Process for converting olefins to jet fuel

Abstract

A process for dimerizing and oligomerizing olefins into distillate fuels, where the oligomerized product is subjected to hydrocracking and optional hydroisomerization to convert heavier oligomers to jet fuel range oligomers. A jet fuel product stream is removed from the side of a stripping tower, and a heavy drag stream is produced from the bottom of the stripping tower to ensure that the end boiling point specification is met.

Description

[Technical Field]

[0001] (Statement of priority) This application claims priority to U.S. Provisional Patent Application No. 63 / 477,395, filed December 28, 2022, which is incorporated herein by reference in its entirety.

[0002] FIELD OF THE INVENTION The field is the conversion of olefins to distillates. The field may particularly relate to the dimerization of olefins and the oligomerization of dimerized olefins to distillate fuels. [Background technology]

[0003] Ethylene can be dimerized and oligomerized to olefins, such as C4, C6, and C8 olefins. Olefin oligomerization is a process by which smaller olefins can be oligomerized to larger olefins. More specifically, olefins, including dimerized olefins, can be converted to distillates, including jet fuel and diesel range products. The oligomerized distillates can be saturated for use as transportation fuels.

[0004] The dimerization reaction of ethylene is highly exothermic. The exotherm generated by the dimerization of ethylene can be difficult to manage.

[0005] Jet fuel is one of the few petroleum fuels that cannot be easily substituted in electric motor systems because fueling an airplane requires a high energy output that cannot be provided by an electric motor. Jet fuel has an end boiling point specification of less than 300 °C using ASTM D86. Significant tax incentives are currently available for green jet fuel in certain regions.

[0006] Hydroprocessing can involve the conversion of hydrocarbons into more valuable products in the presence of a hydroprocessing catalyst and hydrogen. Hydrotreating or hydrogenation is the process of contacting hydrogen with hydrocarbons in the presence of a hydrotreating catalyst that is primarily active for the removal of heteroatoms, such as sulfur, nitrogen, oxygen, and metals, from the hydrocarbon feed. Hydrotreating can saturate hydrocarbons with double and triple bonds, such as olefins.

[0007] Hydrocracking is a hydroprocessing process in which hydrocarbons are cracked into lower molecular weight hydrocarbons in the presence of hydrogen and a hydrocracking catalyst. Depending on the desired production output, the hydrocracking unit may contain one or more beds of the same or different catalysts. Hydroisomerization is a hydroprocessing process that increases the number of alkyl groups on the hydrocarbon chain.

[0008] An efficient process for converting ethylene to distillate fuels is desirable. Summary of the Invention

[0009] The inventors have devised a process for dimerizing and oligomerizing olefins into distillate fuels that cracks larger diesel-range oligomers into jet fuel-range molecules. An isomerization function can also be performed during cracking to improve jet fuel properties and bring the larger oligomers into the jet fuel boiling range. A jet fuel product stream is removed from the side of the stripping tower, and a heavy drag stream is produced from the bottom of the stripping tower to ensure that the end-boiling point specification is met. [Brief explanation of the drawings]

[0010] [Figure 1] FIG. 1 is a schematic diagram of the oligomerization section of the process and apparatus of the present disclosure. [Figure 2] FIG. 1 is a schematic diagram of the hydrogenation section of the process and apparatus of the present disclosure.

[0011] definition The term "communication" means that fluid flow is operably permitted between the listed components, which may be characterized as "fluid communication."

[0012] The term "downstream communication" means that at least a portion of the fluid flowing to an object in downstream communication can operatively flow from the object in fluid communication.

[0013] The term "upstream communication" means that at least a portion of the fluid flowing from the object with which it is in upstream communication can operatively flow to the object with which it is in fluid communication.

[0014] The term "direct communication" means that fluid flow from an upstream component enters a downstream component without passing through any other intervening vessel.

[0015] The term "indirect communication" means that fluid flow from an upstream component passes through an intervening vessel before entering a downstream component.

[0016] The term "bypass" means that an object is out of downstream communication with a bypass subject, at least to the extent that it bypasses.

[0017] As used herein, the term "predominant" or "majority" means more than 50%, suitably more than 75%, preferably more than 90%.

[0018] The term "column" refers to a distillation column or columns for separating one or more components of different volatility. Unless otherwise indicated, each column includes a condenser at the top of the column for condensing and refluxing a portion of the overhead stream returning to the top of the column, and a reboiler at the bottom of the column for vaporizing a portion of the bottom stream and returning it to the bottom of the column. The feed to a column may be preheated. The overhead pressure is the pressure of the overhead vapor at the vapor outlet of the column. The bottom temperature is the liquid bottom outlet temperature. The overhead and bottom lines refer to the net lines from column to column downstream of any reflux or reboil. A stripping column may omit the reboiler at the bottom of the column and instead provide the heating requirements and separation driving force from a fluidized inert medium such as steam. A stripping column typically feeds a top tray and removes the main product from the bottom.

[0019] As used herein, the term "separator" means a vessel having an inlet and at least a top vapor outlet and a bottom liquid outlet, and may also have an aqueous outlet from a boot. A flash drum is a type of separator that may be in downstream communication with a separator that may be operated at a higher pressure. As used herein, the term "boiling point temperature" means the atmospheric equivalent boiling point (AEBP) calculated from the observed boiling point and distillation pressure, calculated using the formula provided in ASTM D1160 Appendix A7, entitled "Practice for Converting Observed Vapor Temperatures to Atmospheric Equivalent Temperatures."

[0020] As used herein, the term "True Boiling Point" (TBP) refers to a test method for determining the boiling point of a substance, which test method corresponds to ASTM D-2892 for producing liquefied gases, distillate fractions, and residual oils of standardized quality for which analytical data can be obtained, and for determining the yield of said fractions by both mass and volume, where a graph of temperature versus mass % distilled is produced using 15 theoretical plates in a column with a reflux ratio of 5:1.

[0021] As used herein, the terms "T5," "T90," or "T95" refer to the temperature at which 5 percent, 90 percent, or sometimes 95 percent by weight of a sample boils using ASTM D-86 or TBP, respectively.

[0022] 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 the case may be.

[0023] As used herein, the term "end point" (EP) means the temperature at which the sample has completely evaporated using ASTM D-7169, ASTM D-86, or TBP, as the case may be.

[0024] As used herein, the term "diesel" refers to hydrocarbons boiling within a range of "diesel cut points" including an IBP of 125°C (257°F) to 175°C (347°F) or a T5 of 150°C (302°F) to 200°C (392°F), and a T95 of 343°C (650°F) to 399°C (750°F) using the TBP distillation method or a T90 of 280°C (536°F) to 340°C (644°F) using ASTM D-86. The term "green diesel" refers to diesel containing hydrocarbons not derived from fossil fuels.

[0025] As used herein, the term "jet fuel" means hydrocarbons that boil in the range of 190°C (374°F) to 215°C (419°F) T10 and 290°C (554°F) to 310°C (590°F). The term "green jet fuel" means jet fuel containing hydrocarbons that do not originate from fossil fuels. DETAILED DESCRIPTION OF THE INVENTION

[0026] The disclosed process involves dimerizing an olefin stream containing ethylene, followed by further oligomerization of the ethylene dimer and ethylene oligomers. The process utilizes a hydrocracking reactor to crack and optionally hydroisomerize the larger oligomers to jet fuel range. The process and apparatus can include an oligomerization section 10 in FIG. 1 and a hydroprocessing section 80 in FIG. 2.

[0027] Referring to oligomerization section 10 of FIG. 1 , an input olefin stream in line 12 is fed to oligomerization section 10. The input olefin stream can comprise substantial ethylene. The input olefin stream can comprise primarily ethylene. In embodiments, the input olefin stream can comprise at least 95 mole % ethylene. The input olefin stream in line 12 may also be referred to as an ethylene stream. The olefin stream may be provided by the dehydration of ethanol or from an MTO unit. The input olefin stream may have a temperature of 60° C. (140° F.) to 150° C. (302° F.), preferably 80° C. (176° F.) to 100° C. (212° F.), and a pressure of 3.5 MPa (500 psig), preferably 5.6 MPa (800 psig) to 8.4 MPa (1200 psig).

[0028] The input olefin stream can be first contacted with a first stage oligomerization catalyst to dimerize ethylene to dimers and oligomerize the dimers to oligomers, and then contacted with a second oligomerization catalyst to dimerize unconverted ethylene and oligomerize the dimerized and oligomerized ethylene from the first stage oligomerization. Alternatively, the olefin stream can be first contacted with a second stage oligomerization catalyst to dimerize ethylene and then contacted with a first stage oligomerization catalyst to oligomerize the dimerized ethylene.

[0029] Dimerization reactions produce large amounts of heat. For example, the dimerization of ethylene can generate 612 kcal / kg (1100 BTU / lb). As a result, this large amount of heat must be managed.

[0030] Thus, the input olefin stream in line 12 can be divided into multiple olefin streams. In FIG. 1 , the olefin stream is divided into four separate streams: a first olefin stream in input line 12a, a second olefin stream in input line 12b, a third olefin stream in input line 12c, and a fourth or final olefin stream in input line 12d. More or fewer separate olefin streams can also be used. Up to six olefin streams are readily contemplated. The input olefin stream in line 12 can be divided into multiple olefin streams of equal amounts. Alternatively, the input olefin stream in line 12 can be divided into unequal streams. For example, the input olefin stream can be divided into streams of increasing flow rates, with the olefin stream to a subsequent reactor having a higher flow rate than the olefin stream to a previous reactor. In an embodiment, the input olefin stream is divided into four streams of equal flow rates, each containing 25% by volume of the input olefin stream. In another embodiment, the input olefin stream is split into two olefin streams: a first olefin stream in input line 12a and a second olefin stream in input line 12b. The first olefin stream may comprise 70-90% by volume of the input olefin stream, and the second olefin stream may comprise 10-30% by volume of the input olefin stream.

[0031] To manage the heat release, the olefin stream can be diluted with a diluent stream to provide a diluent olefin stream for absorbing the heat release. The diluent stream can include a paraffin stream in diluent line 14. The diluent stream in diluent line 14 can be added to the input olefin stream in line 12 before splitting the input olefin stream into multiple olefin streams. Preferably, after splitting into multiple olefin streams, the diluent stream is added to the first olefin stream in line 12a to provide a first diluted olefin stream in line 16a, whereby the diluent stream passes through all of the first-stage oligomerization reactions. Alternatively, the diluent stream can be split into multiple streams, and each diluent stream can be added to a corresponding olefin stream. The diluent stream can have a volumetric flow rate that is 2 to 8 times, preferably 3 to 6 times, the volumetric flow rate of the input olefin stream. The first diluted olefin stream can contain 35 wt.% or less olefins, suitably 30 wt.% or less olefins, and preferably 20 wt.% or less olefins. In one embodiment, the first dilute olefin stream comprises 10 to 30 wt. % C2 to C8 olefins. The first dilute olefin stream may comprise up to 30 wt. % ethylene, suitably up to 25 wt. % ethylene, and preferably up to 20 wt. % ethylene. In one embodiment, the first dilute olefin stream comprises 10 to 20 wt. % ethylene. The first dilute olefin stream in line 16a may be cooled in first feed cooler 18a to provide a first cooled dilute olefin stream in line 20a, which may be combined with an oligomer recycle stream in oligomer recycle line 26 to provide a first-stage oligomerization feed stream in first-stage oligomer feed line 21a, which may be charged to a first bed 22a of first-stage oligomerization catalyst in first-stage oligomerization reactor 22. The first-stage oligomerization feed stream in line 21a may be fed at a temperature of 180°C (356°F) to 260°C (500°F) and a pressure of 3.5MPag (500 psig) to 8.4MPag (1200 psig). Feed cooler 18a may be equipped with a steam generator.

[0032] The first-stage oligomerization reactor 22 can include a series of first-stage oligomerization catalyst beds 22a, 22b, 22c, and 22d for inputting each of the plurality of olefin streams 12a, 12b, 12c, and 12d, respectively. The first-stage oligomerization reactor preferably includes four fixed first-stage oligomerization catalyst beds 22a, 22b, 22c, and 22d. It is also contemplated that each first-stage oligomerization catalyst bed 22a, 22b, 22c, and 22d can be in a dedicated first-stage oligomerization reactor, or that multiple first-stage oligomerization catalyst beds can be in two or more separate first-stage oligomerization reactors. Up to six first-stage oligomerization catalyst beds are readily contemplated. Parallel first stage oligomerization reactors may be used when first stage oligomerization reactor 22 is deactivated while it is regenerated in situ by burning coke from the catalyst. In another embodiment, each first stage oligomerization reactor may include a lead reactor, a lag reactor, and a pre-reactor to facilitate regeneration.

[0033] The first-stage oligomerization input stream may be introduced into the first first-stage catalyst bed 22a in line 21a, preferably in downflow operation. However, upflow operation may also be preferred. As ethylene dimerization occurs in the first first-stage oligomerization catalyst bed 22a, heat is generated due to the exothermic nature of the ethylene dimerization reaction. Dimerization and oligomerization of the first-stage oligomerization input stream, including the first olefin stream, produces a first oligomerized olefin stream in first oligomerization effluent line 24a at an elevated outlet temperature, despite cooling and dilution. The elevated outlet temperature is limited to between 150°C (302°F) and 250°C (482°F).

[0034] The second olefin stream in line 12b can be diluted with the first oligomerized olefin stream in line 24a removed from the first first-stage oligomerization reactor 22 to provide a second diluted olefin stream in line 16b. The first oligomerized olefin stream in line 24a comprises a diluent stream from diluent line 14 added to the first olefin stream in line 12a. The second diluted olefin stream can comprise up to 35 wt% C2-C8 olefins, suitably up to 25 wt% C2-C8 olefins, and preferably up to 20 wt% ethylene. The second diluted olefin stream can comprise up to 30 wt% ethylene, suitably up to 25 wt% ethylene, and preferably up to 20 wt% ethylene. The second dilute olefin stream in line 16b is cooled in a second feed cooler 18b, which may be located external to the dimerization reactor 22, to provide a second cooled dilute olefin stream in line 20b, which may be charged to a second bed 22b of first-stage oligomerization catalyst in the first-stage oligomerization reactor 22. The feed cooler 18b may include a steam generator.

[0035] The second cooled, dilute olefin stream in line 20b may be input at a temperature between 180°C (356°F) and 230°C (446°F) and a pressure between 3.5 and 8.4 MPa (500 and 1200 psig). The second dilute olefin stream comprises diluent and olefins from the first oligomerized olefin stream. The olefins from the first oligomerized olefin stream are dimerized and oligomerized in second catalyst bed 22b. Dimerization and oligomerization of ethylene and oligomers in the second olefin stream in second bed 22b of first-stage oligomerization catalyst produces a second oligomerized olefin stream in second dimerization effluent line 24b at an elevated outlet temperature. The elevated outlet temperature may be limited to between 30°C (54°F) and 50°C (90°F) above the inlet temperature to catalyst bed 22b.

[0036] The third olefin stream in line 12c can be diluted with the second oligomerized olefin stream in line 24b removed from the second first-stage oligomerization reactor 22 to provide a third diluted olefin stream in line 16c. Alternatively, the third olefin stream in line 12c has no flow, and line 16c carries only the second oligomerized olefin stream from line 24b. The second oligomerized olefin stream in line 24b comprises the diluent stream from diluent line 14 added to the first olefin stream in line 12a. The third diluted olefin stream can contain up to 30 wt% ethylene, suitably up to 25 wt% ethylene, and preferably up to 20 wt% ethylene. The third diluted olefin stream can contain up to 30 wt% C2-C8 olefins, suitably up to 25 wt% C2-C8 olefins, and preferably up to 20 wt% C2-C8 olefins. The third dilute olefin stream in line 16c is cooled in a third feed cooler 18c, which may be located external to the dimerization reactor 22, to provide a third cooled dilute olefin stream in line 20c, which may be fed to a third bed 22c of first-stage oligomerization catalyst in the first-stage oligomerization reactor 22. The feed cooler 18c may include a steam generator.

[0037] The third cooled, dilute olefin stream in line 20c may be introduced at a temperature between 180°C (356°F) and 230°C (446°F) and a pressure between 3.5 MPa (500 psig) and 8.4 MPa (1200 psig). The third dilute olefin stream comprises diluent and olefins from the second oligomerized olefin stream. The olefins from the second oligomerized olefin stream are oligomerized in third catalyst bed 22c. Dimerization of ethylene and oligomerization of oligomers in the third olefin stream in third bed 22c of first-stage oligomerization catalyst produces a third oligomerized olefin stream in third oligomerization effluent line 24c at an elevated outlet temperature. In one embodiment, the third oligomerized olefin stream is the penultimate oligomerized olefin stream, and the third oligomerization effluent line 24c is the penultimate oligomerization effluent line 24c. The elevated outlet temperature is limited to between 30°C (54°F) and 50°C (90°F) above the inlet temperature to catalyst bed 22c.

[0038] The fourth olefin stream in line 12d can be diluted with the third or penultimate oligomerized olefin stream in line 24c removed from the first-stage oligomerization reactor 22 to provide a fourth dilution olefin stream in line 16d. The third or penultimate oligomerized olefin stream in line 24c comprises the diluent stream from diluent line 14 added to the first olefin stream in line 12a. The fourth olefin stream in line 12d can be empty, with line 16d carrying only the third oligomerized olefin stream from line 24c. The fourth dilution olefin stream can comprise 35 wt% or less C2-C8 olefins, suitably 30 wt% or less C2-C8 olefins, and preferably 25 wt% or less C2-C8 olefins. The fourth dilution olefin stream can comprise 30 wt% or less ethylene, suitably 25 wt% or less ethylene, and preferably 20 wt% or less ethylene. The fourth dilute olefin stream in line 16d is cooled in a fourth feed cooler 18d, which may be located external to the dimerization reactor 22, to provide a fourth cooled dilute olefin stream in line 20d, which may be fed to a fourth bed 22d of first-stage oligomerization catalyst in the first-stage oligomerization reactor 22. The feed cooler 18d may include a steam generator.

[0039] The fourth cooled dilute olefin stream in line 20d may be introduced at a temperature between 180°C (356°F) and 230°C (446°F) and a pressure between 3.5 MPag (500 psig) and 8.4 MPag (1200 psig). The fourth, or final, dilute olefin stream comprises diluent and olefins from the third, or penultimate, oligomerized olefin stream. The olefins from the third, or penultimate, oligomerized olefin stream are dimerized and oligomerized in fourth catalyst bed 22d. Dimerization of ethylene and oligomerization of olefins in the fourth olefin stream in fourth bed 22d of first-stage oligomerization catalyst produces a fourth oligomerized olefin stream in fourth oligomerization effluent line 24d at an elevated outlet temperature. The elevated outlet temperature is limited to 30°C (54°F) to 50°C (90°F) above the inlet temperature to catalyst bed 22d.

[0040] In one embodiment, the fourth olefin stream is the final olefin stream, the fourth oligomerized olefin stream is the final oligomerized olefin stream, and the fourth oligomerization effluent line 24d is the final oligomerization effluent line 24d.

[0041] The first stage oligomerization reaction takes 0.5 to 10 hours based on olefins. -1 The conversion occurs primarily in the liquid phase or mixed gas-liquid phase at an LHSV of 0.015. The inventors have found that, across each first-stage oligomerization catalyst bed, typically 30-50 wt. % of the ethylene in the olefin stream is converted to higher olefins. Ethylene is first dimerized over the catalyst to butenes. A majority of the butenes in the olefin stream charged to the first-stage oligomerization catalyst bed are oligomerized. In one embodiment, at least 99 mole % of the butenes in the olefin stream are oligomerized.

[0042] The first-stage oligomerization catalyst can include a zeolite catalyst. The first-stage oligomerization catalyst can be considered a solid acid catalyst. The zeolite can comprise 5 to 95 wt. % of the catalyst, e.g., 5 to 85 wt. %. Suitable zeolites include zeolites having a structure from one of the following classes: MFI, MEL, ITH, IMF, TUN, FER, BEA, FAU, BPH, MEI, MSE, MWW, UZM-8, MOR, OFF, MTW, TON, MTT, AFO, ATO, and AEL. Three-letter codes for zeotypes are as defined by the Structure Commission of the International Zeolite Association and maintained at http: / / www.iza-structure.org / databases. UZM-8 is as described in U.S. Patent No. 6,756,030. In a preferred embodiment, the first-stage oligomerization catalyst may comprise a zeolite having a framework with a 10-ring pore structure. Examples of suitable zeolites having a 10-ring pore structure include TON, MTT, MFI, MEL, AFO, AEL, EUO, and FER. In a further preferred embodiment, the first-stage oligomerization catalyst comprising a zeolite having a 10-ring pore structure may comprise a one-dimensional pore structure. A one-dimensional pore structure refers to a zeolite containing non-intersecting pores that are substantially parallel to one of the crystal axes. The pores preferably extend throughout the zeolite crystal. A suitable example of a zeolite having a 10-ring one-dimensional pore structure may include MTT. In a further embodiment, the first-stage oligomerization catalyst comprises an MTT zeolite.

[0043] The first-stage oligomerization catalyst can be formed by combining a zeolite with a binder and then forming the catalyst into pellets. The pellets may optionally be treated with a phosphorus reagent to produce a zeolite having a phosphorus component of 0.5 to 15% by weight of the treated catalyst. The binder is used to impart hardness and strength to the catalyst. Binders include alumina, aluminum phosphate, silica, silica-alumina, zirconia, titania, and combinations of these metal oxides, as well as other refractory oxides and clays such as montmorillonite, kaolin, palygorskite, smectite, and attapulgite. Preferred binders are aluminum-based binders such as alumina, aluminum phosphate, silica-alumina, and clays.

[0044] One of the components of the catalyst binder utilized in the present invention is alumina. The alumina source can be any of a variety of hydrous aluminum oxides or alumina gels, such as alpha-alumina monohydrate with boehmite or pseudo-boehmite structure, alpha-alumina trihydrate with gibbsite structure, or beta-alumina trihydrate with bayerite structure. A suitable alumina is available from UOP LLC under the VERSAL trademark. A preferred alumina is available from Sasol North America Alumina Product Group under the Catapal trademark. This material is an extremely high purity alpha-alumina monohydrate (pseudo-boehmite) that has been shown to yield high purity gamma-alumina after high temperature calcination.

[0045] A suitable first-stage oligomerization catalyst is prepared by mixing proportional volumes of zeolite and alumina to achieve the desired zeolite to alumina ratio. In embodiments, the MTT content may be 5 to 85, e.g., 20 to 82 wt. % MTT zeolite, with the remainder being alumina powder to provide a suitably supported catalyst. Silica supports are also contemplated.

[0046] A monobasic 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 wettability to form a dough with sufficient consistency to be extruded or spray-dried. An extrusion aid such as cellulose ether powder may also be added. A preferred extrusion aid is available from The Dow Chemical Company under the Methocel trademark.

[0047] The paste or dough may be prepared in the form of shaped particles; a preferred method is to extrude the dough through a die having openings of the desired size and shape, followed by dividing the extruded material into extrudates of the desired length and drying. A further calcination step may be used to provide additional strength to the extrudates. Typically, calcination is carried out in a stream of air at a temperature of 260°C (500°F) to 815°C (1500°F). The MTT catalyst is not selective to neutralize acidic sites such as amines.

[0048] The extruded particles can have any suitable cross-sectional shape, i.e., symmetrical or asymmetrical, but most often have a symmetrical cross-sectional shape, preferably spherical, cylindrical, or multilobal. The cross-sectional diameter of the particles can be as small as 40 μm. However, it is usually between 0.635 mm (0.25 inch) and 12.7 mm (0.5 inch), preferably between 0.79 mm (1 / 32 inch) and 6.35 mm (0.25 inch), and most preferably between 0.06 mm (1 / 24 inch) and 4.23 mm (1 / 6 inch).

[0049] In one exemplary embodiment, an MTT-type zeolite catalyst disposed on a high purity pseudoboehmite alumina substrate in a ratio of 90 / 10 to 20 / 80, preferably 20 / 80 to 50 / 50, is provided in a catalyst bed or more in the first stage oligomerization reactor 22.

[0050] The first-stage oligomerization catalyst can be regenerated when deactivated. Suitable regeneration conditions include subjecting the first-stage oligomerization catalyst, for example, in situ, to hot air at 400°C to 500°C for three hours. To facilitate regeneration without downtime, a swing-bed system may be used in conjunction with an alternative first-stage oligomerization reactor. A regeneration gas stream may be introduced into the first-stage oligomerization reactor 22 requiring regeneration. The regeneration gas may comprise air with either increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to that of fresh catalyst.

[0051] Zeolite catalysts are advantageous as first-stage oligomerization catalysts because they have a relatively low susceptibility to oxygenate contamination. As a result, the olefin feed in line 12 requires less oxygenate removal when produced from an ethanol dehydration process.

[0052] The final first-stage oligomerized olefin stream in final first-stage oligomerization effluent line 24d has an increased concentration of ethylene dimer and oligomers compared to the input olefin stream in line 12. The final first-stage oligomerized olefin stream in final first-stage oligomerization effluent line 24d is cooled by steam generation or other heat exchange to provide a second-stage oligomerization input stream, which is fed to second-stage oligomerization reactor 32 in second-stage oligomerization input line 28. To achieve the most desirable olefin products, second-stage oligomerization reactor 32 operates at a temperature of 80°C (176°F) to 180°C (356°F). Second stage oligomerization reactor 32 is operated at a pressure between 2.1 MPa (300 psig) and 7.6 MPa (1100 psig), more preferably between 3.5 MPa (500 psig) and 6.9 MPa (1000 psig).

[0053] The second-stage oligomerization reactor 32 may be in downstream communication with the first-stage oligomerization reactor 22. The second-stage oligomerization reactor 32 preferably operates in downflow operation; however, upflow operation may be preferred. The second-stage oligomerization input stream is contacted with a second-stage oligomerization catalyst to dimerize and trimerize unconverted ethylene from the first-stage oligomerization reactor 22, while also dimerizing, trimerizing, and tetramerizing higher olefins to provide distillate range olefins. For the second-stage oligomerization reactor 32, process conditions are selected to produce a higher proportion of jet range olefins, which, when hydrogenated in subsequent steps as described below, yield desirable jet range hydrocarbon products. The majority of the unconverted ethylene from the first-stage oligomerization reactor 22 is dimerized, trimerized, and tetramerized. In one embodiment, at least 99 wt. % of the ethylene in the second-stage oligomerization feed stream is converted to mostly butenes. A second-stage oligomerization olefin stream having a higher average carbon number than the second-stage oligomerization feed stream in line 28 exits second-stage oligomerization reactor 32 in line 34.

[0054] The second-stage oligomerization catalyst is preferably an amorphous silica-alumina base having a metal from either Group VIII and / or Group VIB of the Periodic Table, using Chemical Abstracts Service notation. In one embodiment, the catalyst has a Group VIII metal promoted with a Group VIB metal. Typically, the silica and alumina are present only in the base, so the silica to alumina ratio is the same for both the catalyst and the base. The metal can be impregnated onto the silica-alumina base or ion-exchanged into the silica-alumina base. Co-mulling is also contemplated. The catalyst of the present invention may have a low temperature acidity ratio of at least 0.15, suitably 0.2, and preferably greater than 0.25, as measured by ammonia temperature programmed desorption (TPD), as described below. Additionally, suitable catalysts have a low temperature acidity ratio of 50 to 400 m as determined by the nitrogen BET method. 2 / g of surface area.

[0055] A preferred second-stage oligomerization catalyst comprises an amorphous silica-alumina support. One of the components of the catalyst support utilized in the present invention is alumina. The alumina can be any of a variety of hydrous aluminum oxides or alumina gels, such as alpha-alumina monohydrate with a boehmite or pseudo-boehmite structure, alpha-alumina trihydrate with a gibbsite structure, or beta-alumina trihydrate with a bayerite structure. A particularly preferred alumina is available from Sasol North America Alumina Product Group under the Catapal trademark. This material is an extremely high-purity alpha-alumina monohydrate (pseudo-boehmite) that has been shown to yield high-purity gamma-alumina after calcination at high temperatures. Another component of the catalyst support is amorphous silica-alumina. Suitable silica-alumina having a silica to alumina ratio of 2.6 is available, for example, from CCIC, a subsidiary of JGC in Japan.

[0056] Another component utilized in the preparation of the second-stage oligomerization catalyst utilized in the present invention is a surfactant. The surfactant is preferably mixed with the alumina and silica-alumina powders described above. The resulting surfactant, alumina, and silica-alumina mixture is then formed, dried, and calcined as described below. Calcination effectively removes the organic components of the surfactant by combustion, but only after the surfactant has faithfully performed its function in accordance with the present invention. Any suitable surfactant may be utilized in accordance with the present invention. A preferred surfactant is selected from the series of commercially available surfactants sold by Solvay SA under the "Antarox" trademark. "Antarox" surfactants are generally characterized as modified linear aliphatic polyethers and are low-foaming biodegradable detergents and wetting agents.

[0057] A suitable silica-alumina mixture is prepared by mixing proportional volumes of silica-alumina and alumina to achieve the desired silica to alumina ratio. In one embodiment, 75-99 wt. % amorphous silica-alumina with a silica:alumina ratio of 2.6 and 10-20 wt. % alumina powder provides a suitable support. In embodiments, other ratios of amorphous silica-alumina to alumina may be suitable.

[0058] Any convenient method can be used to incorporate the surfactant into the mixture of silica-alumina and alumina. The surfactant is preferably mixed during the mixing and formation of the alumina and silica-alumina. A preferred method is to mix an aqueous solution of the surfactant with the alumina and silica-alumina blend prior to final formation of the carrier. The surfactant is preferably present in the paste or dough in an amount of 0.01 to 10% by weight, based on the weight of the alumina and silica-alumina.

[0059] A monobasic acid such as nitric acid or formic acid can be added to the mixture in aqueous solution to peptize the alumina in the binder. Additional water can be added to the mixture to provide sufficient wettability to form a dough with sufficient consistency to be extruded or spray dried.

[0060] The paste or dough can be prepared in the form of shaped particles; a preferred method is to extrude a dough mixture of alumina, silica-alumina, surfactant, and water through a die having openings of the desired size and shape, after which the extruded material is divided into extrudates of the desired length and dried. A further calcination step can be used to provide additional strength to the extrudates. Typically, calcination is carried out in a stream of dry air at a temperature of 260°C (500°F) to 815°C (1500°F).

[0061] The extruded particles can have any suitable cross-sectional shape, i.e., symmetrical or asymmetrical, but most often have a symmetrical cross-sectional shape, preferably spherical, cylindrical, or multilobal. The cross-sectional diameter of the particles can be as small as 40 μm. However, it is usually between 0.635 mm (0.25 inch) and 12.7 mm (0.5 inch), preferably between 0.79 mm (1 / 32 inch) and 6.35 mm (0.25 inch), and most preferably between 0.06 mm (1 / 24 inch) and 4.23 mm (1 / 6 inch).

[0062] Typical properties of the amorphous silica-alumina supports utilized herein are total pore volume, average pore diameter, and a surface area large enough to provide substantial space and area for depositing the active metal components. The total pore volume of the support, as measured by conventional mercury porosimetry, is typically 0.2-2.0 cc / gram, preferably 0.25-1.0 cc / gram, and most preferably 0.3-0.9 cc / gram. Typically, the amount of pore volume of the support in pores with diameters greater than 100 angstroms is less than 0.1 cc / gram, preferably less than 0.08 cc / gram, and most preferably less than 0.05 cc / gram. Surface area, as measured by the BET method, is typically less than 50 m 2 / g or more, e.g. 200m2 / g, preferably at least 250m 2 / gram, most preferably 300m 2 grams ~ 400m 2 / gram.

[0063] To prepare the second-stage oligomerization catalyst, the support material is combined with one or more precursors of at least one metal component from Group VIII or Group VIB of the Periodic Table, such as by single or multiple impregnation of calcined amorphous refractory oxide support particles. The Group VIII metal, preferably nickel, should be present in a concentration of 0.5 to 15 wt. %, and the Group VIB metal, preferably tungsten, should be present in a concentration of 0 to 12 wt. %. Impregnation can be accomplished by any method known in the art, such as spray impregnation, in which a solution containing the metal precursors in dissolved form is sprayed onto the support particles. Another method is a multi-dip procedure, in which the support material is repeatedly contacted with the impregnation solution, with or without intermittent drying. Yet another method involves immersing or circulating the support in a large volume of the impregnation solution, and yet another method is the pore volume or pore saturation technique, in which the support particles are introduced into a volume of impregnation solution just sufficient to fill the pores of the support. In some cases, the pore saturation technique may be modified to utilize an impregnation solution having a volume between 10 percent less and 10 percent more than the volume that would just fill the pores.

[0064] If the active metal precursors are incorporated by impregnation, a subsequent or second calcination at elevated temperatures, for example, 399°C (750°F) to 760°C (1400°F), converts the metals to their respective oxide forms. Optionally, a calcination may be performed after each individual active metal impregnation. Subsequent calcinations result in a catalyst containing the active metals in their respective oxide forms.

[0065] The preferred second-stage oligomerization catalyst of the present invention is in the form of 3.175 mm (0.125 inch) extrudates of an amorphous silica-alumina substrate impregnated with 0.5 to 15 weight percent nickel and has a density of 0.45 to 0.65 g / ml. It is also contemplated that the metal can be incorporated onto the support by other methods, such as ion exchange and co-mixing.

[0066] The second-stage oligomerization catalyst can be regenerated when deactivated. Suitable regeneration conditions include subjecting the catalyst, for example, in situ, to hot air at 400-500°C for 3 hours. To facilitate regeneration without downtime, a swing-bed system may be used with an alternative second-stage oligomerization reactor. The regeneration gas may include air with increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to that of fresh catalyst.

[0067] The oligomerization reaction is also exothermic in nature. The final dimerized olefin stream in line 24d comprises the diluent stream from diluent line 14, added to the first olefin stream in line 12a and conveyed through first-stage oligomerization catalyst beds 22a-22d. The diluent stream is then transported in line 28 to second-stage oligomerization reactor 32 to absorb the exothermic heat generated in the second-stage oligomerization reactor.

[0068] When the oligomerization reaction is carried out according to the above process conditions, C4 olefin conversions of 95% or greater, or even 97% or greater, are achieved. The resulting oligomerized olefin stream in line 34 comprises a plurality of olefin products that are distillate range hydrocarbons.

[0069] The oligomerized olefins stream in line 34, having a higher C8+ olefin concentration compared to the second-stage oligomerization feed stream in line 28, exchanges heat with the olefin splitter bottoms stream in line 30, reduces its pressure, and is fed to olefin splitter column 36. The oligomerized olefins stream in line 34 has a temperature of between 160°C (320°F) and 190°C (374°F) and a pressure of between 3.9 MPa (gauge) (550 psig) and 7 MPa (gauge) (1000 psig).

[0070] In olefin splitter column 36, light oligomerized olefins boiling at a lower temperature than jet range hydrocarbons, typically C7- hydrocarbons having atmospheric boiling points below 150°C, are separated into an olefin splitter overhead stream in top line 38 from a bottoms stream in bottoms line 40 containing heavy oligomerized distillate range C8+ hydrocarbons, typically C8 to C22 olefins. Olefin splitter column 36 can operate at a bottoms temperature of 250°C (482°F) to 310°C (590°F) and a top pressure of 35 kPa (gauge) (5 psig) to 175 kPa (gauge) (25 psig). It is envisioned that olefin splitter column 36 may be two columns.

[0071] The olefin splitter overhead stream may be cooled to 66°C (150°F) to 93°C (200°F), and the resulting condensed liquid portion is refluxed from olefin splitter receiver 42 back to olefin splitter column 36. The net vapor stream in receiver overhead line 44 from olefin splitter receiver 42 can be compressed to oligomerization pressure in off-gas compressor 46 to provide, after cooling, a light oligomer stream in line 48 in either the vapor or liquid phase. Alternatively, the olefin splitter overhead stream in overhead line 38 can be fully condensed, perhaps by cooling in an external refrigeration loop, to provide a liquid light oligomer stream in line 48. The light oligomer stream in line 48 can be split into a light olefin drag stream in line 50 and an oligomer recycle stream in line 26, which can be recycled to first-stage oligomerization reactor 22 or, alternatively, to second-stage oligomerization reactor 32. The light olefin drag stream in line 50 can comprise 1 to 15 weight percent of the light oligomer stream in line 48. The light oligomer stream in line 48 can contain 40 to 80 weight percent C4 to C8 olefins. In one embodiment, the recycled oligomer stream in line 26 can be mixed with the first diluent olefin stream in line 16a, or split into first to fourth diluent olefin streams in lines 16a to 16d, or split only into the penultimate diluent olefin stream in line 16c and / or the final diluent olefin stream in line 16d to dimerize unreacted ethylene.

[0072] The heavy olefins stream in splitter bottoms line 40 may be split into a reboil stream that is reboiled back to olefin splitter column 36 and a heavy olefins stream in net splitter bottoms line 30. The heavy olefins stream in net bottoms line 30 is cooled by heat exchange with the oligomerized olefins stream in line 34 and then transported to hydroprocessing section 80 of Figure 2. The reboil in line 40 may be provided by a thermosiphon reboiler.

[0073] Referring to hydroprocessing section 80 of Figure 2, the heavy oligomerized olefin stream in net olefin splitter bottoms line 30 from Figure 1, which contains C8+ oligomerized olefins in the distillate range, can be hydrogenated in hydrogenation reactor 52 to saturate the olefinic bonds and provide a fuel. This step is carried out to ensure that the product motor fuel meets or exceeds the thermal oxidation requirements specified in ASTM D7566-10a for hydrotreated synthetic paraffinic kerosene (SPK). Furthermore, saturating the heavy oligomerized olefins results in a paraffinic stream that can be used as a diluent stream in line 14.

[0074] The heavy oligomerized olefins stream in line 30 may be cooled by heat exchange with the hydrogenation stream in line 60, cooled in a first steam generator to generate steam, and combined with the hydrogenation stream in line 56 to produce a combined oligomerized hydrogenation feed stream in line 54. The light olefins drag stream containing C2 to C7 olefins in line 50, also from FIG. 1, may also be combined with the heavy oligomerized olefins stream in line 30 to produce a combined oligomerized hydrogenation feed stream in line 54. The combined oligomerized hydrogenation feed stream in line 54 may be cooled and charged to hydrogenation reactor 52 at 125°C (257°F) to 250°C (482°F) and 3.5 MPa (500 psig) to 6.9 MPa (1000 psig). Alternatively, the combined oligomerized hydrogenation feed stream in line 54 is not cooled before being charged to hydrogenation reactor 52. An excess of hydrogen, such as 1.5 to 2.5 stoichiometric hydrogen, can be used to ensure complete saturation.

[0075] Hydrogenation is typically carried out using conventional hydrogenation or hydrotreating catalysts, which may include, for example, metal catalysts containing palladium, rhodium, nickel, ruthenium, platinum, rhenium, cobalt, molybdenum, or combinations thereof, and supported metal catalysts thereof. The catalyst support may be any solid, inert material, including, but not limited to, oxides such as silica, alumina, titania, calcium carbonate, barium sulfate, and carbon. The catalyst support may be in the form of a powder, granules, pellets, etc.

[0076] In an exemplary embodiment, the hydrogenation is carried out in a hydrogenation reactor 52 containing an alumina-supported platinum catalyst, for example, 0.5 wt % to 0.9 wt % alumina-supported platinum catalyst. The hydrogenation reactor 52 converts the olefins by contact with the hydrogenation catalyst in the presence of hydrogen to paraffin products having the same carbon number distribution as the olefins, thereby forming paraffins in the distillate range suitable for use as jet and diesel fuels.

[0077] The saturated heavy oligomerized olefins stream discharged from hydrocracking reactor 52 in line 60 may be cooled by heat exchange with a heavy oligomerized olefins stream in line 30 and heated by heat exchange with a cracked stream in line 92. The heated saturated heavy oligomerized olefins stream is combined with a hydrocracking hydrogen stream in line 94 to provide a combined saturated heavy oligomerized olefins stream in line 96, and heated to cracking temperatures in a fired heater to provide a heated combined saturated heavy oligomerized olefins stream. The heated combined saturated heavy oligomerized olefins stream in line 96 is input to cracking reactor 100.

[0078] Cracking reactor 100 can be a fixed bed reactor containing one or more vessels, single or multiple catalyst beds within each vessel, and various combinations of hydrocracking catalysts within one or more vessels. Hydrocracking reactor 100 can be operated in a conventional continuous gas phase, moving bed, or fluidized bed hydroprocessing reactor.

[0079] The heated combined saturated heavy oligomerized olefin stream in line 96 is hydrocracked over a hydrocracking catalyst in at least one hydrocracking catalyst bed in hydrocracking reactor 100 in the presence of a hydrocracking hydrogen stream from hydrocracking hydrogen line 94 to provide a hydrocracked stream in line 92.

[0080] The hydrocracking reactor can provide an overall conversion of at least 20% by volume, and typically greater than 60% by volume, of the heated combined saturated heavy oligomerized olefin stream in the heated hydrocracking input line 96 to products boiling in the jet fuel range. Diesel-boiling paraffins containing C18+ hydrocarbons are cracked to jet-boiling paraffins containing C8 to C16 hydrocarbons. The hydrocracking reactor 100 can be operated at a partial conversion of greater than 30% by volume of the feed, or at a full conversion of at least 90% by volume, based on overall conversion. The hydrocracking reactor 100 can be operated at mild hydrocracking conditions that provide an overall conversion of 20 to 60% by volume, preferably 20 to 50% by volume, of the hydrocracking input stream to products boiling in the jet fuel range. When hydrocracking reactor 100 is operated at partial conversion, recycle streams removed from the liquid cracked oligomerization stream in hot separator bottoms line 66 or from the small amount of heavy drag stream in net stripping bottoms line 82 may be recycled back to hydrocracking reactor 100, although these recycle lines are not shown in the drawings.

[0081] The hydrocracking catalyst may utilize an amorphous silica-alumina base or a zeolite base combined with one or more Group VIII or Group VIB metal hydrogenation components to selectively produce a balance of light diesel and jet fuel distillates. Alternatively, generally, catalysts comprising any crystalline zeolite cracking base upon which a Group VIII metal hydrogenation component is deposited may be suitable. An additional hydrogenation component may be selected from Group VIB for incorporation into the zeolite base. Additionally, a hydroisomerization catalyst may be used as or added to the hydrocracking catalyst in the hydrocracking reactor 100, but function at the high end of the hydroisomerization temperature range.

[0082] Zeolite cracking bases, sometimes referred to in the art as molecular sieves, are typically composed of silica, alumina, and one or more exchangeable cations, such as sodium, magnesium, calcium, or rare earth metals. They are further characterized by relatively uniform crystalline pore diameters of 4 to 14 angstroms. Zeolites with a relatively high silica / alumina molar ratio of 3 to 12 are preferred. Suitable naturally occurring zeolites include, for example, mordenite, stilbite, heulandite, ferrierite, dacialdite, chabazite, erionite, and faujasite. Suitable synthetic zeolites include, for example, B, X, Y, and L crystal types, such as synthetic faujasite and mordenite. Preferred zeolites have crystalline pore diameters of 8 to 12 angstroms and silica / alumina molar ratios of 4 to 6. One example of a preferred group of zeolites is synthetic Y molecular sieve.

[0083] Naturally occurring zeolites are usually found in sodium, alkaline earth metal, or mixed forms. Synthetic zeolites are almost always prepared in the sodium form. In any case, for use as a cracking base, it is preferred to ion-exchange most or all of the original zeolite's monovalent metal with a polyvalent metal and / or ammonium salt, followed by heating to decompose the ammonium ions associated with the zeolite, leaving hydrogen ions and / or exchange sites at those sites that are effectively decationized by further removal of water. This type of hydrogen or "decationized" Y zeolite is described in more detail in U.S. Pat. No. 3,100,006.

[0084] Mixed polyvalent metal-hydrogen zeolites can be prepared by ion-exchanging with an ammonium salt, followed by partial back-exchanging with a polyvalent metal salt, and then calcining. In some cases, as in the case of synthetic mordenite, the hydrogen form can be prepared by direct acid treatment of the alkali metal zeolite. In one aspect, the preferred cracking base is one that is at least 10% by weight, and preferably at least 20% by weight, deficient in metal cations, based on the initial ion-exchange capacity. In another aspect, a desirable and stable class of zeolites is one in which at least 20% by weight of the ion-exchange capacity is filled with hydrogen ions.

[0085] The active metals used as the hydrogenation component in the preferred hydrocracking catalysts of the present disclosure are those of Group VIII, namely, iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium, and platinum. In addition to these metals, other promoters, including Group VIB metals, such as molybdenum and tungsten, may also be used therewith. The amount of hydrogenation metal in the catalyst may vary within a wide range. Generally, any amount between 0.05 wt. % and 30 wt. % may be used. In the case of noble metals, it is usually preferred to use 0.05 wt. % to 2 wt. % of the noble metal. Although noble metals can deactivate noble metal catalysts, they may be preferred as hydrogenation metals on hydrocracking catalysts to provide selectivity to jet fuel due to the absence of hydrogen sulfide and ammonia being removed upstream in the process.

[0086] A method for incorporating the hydrogenation metals is to contact the base material with an aqueous solution of a suitable compound of the desired metal, where the metal is present in cationic form. After adding the selected hydrogenation metal(s), the resulting catalyst powder is filtered, dried, pelletized with the addition of lubricants, binders, etc., as desired, and calcined in air at temperatures between 371°C (700°F) and 648°C (200°F) to activate the catalyst and decompose the ammonium ions. Alternatively, the base component can be pelletized, followed by the addition of the hydrogenation component and activation by calcination.

[0087] The aforementioned catalysts can be used in undiluted form, or the powder catalyst can be mixed and co-pelletized with other less active catalysts, diluents, or binders, such as alumina, silica gel, silica-alumina cogels, activated clays, etc., in proportions ranging from 5 to 90% by weight. These diluents can be used neat or can contain small amounts of added hydrogenation metals, such as Group VIB and / or Group VIII metals. Additional metal-promoted hydrocracking catalysts can also be utilized in the processes of the present disclosure, including, for example, aluminophosphate molecular sieves, crystalline chromosilicates, and other crystalline silicates. Crystalline chromosilicates are more fully described in U.S. Pat. No. 4,363,178.

[0088] According to one approach, hydrocracking conditions may include a temperature of 290°C (550°F) to 468°C (875°F), preferably 371°C (700°F) to 427°C (800°F), a pressure of 3.1 MPa (gauge) (450 psig) to 4.5 MPa (gauge) (650 psig), a liquid hourly space velocity (LHSV) of 0.4 to less than 2.5 hr, and a hydrogen rate of 337 Nm / m oil (2,000 scf / bbl) to 2,527 Nm / m oil (15,000 scf / bbl). The hydrocracking pressure may vary with the hydrogenation pressure.

[0089] The hydroisomerization, including hydrodewaxing, of the heated combined saturated heavy oligomerized olefin stream in heated hydrocracking input line 96 can also be carried out in one or more beds of hydroisomerization catalyst, either in a hydrocracking catalyst bed in hydrocracking reactor 100, a separate hydroisomerization catalyst bed in hydrocracking reactor 100, or in a separate hydroisomerization reactor. In the embodiment of Figure 2, hydroisomerization can occur simultaneously with the hydrocracking occurring in hydrocracking reactor 100.

[0090] The hydroisomerization catalyst may comprise a dehydrogenation metal, a molecular sieve, and a metal oxide binder. The hydroisomerization catalyst may comprise a dehydrogenation metal comprising a Group VIII metal. The dehydrogenation metal may be selected from platinum, palladium, nickel, nickel molybdenum sulfide, or nickel tungsten sulfide. Preferably, the dehydrogenation metal is selected from platinum or nickel tungsten sulfide. The concentration of the dehydrogenation metal on the hydroisomerization catalyst may comprise 0.05 to 5 wt. % based on the transition metal.

[0091] The dehydrogenation metals are distributed between the molecular sieve and the binder, with 40 to 65 wt. %, preferably 45 to 60 wt. %, of the metals distributed on the molecular sieve and 40 to 65 wt. %, preferably 45 to 60 wt. %, of the metals distributed on the binder. An associated advantage of the hydroisomerization catalyst is high activity and selectivity for hydroisomerization. In a further embodiment, the hydroisomerization catalyst further comprises less than 0.5 wt. % carbon, with the associated benefits of high activity and selectivity for hydroisomerization.

[0092] In one embodiment, the hydroisomerization catalyst comprises one or more molecular sieves having a topology selected from AEI, AEL, AFO, AFX, ATO, BEA, CHA, FAU, FER, MEL, MFI, MOR, MRE, MTT, MWW, or TON, e.g., EU-2, ZSM-11, ZSM-22, ZSM-23, ZSM-48, SAPO-5, SAPO-11, SAPO-31, SAPO-34, SAPO-41, SSZ-13, SSZ-16, SSZ-39, MCM-22, zeolite Y, ferrierite, mordenite, ZSM-5, or zeolite beta, with the associated benefit that the molecular sieve is active in the hydroisomerization of linear hydrocarbons.

[0093] The metal oxide binder may be taken from the group comprising alumina, silica, silica-alumina, and titania, or mixtures thereof. Preferably, the metal oxide binder is alumina, preferably gamma alumina.

[0094] Hydroisomerization catalysts typically comprise particles having diameters of 1 to 5 millimeters. Catalyst production typically involves the formation of a stable porous support followed by impregnation with the active metal. Stable porous supports typically include metal oxides and molecular sieves, which may be zeolites. Stable supports are produced with high porosity to ensure maximum surface area, and it is typically desirable to distribute the active metal over the entire internal and external surface area of ​​the support. DI-100 and / or DI-200, available from UOP LLC (Des Plaines, Illinois), may be suitable hydroisomerization catalysts.

[0095] The hydrocracking catalyst and hydroisomerization catalyst can be mixed together or both charged in separate beds to hydrocracking reactor 100. Alternatively, only the hydroisomerization catalyst, which has some hydrocracking function, can be charged to hydrocracking reactor 100. Under hydrocracking conditions, the hydroisomerization catalyst performs a portion of the hydrocracking function and a majority of the hydroisomerization function on the heated mixed saturated heavy oligomerized olefin stream in hydrocracking input line 96.

[0096] A cracked stream, which may also have been hydroisomerized and may be a hydroisomerized stream, may exit hydrocracking reactor 100 in cracking line 92. The cracked stream in line 92 is cooled by heat exchange with the saturated heavy oligomerized stream in line 60 and supplied to high temperature separator 62. In high temperature separator 62, the cracked stream is separated into a vaporous cracked oligomerized stream in high temperature separator top line 64 extending from the top of the high temperature separator and a liquid cracked oligomerized stream in high temperature separator bottom line 66 extending from the bottom of the high temperature separator. High temperature separator 62 may operate at a temperature of 200°C (400°F) to 320°C (600°F) and at the same pressures as hydrocracking reactor 100.

[0097] The liquid cracked oligomerization stream in line 66 may be heated by heat exchange with the paraffin stream in diluent line 14. The cooled liquid cracked oligomerization stream in line 66 may be fed to stripping column 70 along with the cold liquid cracked oligomerization stream in cold bottoms line 114 to strip light gases in stripping feed line 120. The cold liquid cracked oligomerization stream in cold bottoms line 114 may be combined with the cooled liquid cracked oligomerization stream in line 66 to provide a combined liquid cracked oligomerization stream in stripping feed line 120. The vaporous cracked oligomerization stream may be cooled and fed to cold separator 110. In the cold separator 110, the vaporous cracked oligomerization stream is separated into a cold vaporous cracked oligomerization stream in a cold separator top line 112 extending from the top of the cold separator 110 and a cold liquid cracked oligomerization stream in a cold bottom line 114 extending from the bottom of the cold separator. The cold separator may operate at temperatures between 90°C (194°F) and 125°C (257°F) and at the same pressures as the hot separator 62.

[0098] A reserve hydrogen stream in line 84 supplies hydrogen to make-up compressor 86, providing a make-up hydrogen stream in line 68. A normally no-flow purge stream in line 65 can be removed from the cold vaporous cracked oligomerization stream in line 112, and the remaining recycle stream in recycle line 115 can be compressed in recycle compressor 118 and combined with the make-up hydrogen in line 68 downstream of make-up compressor 86 to provide a hydrogen stream in line 116. A cold liquid cracked oligomerization stream in cold bottoms line 114 can be combined with the liquid cracked oligomerization stream in line 66 to provide a combined liquid cracked oligomerization stream in stripping feed line 120, which can be fed to stripping column 70 for light gas stripping.

[0099] The combined liquid cracked oligomerization stream in stripping feed line 66 may be fed to stripping tower 70 to strip light gases and separate into product streams. A downstream fractionation tower may not be required to obtain a jet fuel product stream after stripping. In stripping tower 70, the combined liquid cracked oligomerization stream may be separated into an off-gas stream in top line 72, a green jet stream in side line 74 from the side of stripping tower 70, and a heavy drag stream in bottom line 76. Stripping tower 70 may operate at a bottom temperature of 370°C (700°F) to 426°C (800°F) and a top pressure of 35 kPa (5 psig) to 350 kPa (50 psig).

[0100] The stripping overhead stream in overhead line 72 may be cooled and the resulting condensate portion is refluxed from stripping receiver 78 back to stripping column 70 in line 79, while a net off-gas stream comprising hydrogen and C7-hydrocarbon cracking products is removed from stripping receiver 78 in receiver overhead line 80. The majority of the hydrocarbons in the net off-gas stream in receiver overhead line 80 are lighter hydrocarbons and can be used to fuel the reboilers of stripping column 70 and / or olefin splitter column 36, or sent to hydrogen recovery, or sent to a fuel gas header.

[0101] The green jet stream removed from the lower half of stripping tower 70 in side line 74 contains C8 to C17 hydrocarbons in the kerosene range. The green jet stream in side line 74 may be split into a first stream in jet product line 75, which may be cooled and removed as a jet product stream meeting applicable SPK standards, including an ASTM D86 endpoint of 300°C (572°F) or less. The jet fuel product stream in jet product line 75 may be fed to a jet fuel pool. The green jet stream in side line 74 may also be split into a second stream containing a diluent stream in line 14. The diluent stream in line 14 may be pumped and cooled by heat exchange with the liquid cracked oligomerization stream in hot separator bottoms line 66. The cooled diluent stream in line 14 can be recycled back and mixed with the olefin stream in line 12, preferably the first olefin stream in line 12a, in oligomerization section 10 of FIG. 1 to provide a first diluent olefin stream in line 16a and absorb heat generation in first-stage oligomerization reactor 22. Alternatively, some or all of the cooled diluent stream in line 14 can be recycled back and mixed with the second-stage oligomerization input stream in line 28. Because the green jet fuel in diluent line 14 is paraffinic, it can be inert to the dimerization, oligomerization, and hydrogenation reactions that the green jet fuel may undergo. It is also contemplated that the liquid cracked oligomerization stream in hot separator bottoms line 66 can be removed as the diluent stream in line 14.

[0102] In an alternative embodiment, the green jet stream can be removed from the condensate stream in line 79 exiting stripping receiver 78 instead of refluxing all of the condensate to stripping column 70. This green jet stream removed in line 79 may have to be stripped to remove light ends. In such an embodiment, since the green jet fuel is recovered from the condensate stream in line 79, no side line 74 is provided to recover the green jet fuel stream.

[0103] The diesel stream is taken off in bottoms line 76 and split into a reboil stream that is reboiled and returned to stripping column 70 in line 77, and a small heavy drag stream in net stripping bottoms line 82 that contains a green diesel product stream to achieve jet cut specifications for the green jet stream. The green diesel product stream in line 82 may be less than 10 wt. %, suitably less than 5 wt. %, and preferably less than 1 wt. % of the combined liquid cracked oligomerization stream in line 120. The diesel stream meets ASTM D975 standards for diesel and can be blended into a diesel pool.

[0104] The disclosed process can efficiently produce green jet fuel starting from ethylene that meets applicable fuel requirements while managing exothermic heat generation. Carbon recovery in the process can exceed 95%. [Example]

[0105] The heavy hydrocarbon cracking performance of the amorphous silica-alumina catalyst was tested using a hydrogenated C18+ hydrocarbon feed mixture at a flow rate of 1.3 kmol / hr. The test was conducted at a liquid hourly space velocity of 1 h with the goal of high hydrocarbon conversion. -1 The experiments were carried out at a pressure of 40.8 atm, a hydrogen to hydrocarbon ratio of 4:1, and at different temperatures. The results are shown in the table below.

[0106] [Table 1]

[0107] As can be seen, the amorphous silica-alumina catalyst can increase the yield of C8-C17 jet range hydrocarbons in the product. The conditions can be further optimized to increase the cracking selectivity for C8-C17 jet range hydrocarbons.

[0108] Specific Embodiments While the following will be described in conjunction with specific embodiments, it will be understood that this description is illustrative of the preceding description and appended claims, and is not intended to limit them.

[0109] A first embodiment of the present disclosure is a process for oligomerizing an olefin stream, comprising: dimerizing the olefin stream with a first stage oligomerization catalyst to produce a first stage oligomerized olefin stream; oligomerizing the first stage oligomerized olefin stream with a second stage oligomerization catalyst to provide a second stage oligomerized stream; and cracking the second stage oligomerized stream to provide a cracked oligomerized stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiments of this paragraph through the first embodiment of this paragraph, further comprising hydrogenating the second stage oligomerized olefin stream prior to cracking the second stage oligomerized stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this paragraph, further comprising separating the second-stage oligomerization stream into a light oligomerization stream and a heavy oligomerization stream, and cracking the heavy oligomerization stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this paragraph, further comprising hydrogenating the heavy oligomerization stream prior to cracking it. An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this paragraph, further comprising hydroisomerizing the second-stage oligomerization stream to provide a cracked oligomerization stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this paragraph, further comprising separating the cracked oligomerization stream to provide a vaporous cracked oligomerization stream and a liquid cracked oligomerization stream. An embodiment of the present invention is one, any, or all of the preceding embodiments to the first embodiment of this paragraph, further comprising stripping the liquid cracked oligomerization stream to provide a green jet stream. An embodiment of the present invention is one, any, or all of the preceding embodiments to the first embodiment of this paragraph, further comprising recycling a portion of the green jet stream to the dimerization step.An embodiment of the present disclosure is one, any, or all of the preceding embodiments to the first embodiment of this paragraph, further comprising stripping the liquid cracked oligomerization stream to provide a heavy drag stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiments to the first embodiment of this paragraph, further comprising cooling and separating the vaporous cracked oligomerization stream to provide a low temperature vaporous cracked oligomerization stream and a low temperature liquid cracked oligomerization stream, and stripping the low temperature liquid cracked oligomerization stream with the liquid oligomerization stream.

[0110] A second embodiment of the present disclosure is a process for oligomerizing an olefin stream, comprising: dimerizing the olefin stream with a first stage oligomerization catalyst to produce a first stage oligomerized olefin stream; oligomerizing the first stage oligomerized olefin stream with a second stage oligomerization catalyst to provide a second stage oligomerized olefin stream; hydrogenating the second stage oligomerized olefin stream to provide a hydrogenated oligomerized stream; and cracking the hydrogenated oligomerized stream to provide a cracked oligomerized stream. An embodiment of the present disclosure is one, any, or all of the previous embodiment to the second embodiment of this paragraph, further comprising separating the second stage oligomerized olefin stream into a light oligomerized olefin stream and a heavy oligomerized olefin stream; and hydrogenating the heavy oligomerized olefin stream. An embodiment of the present disclosure is any one, any, or all of the preceding through second embodiments of this paragraph, further comprising hydroisomerizing the second-stage oligomerization stream to provide a cracked oligomerization stream. An embodiment of the present disclosure is any one, any, or all of the preceding through second embodiments of this paragraph, further comprising separating the cracked oligomerization stream to provide a vaporous cracked oligomerization stream and a liquid cracked oligomerization stream. An embodiment of the present disclosure is any one, any, or all of the preceding through second embodiments of this paragraph, further comprising stripping the liquid cracked oligomerization stream to provide a green jet stream. An embodiment of the present disclosure is any one, any, or all of the preceding through second embodiments of this paragraph, further comprising recycling a portion of the green jet stream to the dimerization step. An embodiment of the present invention is one, any, or all of the previous embodiment of this paragraph through the second embodiment of this paragraph, further comprising stripping the liquid cracked oligomerization stream to provide a heavy drug stream.An embodiment of the present disclosure is one, any, or all of the previous embodiment to the second embodiment of this paragraph, further comprising cooling and separating the vaporous cracked oligomerization stream to provide a low temperature vaporous cracked oligomerization stream and a low temperature liquid vaporous cracked oligomerization stream, and stripping the low temperature liquid cracked oligomerization stream along with the liquid oligomerization stream.

[0111] A third embodiment of the present disclosure is a process for oligomerizing an olefin stream, comprising: dimerizing the olefin stream with a first stage oligomerization catalyst to produce a first-stage oligomerized olefin stream; oligomerizing the first-stage oligomerized olefin stream with a second stage oligomerization catalyst to provide a second-stage oligomerized olefin stream; hydrogenating the second-stage oligomerized olefin stream to provide a hydrogenated oligomerized stream; and cracking and isomerizing the hydrogenated oligomerized stream to provide a cracked oligomerized stream. An embodiment of the present disclosure is one, any, or all of the previous embodiment through the third embodiment of this paragraph, further comprising separating the second-stage oligomerized olefin stream into a light oligomerized olefin stream and a heavy oligomerized olefin stream; and hydrogenating the heavy oligomerized olefin stream.

[0112] Without further elaboration, it is believed that, using the preceding description, one skilled in the art can utilize the present disclosure to its fullest extent and easily ascertain the essential characteristics of the present disclosure, and can make various changes and modifications to the present disclosure to adapt it to various uses and conditions, without departing from the spirit and scope of the present disclosure. 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, and are intended to cover various modifications and equivalent arrangements included within the scope of the appended claims.

[0113] Above, all temperatures are set forth in degrees Celsius and all parts and percentages are by weight unless otherwise indicated.

Claims

1. 1. A process for oligomerizing an olefin stream, comprising: dimerizing the olefin stream over a first-stage oligomerization catalyst to produce a first-stage oligomerized olefin stream; oligomerizing the first-stage oligomerized olefin stream with a second-stage oligomerization catalyst to provide a second-stage oligomerized stream; cracking said second-stage oligomerization stream to provide a cracked oligomerization stream.

2. 10. The process of claim 1, further comprising hydrogenating said second-stage oligomerized olefin stream prior to cracking said second-stage oligomerized stream.

3. 10. The process of claim 1, further comprising separating the second-stage oligomerization stream into a light oligomerization stream and a heavy oligomerization stream, and cracking the heavy oligomerization stream.

4. 4. The process of claim 3, further comprising hydrogenating said heavy oligomerization stream prior to cracking said heavy oligomerization stream.

5. 10. The process of claim 1, further comprising hydroisomerizing said second-stage oligomerization stream to provide said cracked oligomerization stream.

6. 10. The process of claim 1, further comprising separating the cracked oligomerization stream to provide a vaporous cracked oligomerization stream and a liquid cracked oligomerization stream.

7. 7. The process of claim 6, further comprising stripping the liquid cracked oligomerization stream to provide a green jet stream.

8. 8. The process of claim 7, further comprising recycling a portion of said green jet stream to said dimerizing step.

9. 8. The process of claim 7, further comprising stripping the liquid cracked oligomerization stream to provide a heavy drug stream.

10. 9. The process of claim 8, further comprising cooling and separating the vaporous cracked oligomerization stream to provide a low temperature vaporous cracked oligomerization stream and a low temperature liquid cracked oligomerization stream, and stripping the low temperature liquid cracked oligomerization stream along with the liquid oligomerization stream.

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