted with naphtha from the naphtha splitter overhead in line 482 to effect a separation between C3+ and C2− hydrocarbons by absorption of the heavier hydrocarbons into the naphtha stream upon counter-current contact. A debutanized naphtha stream in line 168 from the bottom of a debutanizer column 460 is delivered to the primary absorber column 140 at a higher elevation than the naphtha splitter overhead stream in line 482 to effect further separation of C3+ from C2− hydrocarbons. The primary absorber column 140 utilizes no condenser or reboiler but may have one or more pump-arounds to cool the materials in the column. A liquid C3+ stream in line 142 from the bottoms of the primary absorber column is returned to line 131 upstream of condenser to be cooled and returned to the second compressor receiver 132. An off-gas stream in line 148 from a top of the primary absorber 140 is directed to a lower end of a secondary or sponge absorber 150. A circulating stream of LCO in line 152 diverted from line 110 absorbs most of the remaining C5+ material and some C3-C4 material in the off-gas stream in line 148 by counter-current contact. LCO from a bottom of the secondary absorber in line 156 richer in C3+ material than the circulating stream in line 152 is returned in line 156 to the main column 90 via the pump-around for line 110. The overhead of the secondary absorber 150 comprising dry gas of predominantly C2− hydrocarbons with hydrogen sulfide, amines and hydrogen is removed in line 158 and may be subjected to further separation to recover ethylene and hydrogen.
Liquid from a bottom of the second compressor receiver 132 in line 144 is sent to the stripper column 146. Most of the C2− material is stripped from the C3-C7 material and removed in an overhead of the stripper column 146 and returned to line 131 via overhead line 138 without first undergoing condensation. The overhead gas in line 138 from the stripper column comprising C2− material, LPG and some light naphtha is returned to line 131 without first undergoing condensation. Therefore, only light naphtha is circulated in the gas recovery section 420. The condenser on line 131 will partially condense the overhead stream from line 138 with the gas compressor discharge in line 131 and with the bottoms stream 142 from the primary absorber column 140 will undergo vapor-liquid separation in second compressor receiver 132. The stripper column 146 is in downstream communication with the first reactor 10, a bottom of the second compressor receiver 132, a bottom of the primary absorber 140 and an overhead of the naphtha splitter 480 via the primary absorber column. The bottoms product of the stripper column 146 in line 162 is rich in light naphtha.
In the LPG splitter column 170, C3 materials may be forwarded from the overhead in a line 174 to a C3 splitter to recover propylene product. C4 materials from the bottom in line 476 may be recovered for blending in a gasoline pool as product or further processed.
In an embodiment, C4 material in line 476 may be delivered as a second hydrocarbon feed to a second catalytic reactor 200 which is in downstream communication with an overhead of the main fractionation column 100, a bottom of the primary absorber 140 and a bottom of the LPG splitter 170. In an embodiment, the C4 stream in line 476 may be vaporized in evaporator 488 from which vaporized naphtha exits in line 490 and is preferably superheated before it is fed to the second catalytic reactor 200. The second catalytic reactor 200 is in downstream communication with the vaporizer 488. In an embodiment, a light naphtha stream may be withdrawn from a side of the debutanizer 460 as a side cut in line 483. The side cut may be taken from a vapor side draw to avoid having to vaporize a liquid stream in an evaporator. The side cut naphtha in line 483 may be mixed with the vaporized C4 stream in line 490 to provide second hydrocarbon feed in line 191, so the second reactor 200 may be in downstream communication with the first debutanizer column 460 via the vapor side draw. A heat exchanger on line 191 may superheat the vaporized second hydrocarbon feed. The vapor side draw for line 483 should be in the lower half of the first debutanizer column 460 and below the feed entry for line 162. If a naphtha side cut is taken in line 483, very little flow may be taken through a control valve on line 472 under normal operation and may be omitted. Line 472, however, may still be used to control build up of heavy naphtha if they make their way to debutanizer column 460.
Operation of the second reactor 200 in
In a further embodiment, a bottoms stream from the naphtha splitter may be diverted in line 285 through open control valve thereon to a second naphtha splitter column 290. The second naphtha splitter column may have a dividing wall 292 interposed between a feed inlet and a mid-cut product outlet for line 296. The dividing wall has top and bottom ends spaced from respective tops and bottoms of the second naphtha splitter column 290, so fluid can flow over and under the dividing wall 292 from one side to the opposite side. The naphtha splitter may provide an overhead product of middle naphtha in line 294, an aromatics rich naphtha product through the mid-cut product outlet in the line 296 and a heavy naphtha in bottoms product line 298. The second naphtha splitter column 290 may be used in any of the embodiments herein.
In another embodiment shown in
The gas recovery section 520 is different in
A liquid bottoms stream from the stripper column 146 is sent to a first depropanizer column 560 via line 162. The first depropanizer column 560 is in downstream communication with the first reactor 10 and fractionates a portion of first cracked products from the first reactor 10 to provide a C3− overhead stream and C4+ bottoms stream. The overhead stream in line 564 from the first depropanizer column comprises C3 olefinic product which can be sent to a propane/propylene splitter (not shown) which may be in communication with an overhead of the depropanizer column 560. The bottoms stream in line 566 may be split between line 568 for delivering depropanized naphtha to the primary absorber 140 to assist in the absorption of C3+ materials and line 572 for recycle to the naphtha splitter column 480 or product recovery in line 473.
In an embodiment, a light naphtha stream may be withdrawn from a side of the first depropanizer column 560 as a side cut in line 583 taken below the feed entry point for line 162. The side cut may predominantly comprise C4-C7 hydrocarbons. The side cut may be from a vapor side draw to avoid having to vaporize a liquid stream in an evaporator. The side cut naphtha in line 583 may provide all of the second hydrocarbon feed in line 191 or may be mixed with vaporous depropanized side draw material in recycle line 556 to provide the second hydrocarbon feed in line 191. The second reactor 200 may be in downstream communication with the first depropanizer column 560 via the vapor side draw feeding line 583. A heat exchanger on line 191 may superheat the vaporized second hydrocarbon feed.
Operation of the second reactor 200, in downstream communication with the depropanizer column 560, and the second product recovery section 530 is generally as is described with respect to
Without further elaboration, it is believed that one skilled in the art can, using the preceding description, utilize the present invention to its fullest extent. The preceding preferred specific embodiments are, therefore, to be construed as merely illustrative, and not limitative of the remainder of the disclosure in any way whatsoever.
In the foregoing, all temperatures are set forth in degrees Celsius and, all parts and percentages are by weight, unless otherwise indicated. Additionally, control valves expressed as either open or closed can also be partially opened to allow flow to both alternative lines.
From the foregoing description, one skilled in the art can easily ascertain the essential characteristics of this invention and, without departing from the spirit and scope thereof, can make various changes and modifications of the invention to adapt it to various usages and conditions.
This application is a Continuation of application Ser. No. 12/614,938 filed Nov. 9, 2009, now U.S. Pat. No. 8,354,018, the contents of which are hereby incorporated by reference in its entirety.
| Number | Name | Date | Kind |
|---|---|---|---|
| 3477946 | Borst, Jr. | Nov 1969 | A |
| 8354018 | Leonard et al. | Jan 2013 | B2 |
| 20080035527 | Eng et al. | Feb 2008 | A1 |
| Entry |
|---|
| Sadeghbeigi, R., Fluid Catalytic Cracking Handbook, 2nd Ed., 2000, Gulf Publishing, pp. 25-27. |
| Number | Date | Country | |
|---|---|---|---|
| 20130131406 A1 | May 2013 | US |
| Number | Date | Country | |
|---|---|---|---|
| Parent | 12614938 | Nov 2009 | US |
| Child | 13737506 | US |
