Method for producing propylene from refinery dry gas

Abstract

The process converts ethylene in a diluted ethylene (dry gas) stream, which may be derived from FCC products, to propylene. The process includes an oligomerization step of converting ethylene in the diluted ethylene stream to heavier olefins and a second reaction step of converting the heavier olefins to propylene. The oligomerization catalyst may be an amorphous silica-alumina base having a Group VIII and / or VIB metal. The catalyst is resistant to feed impurities such as hydrogen sulfide, carbon oxides, hydrogen and ammonia. At least 50% by weight of the ethylene in the diluted ethylene stream can be converted to CC olefins and recycled to the FCC for cracking to propylene.

Description

[0001] Priority declaration

[0002] This application claims priority to U.S. Provisional Application No. 63 / 412,396, filed on October 1, 2022, which is incorporated herein in its entirety. Technical Field

[0003] The state of the art is a process for converting dilute ethylene in a hydrocarbon stream to propylene. The process includes oligomerization integrated with an FCC unit. Background Art

[0004] Catalytic cracking can form a variety of products from larger hydrocarbons. Typically, a heavy hydrocarbon feed, such as vacuum gas oil, is provided to a catalytic cracking reactor, such as a fluid catalytic cracking (FCC) reactor. A variety of products can be produced from such systems, including gasoline products and / or light products, such as propylene and / or ethylene.

[0005] In such systems, a secondary reactor can be used to further convert low-value products from the primary reactor. Although the use of a secondary reactor may incur additional capital costs, its catalyst, operating conditions, and feed (recycled low-value streams) can be tailored to maximize high-value products, such as light olefins including propylene and / or ethylene, to provide an overall benefit.

[0006] Light olefin production is essential to produce enough plastics to meet global demand. Propylene is very valuable for conversion into various grades of polypropylene (PP) and for the production of chemicals.

[0007] Dry gas is the common name of the tail gas stream from the fluid catalytic cracking unit, and it comprises all gases with a boiling point lower than ethane. The tail gas stream is compressed to remove as many C3 and C4 gases as possible. Sulfur is also absorbed in a large amount from the tail gas stream in the scrubber utilizing an amine absorbent. The remaining stream is referred to as FCC dry gas. Typical dry gas stream comprises 5 wt % to 50 wt % of ethylene, 10 wt % to 20 wt % of ethane, 0.5 wt % to 20 wt % of hydrogen, 5 wt % to 20 wt % of nitrogen, each 0.1 wt % to 5.0 wt % of carbon monoxide and carbon dioxide, and less than 0.01 wt % of hydrogen sulfide and ammonia, and the remainder is methane.

[0008] Currently, the FCC dry gas stream is sent to the burner as fuel gas. When the FCC unit, which processes 7,949 kiloliters (50,000 barrels) per day, is in operation, it will burn 181,000 kg (200 tons) of dry gas as fuel and 36,000 kg (40 tons) of ethylene per day to maximize gasoline production. However, as FCC operations shift to high propylene production, the amount of dry gas produced can be as high as 427,000 kg (470 tons) and ethylene is 210,000 kg (231 tons). Because there is a large price difference between fuel gas and light olefins, it seems economically advantageous to try to recover this ethylene. However, the dry gas stream contains impurities that can poison the catalyst and is too dilute, so it is not economically reasonable to recover the ethylene through the gas recovery system.

[0009] Olefin oligomerization to liquid products generally involves the use of propylene or butenes, particularly propylene or butenes from liquefied petroleum gas (LPG) or dehydrogenation feedstocks, to produce olefins in the gasoline or diesel range. In these processes, dilute ethylene is rarely used as a feedstock for oligomerization because its reactivity is much lower. Concentrated ethylene streams are oligomerized to liquid products that are not suitable for dry gas ethylene oligomerization because the catalyst is incompatible with trace impurities. U.S. Patent No. 8,575,410 describes a method for converting ethylene in dry gas to diesel range liquid fuels, but this method does not obtain the full value of dry gas ethylene because the value of liquid fuels is still lower than that of light olefins.

[0010] There is a need to further utilize the dilute ethylene in refinery streams. Summary of the invention

[0011] We find that a dry gas stream is fed to an impurity-resistant oligomerization catalyst to convert ethylene into butenes and higher olefins, which can provide a stream for recycling to an FCC unit. We find that most catalysts suitable for ethylene oligomerization can be rapidly deactivated in the presence of impurities such as carbon oxides, ammonia and hydrogen sulfide. These impurities do not substantially affect the catalysts of the 6th, 8th, 9th and 10th Group metals of the periodic table contained on an amorphous silica-alumina carrier. The cracking of the oligomer stream in the FCC can produce a large amount of propylene from the FCC unit. The unconverted gas can then be burned as a fuel gas, but the more valuable ethylene is converted into propylene. BRIEF DESCRIPTION OF THE DRAWINGS

[0012] Figure 1 is a schematic elevational view of the method of the present disclosure.

[0013] definition

[0014] The term "downstream communication" means that at least a portion of the fluid flowing to the body in the downstream communication can operatively flow from the object in fluid communication therewith.

[0015] The term "upstream communication" means that at least a portion of the fluid flowing out of the body in the upstream communication can be operatively flowed to the object in fluid communication therewith.

[0016] The term "direct communication" means that the fluid flows from the upstream component into the downstream component without passing through any other intervening container.

[0017] The term "indirect communication" means that the fluid flow from the upstream component enters the downstream component after passing through an intervening container.

[0018] The term "bypass" means that the object loses downstream communication with the bypassed body at least within the scope of the bypass.

[0019] As used herein, the term "predominantly" or "majority" means greater than 50%, suitably greater than 75%, and preferably greater than 90%.

[0020] The term "tower" means one or more distillation towers for separating one or more components with different volatilities. Unless otherwise specified, each tower includes a condenser on the top of the tower for condensing a portion of the overhead stream and refluxing it back to the top of the tower, and a reboiler at the bottom of the tower for vaporizing a portion of the bottom stream and sending it back to the bottom of the tower. The feed to the tower can be preheated. The top pressure is the pressure of the overhead vapor at the vapor outlet of the tower. The bottom temperature is the liquid bottom outlet temperature. The top line and the bottom line refer to the net line from any reflux or reboil to the tower downstream. The stripping tower can omit the reboiler at the bottom of the tower, but provide heating requirements and separation power for the liquefied inert medium (such as steam). The stripping tower is usually fed from the top tray and the main product is taken out from the bottom.

[0021] As used herein, the term "separator" means a vessel having an inlet and at least one overhead vapor outlet and one bottom liquid outlet, and may also have an outlet for an aqueous stream from a boot. A flash tank is a type of separator that may be connected downstream to a separator that may operate at a higher pressure.

[0022] As used herein, the term "boiling point temperature" means the atmospheric equivalent boiling point (AEBP) calculated from the observed boiling temperature and the distillation pressure, as calculated using the formula provided in ASTM D1160 Appendix A7, entitled "Practice for Converting Observed Vapor Temperatures to Atmospheric Equivalent Temperatures."

[0023] As used herein, the term "true boiling point" (TBP) means a test method for determining the boiling point of a substance in accordance with ASTM D-2892 for producing standardized masses of liquefied gases, distillate fractions, and residues for which analytical data are available, and for determining the yields of the above fractions by both mass and volume from which a plot of distillation temperature versus mass % is obtained in a column having a reflux ratio of 5:1 using fifteen theoretical plates.

[0024] As used herein, "bitumen" means a hydrocarbon material boiling above 524°C (975°F) AEBP as determined by any standard gas chromatography simulated distillation method such as ASTM D2887, D6352, or D7169 (all of which are used by the petroleum industry).

[0025] As used herein, the term "T5" or "T95" means the temperature at which 5 mass percent or 95 mass percent of a sample boils, respectively, using ASTM D-86 or TBP, as the case may be.

[0026] As used herein, the term "initial boiling point" (IBP) means the temperature at which a sample begins to boil using ASTM D-7169, ASTM D-86, or TBP, as appropriate.

[0027] As used herein, the term "endpoint" (EP) means the temperature at which a sample is brought to a complete boil using ASTM D-7169, ASTM D-86, or TBP, as appropriate.

[0028] As used herein, "vacuum gas oil" means a hydrocarbon material produced by vacuum fractionation of atmospheric residue, having an IBP of at least 232°C (450°F), a T5 between 288°C (550°F) and 392°C (700°F), usually not exceeding 343°C (650°F), a T95 between 510°C (950°F) and 570°C (1058°F), and / or an EP not exceeding 626°C (1158°F), as determined by any standard gas chromatography simulated distillation method such as ASTM D2887, D6352, or D7169 (all of which are used by the petroleum industry).

[0029] As used herein, "atmospheric residue" refers to a hydrocarbon material obtained from the bottom of an atmospheric crude distillation column, having an IBP of at least 232°C (450°F), a T5 between 288°C (550°F) and 392°C (700°F), usually not exceeding 343°C (650°F), and a T95 between 510°C (950°F) and 700°C (1292°F).

[0030] As used herein, "atmospheric residue" means a hydrocarbon material boiling at an IBP of at least 500°C (932°F). DETAILED DESCRIPTION

[0031] The present disclosure relates to forming propylene in an FCC unit and can be applied to any hydrocarbon stream containing ethylene and preferably a dilution ratio of ethylene. Suitable dilute ethylene streams can generally contain ethylene between 5 wt % and 65 wt %. FCC dry gas stream is a suitable dilute ethylene stream. Other dilute ethylene streams can also be used in the present disclosure, such as coker dry gas streams. Because the present disclosure is particularly applicable to FCC dry gas, the present application will be described in conjunction with utilizing ethylene from the FCC dry gas stream.

[0032] Now go to Figure 1 , wherein like numbers represent like parts, the method and apparatus generally include an FCC unit section 6 and a product recovery section 90. The FCC unit section 6 includes a first FCC reactor 10, which includes a first reactor unit 12 and a catalyst regenerator 14. The process conditions in the first FCC reactor 10 may include a cracking reaction temperature of 400° C. to 600° C., preferably 538° C. to 593° C. at the reactor outlet, and a catalyst regeneration temperature of 500° C. to 900° C. Cracking and regeneration are both carried out at an absolute pressure between 100 kPa (14 psia) and 650 kPa (94 psia), preferably between 140 kPa (20 psia) and 450 kPa (65 psia).

[0033] Figure 1 A first FCC reactor vessel 12 is shown wherein a first hydrocarbon feedstock in line 15 through distributor 16 is contacted with a first fluid catalyst stream entering from regenerated catalyst standpipe 18 and recycle catalyst standpipe 19. The first hydrocarbon feedstock may comprise vacuum gas oil, atmospheric residue, deasphalted oil, vacuum residue, or any other stream processed in a conventional FCC unit.

[0034] The catalyst can be a single catalyst or a mixture of different catalysts. Typically, the catalyst includes two components or catalysts, i.e., a first component or catalyst and a second component or catalyst. Such catalyst mixtures are disclosed in, for example, U.S. Patent No. 7,312,370B2. Typically, the first component may include any known catalyst used in the FCC field, such as an active amorphous clay-type catalyst and / or a highly active crystalline molecular sieve. Zeolite can be used as a molecular sieve in the FCC process. Preferably, the first component includes a large pore zeolite (such as Y-type zeolite), an active alumina material, a binder material (including silica or alumina) and an inert filler (such as kaolin).

[0035] Typically, the zeolite molecular sieve suitable for the first component has a large average pore size. Typically, the molecular sieve with a large pore size has a hole, wherein the opening of the hole is greater than 0.7nm, and the effective diameter is defined by more than 10 and typically 12 component rings. The pore size index of the macropore may be above 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 a zeolite) may have any suitable amount of rare earth metal or rare earth metal oxide.

[0036] The second component may include a mesoporous or relatively small pore zeolite catalyst, such as an MFI zeolite, for example at least one of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other similar materials. Other suitable mesoporous or relatively small pore zeolites include ferrierite and erionite. Preferably, the second component is a mesoporous or relatively small pore zeolite dispersed on a matrix, the matrix including a binder material (such as silica or alumina) and an inert filler (such as kaolin). The second component may also include some other active materials, such as beta zeolite. These compositions may have a crystalline zeolite content of 10% to 50% by weight or more and a matrix material content of 50% to 90% by weight. Preferably, a component comprising 40% by weight of a crystalline zeolite material, and those components having a larger crystalline zeolite content may be used. Typically, mesoporous and relatively small pore zeolites are characterized by having an effective pore opening diameter less than or equal to 0.7nm, a ring of 10 yuan or less yuan and an aperture index less than 31.

[0037] The total catalyst mixture in the first FCC reactor 12 may include 1 wt % to 25 wt % of the second component, i.e., medium-pore to small-pore crystalline zeolite, preferably greater than or equal to 1.75 wt % of the second component. The first component may account for the remainder of the catalyst composition. In some preferred embodiments, the relative proportions of the first component and the second component in the mixture may remain essentially unchanged throughout the first FCC reactor 12. High concentrations of medium-pore or small-pore zeolites as the second component of the catalyst mixture may improve the selectivity to light olefins. In an exemplary embodiment, the second component may be a ZSM-5 zeolite, and the mixture may include 4 wt % to 10 wt % of the ZSM-5 zeolite, excluding any other components, such as a binder and / or filler.

[0038] Preferably, at least one of the first and / or second catalysts is an MFI zeolite having a silicon to aluminum ratio greater than 15, preferably greater than 75. In an exemplary embodiment, the silicon to aluminum ratio may be from 15:1 to 35:1.

[0039] This contact may occur in a narrow first riser 20 that extends upward to the bottom of a first reactor vessel 22. The contact of the first hydrocarbon feedstock with the first fluid catalyst stream is fluidized by a gas, such as steam, from a fluidizing distributor 24. In one embodiment, heat from the catalyst vaporizes the first hydrocarbon feedstock, which is then cracked in the presence of the first catalyst stream into a first cracked product stream of lighter molecular weight as both are transferred upward along the riser 20 into the reactor vessel 22, thereby providing a first mixture of catalyst and product gas.

[0040] The pressure in the first riser 20 can be 200 kPa (29 psia) to 450 kPa (65 psia), but can also be lower. The steam flow rate added to the first riser 20 is 3% to 7% by weight of the first hydrocarbon feedstock. Inevitable side reactions occur in the first riser 20, leaving coke deposits on the catalyst, thereby reducing the catalyst activity and producing a spent catalyst stream. The first cracked product stream in the first mixture of catalyst and product gas is then separated from the spent catalyst stream using a cyclone separator, which may include a single or two-stage cyclone separator 62 in the reactor vessel 22. The gaseous first cracked product stream leaves the reactor vessel 22 through the first product outlet 31 and enters the pipeline 32 for transportation to the downstream product recovery section 90.

[0041] The spent catalyst or coked catalyst requires regeneration for further use. The spent catalyst stream, after being separated from the first cracked product stream by the disengagement device 54 in the first disengagement chamber 56, falls into the stripping section 34, where steam is injected through the distributor 35 to remove any residual hydrocarbon vapor. The transport conduit 60 transfers the hydrocarbon vapor (including the stripped hydrocarbon, the stripping medium and the entrained catalyst) to the first reactor vessel 22. The cyclone separator separates the coked catalyst from the hydrocarbon vapor stream. After the stripping operation, the stripped coked catalyst is carried to the catalyst regenerator 14 through the spent catalyst riser 36. Another portion of the stripped coked catalyst can be recycled to the riser 20 through the recycle catalyst riser 19 without undergoing regeneration.

[0042] Figure 1A regenerator 14 referred to as a burner is shown. However, other types of regenerators are also suitable. In the catalyst regenerator 14, a stream of an oxygen-containing gas (such as air) is introduced through an air distributor 38 to contact the coked catalyst. The coked catalyst is burned in a combustion chamber 80 to form coke to provide a regenerated catalyst and flue gas. The catalyst regeneration process adds a large amount of heat to the catalyst, thereby providing energy to offset the endothermic cracking reaction occurring in the reactor riser 20. The catalyst and air flow upward together in the combustion chamber 80 of the regenerator 14, and after regeneration, are discharged through the settler 40 for preliminary separation and enter the separation chamber 86. The regenerated catalyst and flue gas leaving the settler 40 are additionally recovered using the first stage cyclone separator 44 and the second stage cyclone separator 46 in the separation chamber 86 of the catalyst regenerator 14, respectively. The catalyst separated from the flue gas is distributed through the feed pipes in the cyclone separators 44, 46, while the relatively light flue gas in the catalyst leaves the cyclone separators 44, 46 in turn and leaves the regenerator vessel 14 through the flue gas outlet 47 in the flue gas line 48. The regenerated catalyst is carried back to the standpipe 20 through the regenerated catalyst riser 18. Due to the combustion of the coke, the flue gas vapor leaving the top of the catalyst regenerator 14 in the pipeline 48 contains CO, CO2, N2 and H2O together with smaller amounts of other substances. The hot flue gas leaves the regenerator 14 through the flue gas outlet 47 in the pipeline 48 for further processing.

[0043] The product recovery section 90 is in downstream communication with the product outlet 31 and the second product outlet 231. "Communication" means that the material flow is operably allowed between the enumerated components. In the product recovery section 90, the gaseous FCC product in the pipeline 32 is guided to the lower section of the FCC main fractionation tower 92 via the combined gaseous FCC product in the pipeline 33. The main tower 92 is in downstream communication with the product outlet 31 and the second product outlet 231. Several fractions of the FCC product can be separated and taken out from the main tower, including the heavy slurry oil from the bottom in the pipeline 93, the heavy cycle oil stream in the pipeline 94, the light cycle oil in the pipeline 95 taken out from the outlet 95a, and the heavy naphtha stream in the pipeline 96 taken out from the outlet 96a. Any one or all of the pipelines from the pipeline 93 to the pipeline 96 can be cooled and pumped back to the main tower 92 to cool the main tower usually at a higher position. Gasoline and gaseous light hydrocarbons are removed from main column 92 in overhead line 97 and condensed before entering main column receiver 99. Main column receiver 99 is in downstream communication with product outlet 31 and second product outlet 231, and main column 92 is in upstream communication with main column receiver 99. "Upstream communication" means that at least a portion of the material flowing out of a component in upstream communication can operatively flow to a component in communication therewith.

[0044] The aqueous stream is removed from a storage tank in receiver 99. In addition, a condensed light naphtha stream is removed in line 101, while an overhead stream is removed in line 102. The overhead stream in line 102 contains gaseous light hydrocarbons, which may contain a diluent ethylene stream. The streams in lines 101 and 102 may enter a vapor recovery section 120 of product recovery section 90.

[0045] The vapor recovery section 120 shown is an absorption based system, but any vapor recovery system may be used, including a cold box system. In order to fully separate the light gas components, the gaseous stream in line 102 is compressed in compressor 104. More than one compressor stage may be used, but typically two-stage compression is used. The compressed light hydrocarbon stream in line 106 is combined with the streams in lines 107 and 108, cooled and delivered to a high pressure receiver 110. The aqueous stream in line 111 from receiver 110 may be sent to the main column receiver 99. The gaseous hydrocarbon stream in line 112 containing the dilution ethylene stream is directed to a primary absorber 114 where it contacts the unstabilized gasoline in line 101 from the main column receiver 99 to achieve C3 + Hydrocarbons and C2 - The primary absorber 114 is connected to the downstream of the main tower receiver 99. The liquid C3 in the pipeline 107 + The stream is returned to line 106 before cooling. For purposes of this disclosure, the primary tail gas stream in line 116 from primary absorber 114 comprises a dilute ethylene stream. However, to further concentrate the ethylene stream and recover heavier components, line 116 may optionally be directed to secondary absorber 118 where a recycle stream of light cycle oil in line 121 diverted from line 95 absorbs most of the remaining C5 + and some C3-C4 material. A secondary absorber 118 is in downstream communication with the primary absorber 114. The C3-rich + The light cycle oil of the material is circulated back to the main tower 92 via the pump of pipeline 95. The top of the secondary absorber 118 contains mainly C2 - The hydrocarbons, along with a dry gas of hydrogen sulfide, ammonia, carbon oxides, and hydrogen, are removed in a secondary tail gas stream in line 122 to contain a diluent ethylene stream.

[0046] Liquid from high pressure receiver 110 in line 124 is sent to stripping column 126. Most of the C2 - The first debutanizer 130 provides an overhead stream in line 132, which comprises a C3-C4 hydrocarbon stream from the first debutanizer. The bottom stream in line 134 may comprise a first debutanizer naphtha stream.

[0047] In a first embodiment, a first recycle light cracked naphtha stream may be taken from the first debutanized naphtha stream in line 134 in line 137 via a control valve thereon, while the remainder of the first debutanized naphtha stream in line 141 may be passed through a control valve thereon for further processing into gasoline or other products. In this embodiment, it is contemplated that the naphtha splitter 136 may be located upstream of the product recovery section 90.

[0048] In an alternative embodiment, the first debutanized naphtha stream in line 134 can be fed to a naphtha splitter 136 in line 142 via a control valve thereon. In this alternative embodiment, the naphtha splitter is located downstream of the product recovery section 90, such as Figure 1 As shown, the control valves on lines 137 and 141 will be closed or mostly closed. The naphtha splitter separates the debutanized naphtha stream into a first separated light naphtha stream containing C5-C7 hydrocarbons in an overhead line 138 and a heavy naphtha stream in a bottoms line 140. An alternative first recycle light cracked naphtha stream may be taken in line 139 via a control valve thereon, and the remainder of the first separated light naphtha stream in the overhead line 138 may be further processed into gasoline or other products.

[0049] The C3-C4 hydrocarbon stream obtained in the pipeline 132 can be separated into a C3 hydrocarbon stream in the tower top pipeline 146 and a C4 hydrocarbon stream in the tower bottom pipeline 148 in the C3-C4 separation tower 144. A first recycled C4 hydrocarbon stream can be obtained in the pipeline 149 through a control valve thereon, and a second C4 hydrocarbon stream in the pipeline 147 can be further processed into other products. The C3 hydrocarbon stream in the tower top pipeline 146 can be further processed for propylene recovery.

[0050] One or both of the first recycled light cracked naphtha stream comprising olefinic C5-C7 hydrocarbons taken from the first debutanizer naphtha stream in line 134 of the first debutanizer 130 connected downstream to the main tower 92 in line 137 or the alternative first recycled light cracked naphtha stream comprising olefinic C5-C7 hydrocarbons taken from the tower top line 138 of the naphtha splitter 136 connected downstream to the main tower 92 in line 139 and the first recycled C4 stream comprising olefinic C4 hydrocarbons taken from the tower bottom line 148 of the C3-C4 splitter also connected downstream to the main tower 92 in line 149 may be circulated to the second FCC reactor 202 as the second hydrocarbon stream in the second loading line 150. The second hydrocarbon stream may be preheated to a temperature of 221° C. (400° F.) to 704° C. (1300° F.) and charged to the second FCC reactor 202.

[0051] The dilute ethylene stream may be a secondary tail gas stream in line 122 which may comprise an FCC dry gas stream comprising between 5 wt% and 65 wt% ethylene, preferably 10 wt% to 55 wt% ethylene. Methane is typically the major component in the dilute ethylene stream, with a concentration between 25 wt% and 55 wt%, with ethane being substantially present typically at 5 wt% to 45 wt%. Hydrogen and nitrogen may each be present in the dilute ethylene stream at between 0.5 wt% and 25 wt% and typically 1 wt% to 20 wt%. Saturated levels of water may also be present in the dilute ethylene stream. If a secondary absorber 118 is used, no more than 5 wt% of C3 + and typically less than 0.5% by weight of propylene.

[0052] In addition to hydrogen, other impurities such as hydrogen sulfide, ammonia, carbon oxides and acetylene may also be present in the dilute ethylene stream.

[0053] We have found that many impurities in the dry gas ethylene stream can poison the oligomerization catalyst. Hydrogen and carbon monoxide can reduce metal sites, causing them to become deactivated. Carbon dioxide and ammonia can attack acid sites on the catalyst. Hydrogen sulfide can attack the metal on the catalyst, producing metal sulfides. Acetylene can polymerize and bind to the catalyst or equipment.

[0054] The secondary tail gas stream in line 122 comprising a diluent ethylene stream may be introduced into an optional amine absorber unit 160 to remove hydrogen sulfide to a lower concentration. A lean aqueous amine solution, such as one comprising monoethanolamine or diethanolamine, is introduced into absorber 160 via line 162 and contacted with the flowing secondary tail gas stream to absorb hydrogen sulfide, and a rich amine aqueous absorbent solution containing hydrogen sulfide is removed from absorption zone 160 via line 163 and recovered and possibly further processed.

[0055] The amine treated dilute ethylene stream in line 164 may be introduced into an optional water wash unit 166 to remove residual amine carried over from the amine absorber 160 and to reduce the concentration of ammonia and carbon dioxide in the dilute ethylene stream in line 164. Water is introduced into the water wash unit in line 165. The water in line 165 is typically slightly acidified to enhance capture of basic molecules such as amines. An aqueous stream in line 167 rich in amines and possibly ammonia and carbon dioxide leaves the water wash unit 166 and may be further processed.

[0056] In one aspect, the optional amine-treated dilute ethylene and possible water-washed stream in line 168 can then be treated in an optional guard bed 170 to remove one or more impurities such as carbon monoxide, hydrogen sulfide, and ammonia to lower concentrations. The guard bed 170 can contain an adsorbent to adsorb impurities that may poison the oligomerization catalyst, such as hydrogen sulfide. The guard bed 170 can contain a variety of adsorbents for adsorbing more than one type of impurity. Typical adsorbents for adsorbing hydrogen sulfide are ADS-12, for adsorbing CO are ADS-106, and for adsorbing ammonia are UOP MOLSIV 3A, all of which are available from UOP, LLC. The adsorbents can be mixed in a single bed, or can be arranged in a continuous bed.

[0057] In another aspect, the optional amine-treated diluted ethylene and possible water-washed stream in line 168 can be treated to recover hydrogen in the diluted ethylene stream in line 168 or reduce its concentration to a lower level. In an exemplary embodiment, unit 170 is a pressure swing adsorption (PSA) unit, rather than a guard bed. The optional amine-treated diluted ethylene and possible water-washed stream in line 168 can be sent to a pressure swing adsorption (PSA) unit 170 to recover hydrogen in the diluted ethylene stream in line 168 or reduce its concentration to a lower level. The optional amine-treated diluted ethylene and possible water-washed stream in line 168 can enter PSA 170 at high pressure. Molecules heavier than hydrogen present in the optional amine-treated diluted ethylene and possible water-washed stream in line 168 are adsorbed on the adsorbent, while hydrogen molecules are not adsorbed on the adsorbent and pass through the bed. The adsorbent can be selected from one or more of silica gel, alumina, activated carbon, molecular sieves (including but not limited to zeolites), metal organic frameworks (MOFs), etc. or mixtures thereof. PSA 170 may include a single or multiple adsorbent beds of varying composition. Light molecules such as hydrogen leave PSA 170 at high pressure in line 173 and are recovered. The adsorbed heavy molecules are then desorbed from PSA 170 at a lower pressure and recovered in line 171 as a PSA tail gas stream containing ethylene.

[0058] In one embodiment, the PSA unit 170 may be operated at an adsorption pressure of 1500 kPa (220 psi) to 3500 kPa (515 psi) and a desorption pressure of 100 kPa (15 psi) to 300 kPa (44 psi).

[0059] In another aspect, the optional diluted ethylene treated through amines in pipeline 168 and the possible stream washed through water can pass through membrane unit 170, to recycle the hydrogen in the diluted ethylene stream in pipeline 168 or reduce its concentration to a lower level. In membrane unit 170, the optional diluted ethylene treated through amines in pipeline 168 and the possible stream washed through water pass through membrane under high pressure. In an exemplary embodiment, membrane unit 170 may include polyimide or polyethersulfone-polyimide blend membrane. Other suitable membranes may include polymer, porous ceramic, dense ceramic or metal film. After contact, the optional diluted ethylene treated through amines in pipeline 168 and the possible stream washed through water are present in the molecules heavier than hydrogen as retentate quilted membrane retention. The heavy molecules comprising ethylene leave membrane unit 170 under high pressure, and are discharged from membrane unit 170 in the form of the retentate stream comprising ethylene in pipeline 171. Hydrogen present in the optional amine treated dilute ethylene and possible water scrubbed stream in line 168 permeates through the membrane with higher selectivity than retained molecules and is recovered on the other side of the membrane at a lower pressure in the hydrogen stream in line 173.

[0060] The membrane unit 170 can be operated at a temperature of 40°C (104°F) to 80°C (176°F). The transmembrane pressure difference can be as low as 70 kPa (10 psi) or as high as 14.5 MPa (2100 psi), depending on many factors, such as the specific membrane used, the flow rate of the inlet stream, and the availability of a compressor to compress the permeate stream (if such compression is desired). In one embodiment, the transmembrane pressure difference can range between 446 kPa (50 psig) and 6996 kPa (1000 psig) feed pressure.

[0061] The dilute ethylene stream in line 171, which may be treated with amines, may be washed with water, and may be treated with adsorption to remove more hydrogen, hydrogen sulfide, ammonia, and carbon monoxide, typically has at least one of the following impurity concentrations: 0.1 wt % and up to 5.0 wt % carbon monoxide and / or 0.1 wt % and up to 5.0 wt % carbon dioxide, and / or at least 1 wppm and up to 500 wppm hydrogen sulfide and / or at least 1 wppm and up to 500 wppm ammonia, and / or at least 0.1 wt % and up to 10 wt % hydrogen. The types of impurities present and their concentrations will vary depending on the processing and source of the dilute ethylene stream.

[0062] Pipeline 171 delivers the diluted ethylene stream to compressor 172 to be pressurized to the reactor pressure. Compressor 172 is connected to main tower 92, product recovery section 90, first product outlet 31 and second product outlet 231 downstream. The compressed diluted ethylene stream can be compressed to at least 3,550kPa (500psia) and may not exceed 10,445kPa (1500psia) and is suitably between 4,930kPa (700psia) and 7,687kPa (1100psia). Preferably the diluted ethylene stream is pressurized to a critical pressure higher than ethylene, which is 4,992kPa (724psia) for pure ethylene, to avoid rapid catalyst deactivation. Compressor 172 may include one or more stages with interstage cooling. A heater may be required so that the compressed stream reaches the reaction temperature. The compressed diluted ethylene is delivered to oligomerization reactor 176 in pipeline 174.

[0063] The oligomerization reactor 176 is connected to the compressor 172 and the primary and secondary absorbers 114 and 118 downstream, respectively. The oligomerization reactor preferably comprises a fixed catalyst bed 178. The dilute ethylene feed stream preferably contacts the catalyst in a downward flow operation. However, an upward flow operation may be suitable. In one embodiment, the second C4 hydrocarbon stream in the pipeline 147 can also contact the catalyst via a control valve thereon through the pipeline 175. The catalyst is preferably an amorphous silica-alumina base material, with metals from the 6th, 8th, 9th and 10th groups of the periodic table of elements using IUPAC symbols. In one aspect, the catalyst has a metal of the 8th, 9th or 10th (hereinafter referred to as 8 to 10) group promoted with a 6th group metal. In one aspect, the catalyst may have a silica-alumina ratio of no more than 30 and preferably no more than 20. Typically, silica and alumina are only in the base material, so the silica / alumina ratio of the catalyst is the same as the base material. The metal can be impregnated on the silica-alumina base material or ion exchanged with the silica-alumina base material. Co-grinding is also contemplated. Additionally, suitable catalysts will have a 2 The surface area is between 1000 Å / g as determined by nitrogen BET.

[0064] The preferred oligomerization catalyst of the present disclosure is described as follows. The preferred oligomerization catalyst comprises an amorphous silica-alumina support. One of the components of the catalyst support used in the present disclosure is alumina. Alumina can be any of various hydrated aluminum oxides or alumina gels such as α-alumina monohydrate of boehmite or pseudo-boehmite structure, α-alumina trihydrate of gibbsite structure, β-alumina trihydrate of bayerite structure, etc. Particularly preferred alumina can be purchased from Sasol North America Alumina Product Group under the trade name Catapal. This material is an extremely high purity α-alumina monohydrate (pseudo-boehmite) that has been shown to produce high purity γ-alumina after calcination at high temperatures. Another component of the catalyst support is amorphous silica alumina. Suitable silica alumina with a silica to alumina ratio of 2.6 was purchased from CCIC of Japan (a subsidiary of JGC).

[0065] Another component for preparing the catalyst used in the present disclosure is a surfactant. Preferably, the surfactant is mixed with the above-mentioned aluminum oxide and silica alumina powder. The resulting mixture of surfactant, aluminum oxide and silica alumina is then formed as described below, dried and calcined. Calcination effectively removes the organic components of the surfactant by combustion, but only after the surfactant has faithfully performed its function according to the present disclosure. Any suitable surfactant can be used according to the present disclosure. A preferred surfactant is a surfactant selected from a series of commercial surfactants sold by Solvay SA under the trade name "Antarox". "Antarox" surfactants are generally characterized as modified straight chain aliphatic polyethers and are low foaming biodegradable detergents and wetting agents.

[0066] Suitable silica-alumina mixtures can be prepared by mixing proportions of silica-alumina and alumina in volume to achieve the desired silica to alumina ratio. In one embodiment, 85 wt. % amorphous silica-alumina (wherein the silica to alumina ratio is 2.6) and 15 wt. % alumina powder will provide a suitable support. In one embodiment, other ratios of amorphous silica-alumina to alumina other than 85:15 are also suitable.

[0067] The surfactant may be combined with the silica alumina and alumina mixture using any convenient method. The surfactant is preferably mixed during the mixing and forming of the alumina and silica alumina. A preferred method is to mix an aqueous solution of the surfactant with the blend of alumina and silica alumina prior to final formation of the carrier. Preferably, the surfactant is present in the paste or dough in an amount of 0.01 wt % to 10 wt % based on the weight of the alumina and silica alumina.

[0068] A monoprotic acid such as nitric acid or formic acid may be added to the mixture in aqueous solution to peptize the alumina in the binder.Additional water may be added to the mixture to provide sufficient moisture to form a dough of sufficient consistency to be extruded or spray dried.

[0069] The paste or dough may be prepared in the form of shaped particles, preferably by extruding a dough mixture of alumina, silica alumina, surfactant and water through a die having openings of the desired size and shape therein, after which the extruded mass is broken into extrudates of the desired length and dried. A further calcination step may be employed to impart increased strength to the extrudates. Typically, calcination is carried out at a temperature of 260°C (500°F) to 815°C (1500°F) in a stream of dry air.

[0070] The extruded particles may have any suitable cross-sectional shape, i.e., symmetrical or asymmetrical, but most typically have a symmetrical cross-sectional shape, preferably spherical, cylindrical, or multi-lobed. The cross-sectional diameter of the particles may be as small as 40 μm; however, it is typically 0.635 mm (0.025 inch) to 12.7 mm (0.5 inch), preferably 0.79 mm (1 / 32 inch) to 6.35 mm (0.25 inch), and most preferably 1.06 mm (1 / 24 inch) to 4.23 mm (1 / 6 inch). In a preferred catalyst configuration, the cross-sectional shape resembles a clover, as shown in, for example, FIG. 8 and FIG. 8A of U.S. Patent No. 4,028,227. Preferred clover-shaped particles are such that each "leaf" of the cross section is defined by a 270° arc of a circle having a diameter between 0.51 mm (0.02 inch) and 1.27 mm (0.05 inch). Other preferred particles are those having a four-lobed cross-sectional shape, including asymmetric shapes and symmetric shapes, such as shown in Figure 10 of US Patent No. 4,028,227.

[0071] The typical characteristics of the amorphous silica alumina support used herein are that the total pore volume, average pore size and surface area are large enough to provide a large amount of space and area for depositing active metal components. As measured by conventional mercury intrusion method, the total pore volume of the support is generally 0.2mL / g to 2.0mL / g, preferably 0.25mL / g to 1.0mL / g, and most preferably 0.3mL / g to 0.9mL / g. Typically, the amount of pore volume of the support in pores greater than 100 angstroms is less than 0.1mL / g, preferably less than 0.08mL / g, and most preferably less than 0.05mL / g. As measured by the BET method, the surface area is generally above 50m 2 / g, for example, higher than 200m 2 / g, preferably at least 250m 2 / g, and most preferably 300m 2 / g to 400m 2 / g.

[0072] In order to prepare the catalyst, the support material is compounded with one or more precursors of at least one metal component from the 6th, 8th, 9th and 10th groups of the periodic table of elements, such as by a single impregnation or multiple impregnations of calcined amorphous refractory oxide support particles. The 8th to 10th group metal (preferably nickel) should be present in a concentration of 0.5% to 15% by weight, and the 6th group metal (preferably tungsten) should be present in a concentration of 0% to 12% by weight. Impregnation can be achieved by any method known in the art (such as, for example, by spray impregnation), wherein a solution containing a metal precursor in a dissolved form is sprayed onto the support particles. Another method is a multiple dip (multi-dip) procedure, wherein the support material is repeatedly contacted with the impregnation solution, with or without intermittent drying. Other methods involve immersing the support in a large volume of impregnation solution or circulating the support therein, and another method is a pore volume or pore saturation technique, wherein the support particles are introduced into an impregnation solution having a volume just sufficient to fill the pores of the support. Sometimes, the pore saturation technique can be modified to utilize an impregnation solution having a volume that is 10% less to 10% greater than that which just fills the pores.

[0073] The catalyst may be further modified by compounding onto the support particles one or more precursors of metal components of Groups 1 and 2 of the Periodic Table of the Elements, such as sodium. As previously mentioned, the Group 1 and Group 2 metal precursors may be introduced by impregnation according to any method known in the art, and it may be done together with or separately from the impregnation of the Group 6, 8, 9, and 10 metal precursors.

[0074] If the active metal precursors are combined by impregnation, a subsequent or second calcination at an elevated temperature, such as, for example, between 399°C and 760°C (750°F and 1400°F), can convert the metals to their corresponding oxide forms. In some cases, calcination can be performed after each impregnation of the individual active metals. Subsequent calcinations will produce a catalyst containing the active metals in their corresponding oxide forms.

[0075] The preferred oligomerization catalyst of the present disclosure has an amorphous silica-alumina base material impregnated with 0.5 wt % to 15 wt % nickel in the form of 3.175 mm (0.125 inch) extrudates and a density of 0.45 g / mL to 0.65 g / mL. It is also contemplated that the metals may be bound to the support by other methods such as ion exchange and co-grinding.

[0076] An alternative catalyst suitable for use in the present disclosure utilizes a co-gelled silica-alumina support prepared by the well-known oil drop method, which allows the use of a support in the form of large spheres. For example, an alumina sol used as an alumina source is mixed with an acidified water glass solution as a silica source, and the mixture is further mixed with a suitable gelling agent (e.g., urea, hexamethylenetetramine, or a mixture thereof). While still below the gelling temperature, the mixture is discharged through a nozzle or a rotating disk into a hot oil bath maintained at the gelling temperature. The mixture is dispersed into the oil bath in the form of droplets, which form spherical gel particles during passage through the oil bath. The alumina sol is preferably prepared by mixing aluminum pellets with a certain amount of treated water or deionized water, to which sufficient hydrochloric acid is added to digest a portion of the aluminum metal and form the desired sol. At the reflux temperature of the mixture, a suitable reaction rate is affected.

[0077] Spherical gel particles prepared by the oil drop method are typically aged in an oil bath for at least 10 to 16 hours, then aged in a suitable alkali-containing or alkaline medium for at least 3 to 10 hours, and finally washed with water. Proper gelation of the mixture in the oil bath and subsequent aging of the gel spheres is not readily achieved at temperatures below 48.9°C (120°F), while at 98.9°C (210°F), the rapid release of gas tends to rupture and otherwise weaken the spheres. By maintaining sufficient superatmospheric pressure during the forming and aging steps to keep the water in the liquid phase, higher temperatures can be used, generally with improved results. If the gel particles are aged at superatmospheric pressure, the alkaline aging step is not required.

[0078] The balls are washed with water, preferably with water containing a small amount of ammonium hydroxide and / or ammonium nitrate. After washing, the balls are dried at a temperature of 93.3°C (200°F) to 315°C (600°F) for 6 to 24 hours or more, and then calcined at a temperature of 426.67°C (800°F) to 760°C (1400°F) for 2 to 12 hours or more.

[0079] The Group 8 to 10 components and the Group 6 components and optionally the Group 1 and 2 components are composited with the co-gelled silica-alumina support material by any suitable co-impregnation technique. Thus, the support material may be soaked, impregnated, suspended or otherwise immersed in an aqueous impregnation solution containing a soluble Group 8 to 10 salt and a soluble Group 6 salt. One suitable method comprises immersing the support material in the impregnation solution and evaporating it to dryness in a rotary steam dryer, the concentration of the impregnation solution being such that the atomic ratio of nickel to nickel plus tungsten in the final catalyst composite is from 0.1 to 0.3. Another suitable method comprises immersing the support material in the aqueous impregnation solution at room temperature until the support is completely penetrated by the solution. After the support absorbs the impregnation solution, the free surface liquid is drained and dried in a moving belt calciner. Another suitable method comprises adding the Group 8 to 10 components and / or the Group 6 components to a mixture comprising an alumina source, a silica source and a gelling agent prior to the co-gelling step.

[0080] The catalyst composite is typically dried at a temperature of 93.3°C (200°F) to 260°C (500°F) for 1 to 10 hours prior to calcination. According to the present disclosure, calcination is conducted in an oxidizing atmosphere at a temperature of 371°C (700°F) to 650°C (1200°F). The oxidizing atmosphere is suitably air, but other gases including molecular oxygen may also be used.

[0081] A suitable alternative catalyst is an oil droplet silica-alumina spherical support 3.175 mm (0.125 inch) in diameter impregnated with 0.5 to 15 wt % nickel and 0 to 12 wt % tungsten. Other metal bonding methods are also suitable and contemplated. Suitable density ranges for alternative catalysts are between 0.60 g / mL and 0.70 g / mL.

[0082] The dilute ethylene feed in line 174 may contact the oligomerization catalyst in catalyst bed 178 at a temperature between 100°C and 400°C. The reaction is primarily in the vapor phase at 50 hr based on olefins. -1 Up to 1000hr -1GHSV occurs. We surprisingly found that despite the presence of catalyst poisoning impurities in the feed and dilution of ethylene in the feed, at least 40 wt % and up to 75 wt % of the ethylene in the feed stream is converted to heavier hydrocarbons. Ethylene will first oligomerize to heavier olefins on the catalyst. Some of the heavier olefins can cyclize on the catalyst, and the presence of hydrogen can promote the conversion of olefins to paraffins, which are all hydrocarbons heavier than ethylene. However, cyclization and saturated products are undesirable for cracking to propylene and should be minimized by appropriate selection of catalyst and reaction conditions.

[0083] The catalyst can remain stable despite impure feeds, but it can be regenerated when deactivated. Suitable regeneration conditions may include subjecting the catalyst to hot air at 500°C for 3 hours, for example, in situ. The activity and selectivity of the regenerated catalyst may be comparable to that of the fresh catalyst.

[0084] The oligomerization product stream from the oligomerization reactor in line 180 can be transported to an oligomerization separation unit 182. The oligomerization separation unit can be a simple flash tank to separate the gaseous stream from the liquid stream, but can also be more complex and include additional input streams (not shown), such as an absorption tower with unstable gasoline or light cycle oil to improve the recovery efficiency of the oligomers. The oligomerization separation unit 182 is connected to the downstream of the oligomerization reactor 176. The light raffinate stream containing light gases such as hydrogen, methane, ethane, unreacted olefins and light impurities in the tower top line 184 can be transported to the combustion unit 186 to generate steam in the pipeline 187. Alternatively, the light raffinate stream in the tower top line 184 can be used in other ways, such as combustion, to ignite a heater (not shown), and / or provide a flue gas source to turn a gas turbine (not shown) to generate electricity, and / or for producing additional hydrogen in another reforming reactor (not shown). The tower top line 184 is connected to the combustion unit 186 upstream. Line 185 contains C 4+ The liquid bottoms stream of olefins, preferably C4 and C6 olefins, may be further used to produce additional propylene.

[0085] In an exemplary embodiment, before the light raffinate stream in line 184 is sent to the combustion unit 186, the light raffinate stream in the overhead line 184 can be passed through a pressure swing adsorption (PSA) unit 194 to remove or recover hydrogen. The molecules heavier than hydrogen present in the light raffinate stream in line 184 are adsorbed on the adsorbent, while the hydrogen molecules are not adsorbed on the adsorbent and pass through the bed. The adsorbent can be selected from one or more of silica gel, alumina, activated carbon, molecular sieves (including but not limited to zeolites), metal organic frameworks (MOFs), etc. or mixtures thereof. PSA 194 may include a single layer or multiple layers of adsorbent beds of different compositions. Light molecules such as hydrogen leave PSA 194 at high pressure in line 195 and are recovered. The adsorbed heavy molecules are then desorbed from PSA 194 at a lower pressure and recovered in the form of a PSA tail gas stream in line 193. The PSA tail gas stream in line 193 may be conveyed to the combustion unit 186 to generate steam in line 187 .

[0086] In one embodiment, PSA 194 may be operated at an adsorption pressure of 1500 kPa (220 psi) to 3500 kPa (515 psi) and a desorption pressure of 100 kPa (15 psi) to 300 kPa (44 psi).

[0087] In an alternative embodiment, before the light raffinate stream in line 184 is sent to the combustion unit 186, the light raffinate stream in the tower top line 184 can be passed through a membrane unit 194 to recover or remove hydrogen. In the membrane unit 194, the light raffinate stream in the tower top line 184 is passed through a membrane at high pressure. In an exemplary embodiment, the membrane unit 194 may include a polyimide or polyethersulfone blend membrane. Other suitable membranes may include polymers, porous ceramics, dense ceramics or metal membranes. After contact, most of the molecules heavier than hydrogen present in the light raffinate stream in the tower top line 184 are blocked by the membrane as a retentate. The heavy molecules leave the membrane unit 194 at high pressure and are discharged from the membrane unit 194 in the form of a retentate stream in the pipeline 193. The retentate stream in the pipeline 193 can be transported to the combustion unit 186 to generate steam in the pipeline 187. The hydrogen present in the light raffinate stream in overhead line 184 permeates through the membrane with a higher selectivity than the retained molecules and is recovered on the other side of the membrane at a lower pressure in the hydrogen stream in line 195 .

[0088] The membrane unit 194 can be operated at a temperature of 40°C (104°F) to 80°C (176°F). The transmembrane pressure difference can be as low as 70 kPa (10 psi) or as high as 14.5 MPa (2100 psi), depending on many factors such as the specific membrane used, the flow rate of the inlet stream, and the availability of a compressor to compress the permeate stream (if such compression is desired). In one embodiment, the transmembrane pressure difference can range between 446 kPa (50 psig) and 6996 kPa (1000 psig) feed pressure.

[0089] Referring again to oligomerization separation unit 182, the C-containing 4+ The oligomerization liquid bottoms stream of olefins is combined with one or more streams to be recycled to the second FCC reactor 202 as a second hydrocarbon stream in the second loading line 150. The second hydrocarbon stream may be preheated to a temperature of 221°C (400°F) to 704°C (1300°F) and charged to the second FCC reactor 202.

[0090] Figure 1 A second FCC reactor 200 including a second reactor unit 202 is shown, wherein the second hydrocarbon stream in line 150 may be loaded through a distributor 216 or at a location closer to the base of a second riser 220 and contacted with a second fluid catalyst stream entering the second riser 220 from a regenerated catalyst riser 218 and a recirculated catalyst riser 219. The second hydrocarbon stream may contain at least 20 wt% olefins, suitably at least 60 wt% olefins, and preferably at least 70 wt% olefins. The second hydrocarbon stream may contain at least 1 wt% paraffins, suitably at least 15 wt% paraffins, and preferably at least 25 wt% paraffins. The second hydrocarbon stream may contain a primary cracked C4 to C7 hydrocarbon. The second riser 220 and the second reactor unit 202 may be connected upstream and downstream to the first riser 20 and the secondary absorber 118.

[0091] This contact may occur in a narrow second riser 220 that extends upward to the bottom of the second reactor vessel 222. The contact of the second hydrocarbon feedstock with the second fluid catalyst stream is fluidized by a gas, such as steam, from a fluidizing distributor 224. In one embodiment, heat from the catalyst vaporizes the second hydrocarbon feedstock, which is then cracked in the presence of the second catalyst stream into a second cracked product stream of lighter molecular weight as both are transferred upward along the riser 220 into the second reactor vessel 222, thereby providing a second mixture of catalyst and product gas.

[0092] The pressure in the second riser 220 can be 200 kPa (29 psia) to 450 kPa (65 psia), but can also be lower. The steam flow rate added to the second riser 220 is 3% to 7% by weight of the second hydrocarbon feedstock. Inevitable side reactions occur in the second riser 220, leaving coke deposits on the catalyst, thereby reducing the catalyst activity and producing a spent catalyst stream. The second cracked product stream in the second mixture of catalyst and product gas is then separated from the spent catalyst stream using a cyclone separator, which may include a single or two-stage cyclone separator 262 in the second reactor vessel 222. The gaseous second cracked product stream leaves the reactor vessel 222 through the second product outlet 231 and enters the pipeline 232 to be transported to the downstream product recovery section 90.

[0093] The second fluid catalyst stream may comprise less than 20 wt%, preferably less than 5 wt%, of the first component and at least 20 wt% of the second component. In a preferred embodiment, the second fluid catalyst stream may comprise at least 20 wt% of ZSM-5 zeolite and less than 20 wt%, preferably less than 5 wt%, of Y-zeolite. In another preferred embodiment, the second fluid catalyst stream may comprise primarily the second component, and in another embodiment, may comprise only the second component, preferably ZSM-5 zeolite, as a catalyst.

[0094] The process conditions in the second riser 212 will be more severe than in the first riser 20 because the second hydrocarbon stream has a lower average molecular weight and a higher olefin content, making it more selective for cracking to propylene under more severe conditions than the first hydrocarbon stream, but producing lower levels of byproducts such as coke. Olefins in the second hydrocarbon stream that cannot be directly cracked to form propylene, such as C4 olefins, can first undergo oligomerization and conversion to suitable heavier olefins and then be re-cracked to propylene in the second riser under process conditions. The second riser 220 can be operated under one or more of the following conditions relative to the first riser 20: a higher outlet temperature, a lower hydrocarbon partial pressure, or a different catalyst density. The hydrocarbon partial pressure is reduced by reducing the total pressure in the second riser 220 independently of the pressure in the first FCC reactor 10 and possibly adjusting the steam rate flowing to the second riser 212.

[0095] The conditions in the second riser 220 may include a cracking reaction temperature at the reactor outlet of 400° C. to 650° C., preferably 565° C. to 635° C. Cracking is carried out at an absolute pressure between 100 kPa (14 psia) and 506 kPa (74 psia), preferably between 138 kPa (20 psia) and 310 kPa (45 psia). The steam flow rate added to the second riser 220 is 5% to 25% by weight of the second hydrocarbon stream.

[0096] Spent or coked catalyst requires regeneration for further use. Figure 1 A single regenerator 14 is shown, but it is also contemplated that two separate regenerators may be used. The spent catalyst stream, after being separated from the second cracked product stream by the disengagement device 254 in the second disengagement chamber 256, falls into the stripping section 234, where steam is injected through the distributor 235 to remove any residual hydrocarbon vapor. The transport conduit 260 transfers the hydrocarbon vapor (including the stripped hydrocarbons, the stripping medium and the entrained catalyst) to a single or double cyclone separator 262 in the second reactor vessel 222, which separates the coked catalyst from the hydrocarbon vapor stream. After the stripping operation, the stripped coked catalyst is carried to the catalyst regenerator 14 via the spent catalyst riser 236. Another portion of the stripped coked catalyst can be recycled to the second riser 220 via the recycle catalyst riser 219 without undergoing regeneration.

[0097] Line 185 contains C 4+ All or part of the liquid bottoms stream of olefins, preferably C4 and C6 olefins, may be fed to a catalytic reaction unit 190 via line 189 controlled by a valve thereon. The catalytic reaction unit 190 may receive a second C4 hydrocarbon stream via line 179 controlled by a valve thereon. Additional feed to the catalytic reaction unit 190 may be fed via line 191. The catalytic reaction unit 190 may catalytically convert the feed stream into a conversion product stream in line 196. The conversion product stream in line 196 may be recycled to the second FCC reactor 202 as at least a portion of the second hydrocarbon stream via line 198 in the second loading line 150. All or part of the conversion product stream may be conveyed to the product recovery section 90 via line 197 controlled by a valve thereon for further processing.

[0098] In one aspect, the catalytic reaction unit 190 may include a cracking process. An exemplary cracking process is an olefin cracking process (OCP) available from UOP LLC. The catalyst used in the cracking process in the catalytic reaction unit 190 may include an MFI zeolite. The MFI zeolite may have a silica-alumina ratio between 400 and 1200, or between 600 and 1100. The catalyst may also include a binder, which may include alumina. The catalytic reaction temperature in the catalytic reaction unit 190 may be between 500°C and 650°C. The catalytic reaction pressure may be between 1 psig and 20 psig, and the catalytic reaction unit may be operated at 1 h. -1 With 20h -1 Between or 5h -1 with 15h -1 Operate under WHSV between.

[0099] In one aspect, the catalytic reaction unit 190 may include an oligomerization process. An exemplary oligomerization process is the Catolene process available from UOP LLC. The catalyst used in the cracking process in the catalytic reaction unit 190 may include solid phosphoric acid (SPA). SPA catalyst refers to a solid catalyst with phosphoric acid (such as orthophosphoric acid, pyrophosphoric acid or tetraphosphoric acid) as the main component. SPA catalysts are usually formed by mixing phosphoric acid with a silicic acid solid carrier to form a wet paste. The paste can be calcined and then crushed to produce catalyst particles, or the paste can be extruded or granulated before calcining to produce more uniform catalyst particles. The carrier is preferably a naturally occurring porous silica-containing material, such as kieselguhr, kaolin, infusorial earth and diatomaceous earth. A small amount of various additives such as mineral talc, bleaching earth and iron compounds including iron oxide can be added to the carrier to improve its strength and hardness. The combination of the carrier and the additive preferably accounts for 15% to 30% by weight of the catalyst, and the rest is phosphoric acid. The additive may comprise 3% to 20% by weight of the total support material. Variations of the composition, such as lower phosphoric acid content, may also be used. Further details on the composition and production of the SPA catalyst may be obtained from U.S. Pat. No. 3,050,472, U.S. Pat. No. 3,050,473, and U.S. Pat. No. 3,132,109, as well as from other references. The catalytic reaction temperature in the catalytic reaction unit 190 may be between 100° C. and 300° C. or between 115° C. and 225° C. The catalytic reaction pressure may be between 100 psig and 1000 psig or between 400 psig and 700 psig, and the catalytic reaction unit may be operated at 0.5 h. -1 With 5h -1 Between or 1h -1 With 2h -1 Operate at LHSV between.

[0100] In one aspect, the catalytic reaction unit 190 may include a metathesis process. The metathesis process converts two olefins into a third olefin. The catalyst used in the metathesis process in the catalytic reaction unit 190 may include tungsten impregnated on a silica support. It may be necessary to pretreat the metathesis catalyst before the reaction. The pretreatment conditions may include reduction in hydrogen at a temperature of 500°C, followed by treatment with ethylene at a temperature of 250°C. Ethylene may be fed via an alternative feed line 191. When the catalytic reaction unit 190 includes metathesis, a feed stream rich in butenes is preferred, such as the second C4 hydrocarbon stream in line 147. The catalytic reaction temperature in the catalytic reaction unit 190 may be between 300°C and 400°C. The catalytic reaction pressure may be between 300 and 700 psig, and the catalytic reaction unit may be operated at 0.1 h-1 With 10h -1 Between or 0.5h -1 With 3h -1 Operate at LHSV between.

[0101] Example

[0102] The following examples are provided to illustrate one or more preferred embodiments of the present disclosure, but are not limited to the embodiments thereof.Many variations may be made to the following examples that fall within the scope of the present disclosure.

[0103] Example 1

[0104] The nickel on amorphous silica-alumina extrudate catalyst was prepared by the following steps:

[0105] The mixed mass was extruded through a 1.59 mm opening in a cylindrical die plate and broken into pieces, then calcined at 550°C to form a silica-alumina base. The mass was prepared by combining amorphous silica-alumina having a silica to alumina ratio of 2.6 supplied by CCIC and pseudo-boehmite supplied under the Catapal trademark in a weight ratio of 85:15. The pseudo-boehmite was peptized with nitric acid prior to mixing with the amorphous silica-alumina. A surfactant supplied under the Antarox trademark and water in an amount sufficient to wet the mass were also added to the mixture.

[0106] 3.37 g of Ni(NO3)2·6H2O was then dissolved in 32.08 g of deionized water and the nickel solution was added to the amorphous silica-alumina base in four portions with vigorous shaking between additions to introduce the nickel into the base. The impregnated base was then dried at 110°C for 3 hours and then calcined at 500°C for 3 hours to convert the nickel to the oxide form. The final catalyst A contained 1.63 wt% nickel.

[0107] Example 2

[0108] A catalyst of nickel on an amorphous silica-alumina oil droplet sphere substrate was synthesized by the procedure given above as an alternative catalyst. After forming an amorphous silica-alumina oil droplet sphere substrate having an average diameter of 3.175 mm and a silica to alumina ratio of 3, nickel was introduced into the substrate in the same manner as described in Example 1, except that the solution concentration and volume were adjusted to match the oil droplet sphere substrate and a metal level of 0.44 wt% nickel was achieved on the final catalyst B.

[0109] Example 3

[0110] A sodium modified nickel on amorphous silica-alumina extrudate catalyst was prepared using the same calcined silica-alumina extrudates as described in Example 1, but using excess solution impregnation for metal loading:

[0111] A nickel solution was prepared by dissolving 2.98 g of Ni(NO3)2·6H2O in 124 g of deionized water. The calcined extruded base from Example 1 was contacted with the nickel solution in an evaporation flask, and the solution-base mixture was rolled at 50°C for 2 hours and then at 100°C for another 2 hours to evaporate the water.

[0112] The nickel impregnated substrate was then further dried at 100°C for 3 hours and calcined at 500°C for 4 hours before sodium impregnation following the same procedure as for nickel impregnation, except that the solution was prepared by dissolving 6.84 g of NaNO3 in 131 g of deionized water. The sodium and nickel impregnated substrates were again dried at 100°C for 3 hours and calcined at 500°C for 4 hours. The final catalyst C contained 1 wt% Ni and 3 wt% Na.

[0113] Example 4

[0114] Catalyst A was tested for olefin oligomerization on 10 ml of catalyst at 300° C., 1000 psig and 4090 GHSV (olefin gas hourly space velocity) in fixed bed operation. The feed consisted of 30.6 wt % ethylene, 20.0 wt % ethane, 26.2 wt % methane, 1.25 wt % hydrogen, 17.5 wt % nitrogen, 1.3 wt % carbon monoxide, 3.2 wt % carbon dioxide, 17.5 wppm hydrogen sulfide, 10 wppm ammonia and saturation level of water. Online GC analysis could not distinguish between olefin and paraffin selectivity for C9+ products because they were too heavy and constituted the majority of the products using Catalyst A.

[0115] Catalyst B was tested for olefin oligomerization over 27.5 ml of catalyst in fixed bed operation at 299° C., 900 psig and 2020 GHSV. The feed consisted of 46.2 wt % ethylene, 22.9 wt % ethane, 25.4 wt % methane, 3.0 wt % hydrogen, 1.0 wt % carbon monoxide, 1.5 wt % carbon dioxide, 37 wppm hydrogen sulfide and 108 wppm acetylene.

[0116] Catalyst C was tested for olefin oligomerization over 32.0 ml of catalyst in fixed bed operation at 285°C, 900 psig and 1730 GHSV. The feed was the same as described above for the test of Catalyst B. The results of these tests are listed in Table 1.

[0117] Table 1: Olefin oligomerization product distribution

[0118]

[0119] The results show that the right combination of catalyst and reaction conditions can increase the C4-C8 hydrocarbon product selectivity and product olefin levels of ethylene oligomerization in the presence of contaminants. When the oligomers are sent to fluid catalytic cracking, higher levels of C4-C8 olefin products are expected to increase the production of propylene.

[0120] Specific implementation plan

[0121] While the following is described in conjunction with specific embodiments, it should be understood that this description is intended to illustrate and not to limit the scope of the foregoing description and the appended claims.

[0122] The first embodiment of the present disclosure is a method for producing propylene, comprising contacting a hydrocarbon feed stream with a fluidized catalyst to produce a cracked product stream; separating the cracked product stream in a main tower; separating the overhead stream from the main tower into a dilute ethylene stream; contacting the dilute ethylene stream with an olefin oligomerization catalyst; converting ethylene in the dilute ethylene stream into an oligomerization product stream; and contacting at least a portion of the oligomerization product stream with a fluidized catalyst to convert at least a portion of the oligomerization product stream into propylene. The embodiment of the present disclosure is one, any one, or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, and further comprises separating the oligomerization product stream to obtain a C4+ olefin stream and a light raffinate stream, and contacting the C4+ olefin stream with a fluidized catalyst to convert at least a portion of the C4+ olefin into propylene. The embodiment of the present disclosure is one, any one, or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the step of contacting the hydrocarbon feed stream with a fluidized catalyst is performed in a first riser under first process conditions, and the step of contacting the C4+ olefin stream with a fluidized catalyst is performed in a second riser under second process conditions. The embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the second process conditions are more severe than the first process conditions. The embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the generation of the cracked product stream produces a spent catalyst stream. The embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, further comprising regenerating the spent catalyst stream by burning the coke in the spent catalyst to provide a regenerated catalyst. The embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the combustion step is carried out at a temperature between 670°C and 750°C. The embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the olefin oligomerization catalyst comprises an amorphous silica-alumina base material and one or more metals selected from Groups 6, 8, 9 and 10 of the periodic table. Embodiments of the present disclosure are one, any, or all of the prior embodiments in this paragraph through the first embodiment in this paragraph, wherein the olefin oligomerization catalyst further comprises one or more metals selected from Groups 1 and 2 of the Periodic Table of the Elements. Embodiments of the present disclosure are one, any, or all of the prior embodiments in this paragraph through the first embodiment in this paragraph, wherein the ethylene in the dilute ethylene stream is converted to C4-C8 hydrocarbons with a selectivity of at least 60 wt%.The embodiment of the present disclosure is one, any one or all of the aforementioned embodiments in this paragraph to the first embodiment in this paragraph, and further includes separating the top stream from the main tower into a second hydrocarbon stream. The embodiment of the present disclosure is one, any one or all of the aforementioned embodiments in this paragraph to the first embodiment in this paragraph, and further includes contacting the light raffinate stream with an adsorbent in a pressure swing adsorption (PSA) unit to remove hydrogen from the light raffinate stream and provide a PSA tail gas stream containing ethylene, and contacting the PSA tail gas stream with the olefin oligomerization catalyst. The embodiment of the present disclosure is one, any one or all of the aforementioned embodiments in this paragraph to the first embodiment in this paragraph, and further includes contacting the dilute ethylene stream with an adsorbent in a pressure swing adsorption (PSA) unit to recover hydrogen from the dilute ethylene stream and provide a PSA tail gas stream containing unreacted olefins.

[0123] The second embodiment of the present disclosure is a method for producing propylene, comprising contacting a hydrocarbon feed stream with a fluidized catalyst to produce a cracked product stream; separating the cracked product stream in a main tower; separating the top stream from the main tower into a dilute ethylene stream; separating the top stream from the main tower into a second hydrocarbon stream containing C4 olefins; contacting the dilute ethylene stream with an olefin oligomerization catalyst; converting ethylene in the dilute ethylene stream into an oligomerization product stream; separating the oligomerization product stream to obtain a C4+ olefin stream and a light raffinate stream; reacting at least a portion of the C4+ olefin stream in a catalytic reaction unit to provide a conversion product stream; and contacting at least a portion of the conversion product stream with a fluidized catalyst to convert at least a portion of the C4+ olefins into propylene. The embodiments of the present disclosure are one, any one, or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the step of contacting the hydrocarbon feed stream with a fluidized catalyst is performed in a first riser under first process conditions. The embodiment of the present disclosure is one, any one or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the step of contacting the conversion product stream with the fluidized catalyst is performed in the second riser under the second process conditions. The embodiment of the present disclosure is one, any one or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the second process conditions are more severe than the first process conditions. The embodiment of the present disclosure is one, any one or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the generation of the cracked product stream produces a spent catalyst stream. The embodiment of the present disclosure is one, any one or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, further comprising regenerating the spent catalyst stream by burning the coke in the spent catalyst to provide a regenerated catalyst. The embodiment of the present disclosure is one, any one or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the combustion step is performed at a temperature between 670°C and 750°C. Embodiments of the present disclosure are one, any one, or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the olefin oligomerization catalyst comprises an amorphous silica-alumina base material and one or more metals selected from Groups 6, 8, 9, and 10 of the periodic table. Embodiments of the present disclosure are one, any one, or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the ethylene in the dilute ethylene stream is converted to C4-C8 hydrocarbons with a selectivity of at least 60 weight %. Embodiments of the present disclosure are one, any one, or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the catalytic reaction unit comprises a cracking process.The embodiment of the present disclosure is one, any one or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the catalytic reaction unit includes oligomerization. The embodiment of the present disclosure is one, any one or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the catalytic reaction unit includes metathesis, and further processes at least a portion of the conversion product stream to recover a propylene stream. The embodiment of the present disclosure is one, any one or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, further comprising contacting the second hydrocarbon stream with the olefin oligomerization catalyst. The embodiment of the present disclosure is one, any one or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the diluted ethylene stream comprises between 10 wt% and 65 wt% ethylene, between 20 wt% and 55 wt% methane and at least one impurity selected from the following: at least 0.1 wt% carbon monoxide, at least 1wppm hydrogen sulfide, at least 1wppm ammonia, at least 5 wt% hydrogen and at least 0.1 wt% carbon dioxide.

[0124] A third embodiment of the present disclosure is a method for producing propylene, comprising contacting a hydrocarbon feed stream with a fluidized catalyst to produce a cracked product stream; separating the cracked product stream in a main tower; separating the overhead stream from the main tower into a dilute ethylene stream, the dilute ethylene stream comprising between 10 wt% and 65 wt% ethylene, between 20 wt% and 55 wt% methane, and at least one impurity selected from the group consisting of at least 0.1 wt% carbon monoxide, at least 1 wppm hydrogen sulfide, at least 1 wppm ammonia, at least 0.5 wt% hydrogen, and at least 0.1 wt% % carbon dioxide; separating the top stream from the main tower into a second hydrocarbon stream comprising C4 olefins; contacting the dilute ethylene stream and the second hydrocarbon stream with an olefin oligomerization catalyst; converting the ethylene in the dilute ethylene stream into an oligomerization product stream; separating the oligomerization product stream to obtain a C4+ olefin stream and a light raffinate stream; reacting at least a portion of the C4+ olefin stream in a catalytic reaction unit to provide a conversion product stream; and contacting at least a portion of the conversion product stream with a fluidized catalyst to convert at least a portion of the C4+ olefins into propylene. An embodiment of the present disclosure is one, any one, or all of the previous embodiments in this paragraph to the third embodiment in this paragraph, wherein the step of contacting the hydrocarbon feed stream with a fluidized catalyst is performed in a first riser under first process conditions. An embodiment of the present disclosure is one, any one, or all of the previous embodiments in this paragraph to the third embodiment in this paragraph, wherein the step of contacting the conversion product stream with a fluidized catalyst is performed in a second riser under second process conditions. The embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the second process condition is more severe than the first process condition. The embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the third embodiment in this paragraph, wherein the generation of the cracked product stream produces a spent catalyst stream. The embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the third embodiment in this paragraph, and also include regenerating the spent catalyst stream by burning the coke in the spent catalyst to provide a regenerated catalyst. The embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the third embodiment in this paragraph, wherein the combustion step is carried out at a temperature between 670°C and 750°C. The embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the third embodiment in this paragraph, wherein the olefin oligomerization catalyst comprises an amorphous silica-alumina base material and one or more metals selected from Groups 6, 8, 9 and 10 of the periodic table.Embodiments of the present disclosure are one, any one, or all of the preceding embodiments in this paragraph to the third embodiment in this paragraph, wherein the ethylene in the diluted ethylene stream is converted to C4-C8 hydrocarbons with a selectivity of at least 60 wt%. Embodiments of the present disclosure are one, any one, or all of the preceding embodiments in this paragraph to the third embodiment in this paragraph, wherein the catalytic reaction unit comprises a cracking process. Embodiments of the present disclosure are one, any one, or all of the preceding embodiments in this paragraph to the third embodiment in this paragraph, wherein the catalytic reaction unit comprises oligomerization. Embodiments of the present disclosure are one, any one, or all of the preceding embodiments in this paragraph to the third embodiment in this paragraph, wherein the catalytic reaction unit comprises metathesis, and further processes at least a portion of the conversion product stream to recover a propylene stream.

[0125] A fourth embodiment of the present disclosure is a method for producing propylene, comprising contacting a hydrocarbon feed stream with a fluidized catalyst in a first riser under first process conditions to produce a first cracked product stream and a spent catalyst stream; separating the cracked product stream in a main tower; separating an overhead stream from the main tower into a dilute ethylene stream, the dilute ethylene stream comprising between 10 wt% and 65 wt% ethylene, between 20 wt% and 55 wt% methane, and at least one impurity selected from the group consisting of at least 0.1 wt% carbon monoxide, at least 1 wppm hydrogen sulfide, at least 1 wppm ammonia, at least 0.5 wt% hydrogen, and at least 0.1 wt% carbon dioxide; separating The top stream from the main tower is separated into a second hydrocarbon stream containing C4 olefins; the dilute ethylene stream and the second hydrocarbon stream are contacted with an olefin oligomerization catalyst; ethylene in the dilute ethylene stream is converted into an oligomerization product stream; the oligomerization product stream is separated to obtain a C4+ olefin stream and a light raffinate stream; at least a portion of the C4+ olefin stream is reacted in a catalytic reaction unit to provide a conversion product stream; and at least a portion of the conversion product stream is contacted with a fluidized catalyst in a second riser under second process conditions to convert at least a portion of the C4+ olefins into propylene, wherein the second process conditions are more severe than the first process conditions.

[0126] Although there is no further detailed description, it is believed that those skilled in the art can make the most of the present disclosure by using the foregoing description and can easily determine the essential characteristics of the present disclosure without departing from the spirit and scope of the present invention, and can make various changes and modifications of the present disclosure and adapt it to various usages and conditions. Therefore, the aforementioned preferred specific embodiments should be understood as merely illustrative and not to limit the rest of the present disclosure in any way, and are intended to cover various modifications and equivalent arrangements included within the scope of the appended claims.

[0127] In the foregoing, all temperatures are set forth in degrees Celsius and all parts and percentages are by weight unless otherwise indicated.

Claims

1. A method for producing propylene, comprising: contacting a hydrocarbon feed stream with a fluidized catalyst to produce a cracked product stream; separating the cracked product stream in a main column; separating an overhead stream from the main column into a dilute ethylene stream; contacting the dilute ethylene stream with an olefin oligomerization catalyst; converting ethylene in the dilute ethylene stream into an oligomerization product stream; as well as At least a portion of the oligomerization product stream is contacted with a fluidized catalyst to convert at least a portion of the oligomerization product stream to propylene.

2. The method according to claim 1, further comprising separating the oligomerization product stream to obtain C4 + olefin stream and light raffinate stream, and making the C4 + The olefin stream is contacted with a fluidized catalyst to convert the C4 + At least a portion of the olefins are converted to propylene.

3. The process of claim 1, wherein the step of contacting the hydrocarbon feed stream with a fluidized catalyst is performed in a first riser under first process conditions and the C4 + The step of contacting the olefin stream with the fluidized catalyst is performed in the second riser under second process conditions. The method according to claim 3 , wherein the second process condition is more severe than the first process condition.

5. The process of claim 1, wherein the olefin oligomerization catalyst comprises an amorphous silica-alumina base material and one or more metals selected from Groups 6, 8, 9 and 10 of the Periodic Table of Elements.

6. The process of claim 5, wherein the olefin oligomerization catalyst further comprises one or more metals selected from Groups 1 and 2 of the Periodic Table of Elements.

7. The process of claim 1, wherein the ethylene in the dilute ethylene stream is converted to C4-C8 hydrocarbons with a selectivity of at least 60 wt%.

8. The method of claim 1, further comprising separating an overhead stream from the main column into a second hydrocarbon stream.

9. The method according to claim 2, further comprising: contacting the light raffinate stream with an adsorbent in a pressure swing adsorption (PSA) unit to remove hydrogen from the light raffinate stream and provide a PSA tail gas stream comprising ethylene; and The PSA tail gas stream is contacted with the olefin oligomerization catalyst.

10. The process of claim 1 further comprising contacting the dilute ethylene stream with an adsorbent in a pressure swing adsorption (PSA) unit to recover hydrogen from the dilute ethylene stream and provide a PSA tail gas stream comprising unreacted olefins.

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