Multiphase Oregonation process for converting olefins to jet fuel.

Abstract

1. A process for oligomerizing an olefin stream, comprising: passing an olefin stream, the olefin stream comprising a C2- olefin vapor stream and a C3+ olefin liquid stream, to a first stage oligomerization reactor comprising a solid acid oligomerization catalyst to produce a first stage oligomerization stream; and oligomerizing the first stage oligomerization stream in a second stage oligomerization reactor comprising a metal catalyst to provide a second stage oligomerization stream.

Description

[Technical Field]

[0001] The field is the conversion of olefins to distillates. The field may in particular relate to the oligomerization of olefins to distillate fuels. [Background technology]

[0002] Molecular sieves, such as microporous crystalline zeolites, and non-zeolitic catalysts, particularly silicoaluminophosphate (SAPO), are known to promote the conversion of oxygenates, such as methanol, to light olefins. The highly efficient methanol-to-olefin (MTO) process can convert oxygenates to light olefins and has typically been explored for plastics production. Light olefins produced from the MTO process are highly enriched in ethylene and propylene, and also contain significant concentrations of butenes, pentenes, and hexenes. When methanol derived from low-carbon-intensity feedstocks, such as carbon dioxide or municipal solid waste, is fed to an MTO unit, renewable light olefins are produced.

[0003] Ethylene can be dimerized and oligomerized to olefins such as C4, C6, and C8 olefins. Propylene can be dimerized and oligomerized to olefins such as C6, C9, and C12 olefins. Ethylene and propylene can be co-oligomerized to olefins such as C5 and C7 olefins. Olefin oligomerization is an exothermic process that can oligomerize smaller olefins to larger olefins. More specifically, olefins, including oligomerized 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] Unlike automobile engines, jet engines cannot be easily replaced by electric motor systems because fueling an aircraft requires a high energy density that batteries cannot provide. In certain regions, significant incentives are currently available for renewable jet fuel. Other regions have announced planned mandates for renewable jet fuel to be implemented gradually over the next few decades in order to meet carbon dioxide emission reduction targets.

[0005] Efficient processes for converting renewable olefin feeds into distillate fuels are desirable. Summary of the Invention

[0006] The present inventors have devised a process for oligomerizing olefin streams into distillate fuels.

[0007] The input olefin stream is a C2 olefin vapor stream that is compressed and mixed with a C3+ liquid stream. In a first embodiment, the fresh vapor olefin stream and the fresh liquid olefin stream are both input to the first oligomerization bed where they are oligomerized. In alternative embodiments, the two fresh olefin streams can be split in various ways between the oligomerization beds to provide optimal yield and heat management. [Brief explanation of the drawings]

[0008] [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. [Figure 3] 1 is a plot of ethylene and propylene conversion over time. [Figure 4] 1 is a plot of product selectivity over time. [Figure 5] 1 is a plot of carbon loss for ethane, propane, and total C2-C7 paraffins over time.

[0009] definition The term "communication" means that fluid flow is operably permitted between the listed components, which may be characterized as "fluid communication."

[0010] 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.

[0011] 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.

[0012] The term "direct communication" means that fluid flow from an upstream component enters a downstream component without passing through any other intervening vessel.

[0013] The term "indirect communication" means that fluid flow from an upstream component passes through an intervening vessel before entering a downstream component.

[0014] The term "bypass" means that an object is out of downstream communication with a bypass subject, at least to the extent that it bypasses.

[0015] As used herein, the term "predominant" or "majority" means more than 50%, suitably more than 75%, preferably more than 90%.

[0016] 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 the column to the 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. Stripping columns typically feed to a top tray and remove the main product from the bottom.

[0017] 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."

[0018] 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 bottoms 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.

[0019] 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.

[0020] 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.

[0021] 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.

[0022] As used herein, the term "diesel" refers to hydrocarbons boiling within a range of "diesel cut points" including an IBP of about 125°C (257°F) to about 175°C (347°F) or a T5 of about 150°C (302°F) to about 200°C (392°F), and a T95 of about 343°C (650°F) to about 399°C (750°F) using the TBP distillation method or a T90 of 280°C (536°F) to about 340°C (644°F) using ASTM D-86. The term "green diesel" refers to diesel containing hydrocarbons not derived from fossil fuels.

[0023] As used herein, the term "jet fuel" means a hydrocarbon that boils in the range of about 190°C (374°F) to about 215°C (419°F) T10 and about 290°C (554°F) to about 310°C (590°F). The term "green jet fuel" means a jet fuel that contains hydrocarbons that do not originate from fossil fuels. DETAILED DESCRIPTION OF THE INVENTION

[0024] The disclosed process involves oligomerizing an olefin stream containing ethylene and / or propylene. The olefin stream may also contain larger C4 to C8 olefins. The inventors have discovered that C2 to C3 alkanes can be produced in the oligomerization process, which requires removal in a dealkanizer column. An olefin splitter column can be used to recover light olefins for recycle to the oligomerization process. The process and apparatus may include an oligomerization section 10 shown in FIG. 1 and a hydrogenation section 110 shown in FIG. 2.

[0025] Referring to oligomerization section 10 of Figure 1, a preliminary vapor olefin stream in line 1 containing C2 olefins is fed to pre-separator 6. The preliminary vapor olefin stream is separated in the pre-separator to provide a vapor olefin stream in line 3 and a liquefied olefin stream in line 4. The pre-separator may be olefin, and may operate at a temperature of from about 16°C (60°F) to about 38°C (100°F) and a pressure of from about 2.1 MPa(g) (300 psig) to about 2.8 MPa(g) (400 psig). The vapor olefin stream in line 3 may be compressed in compressor 8 to an oligomerization pressure in line 12.

[0026] A preliminary liquid olefin stream in line 2 containing C3+ olefins is fed to oligomerization section 10. The preliminary olefin stream may contain C3 to C8 olefins. The preliminary olefin stream in line 2 may be combined with a liquefied olefin stream in line 4, which may contain C3 to C8 olefins, and fed to liquid feed surge drum 7. A liquid olefin stream in line 5 from liquid feed surge drum 7 containing C3 to C8 liquid olefins may be combined with a compressed vapor olefin stream in line 11 to provide a vapor olefin stream in line 12.

[0027] The vapor and liquid olefin streams may comprise substantial ethylene and propylene. The vapor and liquid olefin streams may comprise primarily ethylene and / or propylene. In embodiments, the vapor and liquid olefin streams may comprise at least 95 mole percent ethylene and / or propylene. The vapor and liquid olefin streams in lines 5 and 12 may be styled light olefin streams. Additional olefin species having carbon numbers ranging from C4 to C8 may be expected to be present in the input stream. The light olefin stream may be provided by the dehydration of ethanol or from an MTO unit. The light olefin stream may be at a temperature of about 20°C (68°F) to about 150°C (302°F) and a pressure of about 2.16 MPa (350 psig), preferably about 3.5 MPa (500 psig) to about 8.4 MPa (1200 psig).

[0028] The light olefin stream may first be contacted with a first stage oligomerization catalyst to oligomerize the ethylene and propylene into oligomers and then with a second oligomerization catalyst to oligomerize the unconverted ethylene and propylene from the first stage oligomerization.

[0029] The oligomerization reaction generates a large amount 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. Therefore, the light olefin streams in lines 5 and 12 may each be split into multiple olefin streams. In FIG. 1, the light olefin streams are each split into two separate streams. The compressed vapor olefin stream in line 12 may be split into a first vapor olefin stream in line 12a and a second vapor olefin stream in line 12b. The liquid olefin stream in line 5 is split into a first liquid olefin stream in line 5a and a second liquid olefin stream in line 5b. The first vapor olefin stream in line 12a is mixed with the first liquid olefin stream in line 5a to provide a first input olefin stream in first olefin input line 13a. The second vapor olefin stream in line 12a is mixed with the second liquid olefin stream in line 5b to provide a second input olefin stream in second input olefin line 13b. More or fewer separate olefin streams may be used; up to six input olefin streams are readily contemplated.

[0030] The compressed vapor olefin stream in line 12 may be split into multiple olefin streams of equal size in lines 12a and 12b. The liquid olefin stream in line 5 may be split into multiple olefin streams of equal size in lines 5a and 5b. Alternatively, the compressed vapor olefin stream in line 12 may be split into unequal streams. Similarly, the liquid olefin stream in line 5 may be split into unequal streams. For example, either the vapor olefin stream or the liquid olefin stream, or both, may be split into streams of decreasing flow rates, with the flow rate of the input olefin stream to a preceding reactor being greater than the input olefin stream to a subsequent reactor. In an embodiment, both the vapor olefin stream or the liquid olefin stream may be split into two streams of equal flow rates, each containing 50% by volume of the input olefin stream.

[0031] In another embodiment, the first input olefin stream in first input olefin line 13a may comprise about 70 to about 90 volume percent of the input olefin stream, and the second input olefin stream in second olefin line 13b may comprise about 10 to about 30 volume percent of the input olefin stream. In another embodiment, the input streams are each split in different proportions. For example, the liquid olefin stream in line 5 may be split into two streams such that the first liquid stream in line 5a comprises 70 to 90 volume percent of the liquid olefin stream in line 5 and the second liquid stream in line 5b comprises 10 to 30 volume percent of the liquid olefin stream, while the compressed vapor stream in line 12 is split equally into two streams such that the first vapor olefin stream in line 12a and the second olefin stream in line 12b both comprise 50 volume percent of the compressed vapor olefin stream.

[0032] To manage the heat release, the input olefin stream may be diluted with a diluent stream to provide a diluted olefin stream for absorbing the heat release. The diluent stream may include a paraffin stream in diluent line 14. The diluent stream in diluent line 14 may be added to the first input olefin stream in first input olefin line 13a before being charged to first-stage oligomerization reactor 22. Preferably, the input olefin stream in lines 5 and 12 is split into multiple olefin streams before the diluent stream is added to the first input olefin stream in line 13a to provide the first diluted input olefin stream in line 16a, so that the diluent stream passes through all of the first-stage oligomerization reaction. Alternatively, the diluent stream may be split into multiple streams, with each diluent stream being added to one or more of the corresponding input olefin streams. The diluent stream may have a mass flow rate that is about 2 to about 8 times, preferably about 3 to about 6 times, the combined mass flow rate of the first input olefin stream in first input olefin line 13a and the second input olefin stream in second input olefin line 13b.

[0033] A recycle olefin stream in recycle line 26 comprising C4 to C8 olefins may be mixed with the input olefin stream and oligomerized in first-stage oligomerization reactor 22. In an embodiment, the recycle olefin stream in line 26 is split into multiple recycle olefin streams 26a-26d. The recycle olefin stream in first recycle olefin line 26a may be mixed with the first input olefin stream in line 13a and input to first-stage oligomerization reactor 22. In a further embodiment, the first recycle olefin stream in first recycle olefin line 26a is mixed with the first input olefin stream in line 13a and the diluent stream in line 14 to provide a diluted first input olefin stream in line 16a.

[0034] The first dilute input olefin stream may comprise up to 50 wt.% olefins, suitably up to 30 wt.% olefins, and preferably up to 20 wt.% olefins. In embodiments, the first dilute olefin stream comprises from about 10 to about 35 wt.% C2 to C8 olefins. The first dilute olefin stream may comprise up to 50 wt.% ethylene, suitably up to 25 wt.% ethylene, and preferably up to 20 wt.% ethylene. In embodiments, the first dilute input olefin stream comprises from about 10 to about 20 wt.% propylene. The first dilute input olefin stream may comprise up to 50 wt.% propylene, suitably up to 25 wt.% propylene, and preferably up to 20 wt.% propylene. In embodiments, the first dilute input olefin stream comprises from about 10 to about 20 wt.% propylene.

[0035] First-stage oligomerization reactor 22 may comprise a series of first-stage oligomerization catalyst beds 22a, 22b, 22c, and 22d, each for receiving an olefin input stream. First-stage oligomerization reactor 22 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 may be in a dedicated first-stage oligomerization reactor, or multiple first-stage oligomerization catalyst beds may be in two or more separate first-stage oligomerization reactors. Up to six first-stage oligomerization catalyst beds are readily contemplated. In FIG. 1, two first-stage oligomerization reactors 21a and 21b are utilized.

[0036] 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. Only two reactor vessels 21a, 21b are shown in FIG. 1.

[0037] The diluted first feed olefin stream in line 16a may be cooled in first feed cooler 18a to provide a cooled, diluted first feed olefin stream in line 20a, which may be charged to a first bed 22a of first-stage oligomerization catalyst in a first first-stage oligomerization reactor vessel 21a of first-stage oligomerization reactor 22. The cooled, diluted first feed olefin stream in line 20a may be charged at a temperature of from about 180°C (356°F) to about 260°C (500°F) and a pressure of from about 3.5 MPag (500 psig) to about 8.4 MPag (1200 psig). Feed cooler 18a may be equipped with a steam generator.

[0038] The diluted first input olefin stream may be introduced into the first first-stage catalyst bed 22a in line 20a, preferably in downflow operation. However, upflow operation may be preferred. The diluted first input olefin stream is in a mixed vapor-liquid phase, with the vapor phase comprising primarily ethylene. As oligomerization of ethylene, propylene, and recycle olefins occurs in the first first-stage oligomerization catalyst bed 22a, an exotherm is generated due to the highly exothermic nature of the olefin oligomerization reaction. As the first input olefin stream is oligomerized, a first oligomerization effluent stream is produced 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 about 260°C (500°F).

[0039] The second input olefin stream in line 13b may be combined with a second recycle olefin stream in second recycle olefin line 26b and a first oligomerization effluent stream in first oligomerization effluent line 24a removed from first first-stage oligomerization catalyst bed 22a in first first-stage reactor 21a to provide a combined second input olefin stream in line 16b. The first oligomerization effluent stream in line 24a comprises a diluent stream from diluent line 14 added to the first input olefin stream in line 13a. The second input olefin stream may comprise up to 35 wt% C2-C8 olefins, suitably up to 25 wt% C2-C8 olefins, and preferably up to 20 wt% C2-C8 olefins. The second input olefin stream may comprise up to 30 wt% ethylene, suitably up to 25 wt% ethylene, and preferably up to 20 wt% ethylene. The second input olefin stream may contain up to 30 wt.% propylene, suitably up to 25 wt.% propylene, and preferably up to 20 wt.% propylene. The second mixed input olefin stream in line 16b may be cooled in a second feed cooler 18b, which may be external to the first first-stage oligomerization reactor 21a, to provide a cooled second input olefin stream in line 20b, which may be input to a second bed 22b of first-stage oligomerization catalyst in the first first-stage oligomerization reactor 21a. The feed cooler 18b may include a steam generator.

[0040] The second cooled input olefin stream in line 20b may be input at a temperature of from about 180°C (356°F) to about 230°C (446°F) and a pressure of from about 3.5 MPag (500 psig) to about 8.4 MPag (1200 psig). The second cooled input olefin stream comprises a diluent and the olefins from the first oligomerization stream. The diluted second input olefin stream is in a mixed vapor-liquid phase, with the vapor phase comprising primarily ethylene. The olefins from the first oligomerization stream oligomerize in second first-stage catalyst bed 22b. Oligomerization of ethylene, propylene, recycle olefins, and oligomers in the second olefin stream in second first-stage oligomerization catalyst bed 22b produces a second oligomerized olefin effluent stream in second oligomerization effluent line 24b at an elevated outlet temperature. The elevated outlet temperature may be limited to about 30° C. (54° F.) to about 50° C. (90° F.) above the inlet temperature to catalyst bed 22b.

[0041] A second oligomerization effluent stream in line 24b removed from second first-stage oligomerization catalyst bed 22b in first first-stage reactor vessel 21a may be combined with a third recycle olefin stream in third recycle olefin line 26c to provide a first recycle olefin input stream in line 16c. In embodiments, neither the first input olefin stream in line 13a nor the second input olefin stream in line 13b are added directly to the first recycle olefin input stream in line 16c. Alternatively, portions of the input olefin streams in lines 13a and 13b may be added to the second oligomerization effluent stream along with the first recycle olefin input stream in line 16c. The second oligomerization effluent stream in line 24a comprises the diluent stream from diluent line 14 added to the first input olefin stream in line 13a. The first recycle olefins feed stream in line 16c may contain up to 30 wt.% ethylene, suitably up to 25 wt.% ethylene, and preferably up to 20 wt.% ethylene. The first recycle olefins feed stream in line 16c may contain up to 30 wt.% propylene, suitably up to 25 wt.% propylene, and preferably up to 20 wt.% propylene. The first recycle olefins feed stream in line 16c may 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 first recycle olefins feed stream in line 16c is cooled in a third feed cooler 18c, which may be external to oligomerization reactor 22, to provide a cooled first recycle olefins feed stream in line 20c, which may be fed to a third bed 22c of first-stage oligomerization catalyst in first-stage oligomerization reactor 22. In an embodiment, a third bed 22c of first-stage oligomerization catalyst is provided in second first-stage oligomerization reactor vessel 21b. Feed cooler 18c may comprise a steam generator.

[0042] The cooled first recycle olefins input stream in line 20c may be input at a temperature of from about 180°C (356°F) to about 230°C (446°F) and a pressure of from about 3.5 MPag (500 psig) to about 8.4 MPag (1200 psig). The first recycle olefins input stream comprises diluent and olefins from the second oligomerized olefins stream and the third recycle olefins stream. The olefins oligomerize in third catalyst bed 22c. Oligomerization of ethylene and propylene and oligomers in the first recycle olefins input stream in third bed 22c of first-stage oligomerization catalyst produces a third oligomerization effluent stream in third oligomerization effluent line 24c at an elevated outlet temperature. In an embodiment, the third oligomerization effluent stream is the penultimate oligomerization effluent stream, and the third oligomerization effluent line 24c is the penultimate oligomerization effluent line 24c. The elevated outlet temperature is limited to about 30°C (54°F) to about 50°C (90°F) above the inlet temperature to catalyst bed 22c.

[0043] A third oligomerization effluent stream in line 24c removed from second first-stage oligomerization reactor vessel 21b of first-stage oligomerization reactor 22 may be combined with a fourth recycle olefin stream in line 26d to provide a second recycle olefin input stream in line 16d. The third oligomerization effluent stream in line 24c includes the diluent stream from diluent line 14 added to the first olefin stream in line 13a. Neither input olefin stream in lines 13a nor 13b is added directly to the second recycle olefin input stream in line 16d. In embodiments, the third oligomerization effluent stream in line 24c may also be combined with and oligomerized with a portion of the input olefin stream in lines 13a and 13b. The second recycle olefin input stream may contain up to 35 wt.% C2-C8 olefins, suitably up to 30 wt.% C2-C8 olefins, and preferably up to 25 wt.% C2-C8 olefins. The second recycle olefins feed stream may contain up to 30 wt.% ethylene, suitably up to 25 wt.% ethylene, and preferably up to 20 wt.% ethylene. The second recycle olefins feed stream may contain up to 30 wt.% propylene, suitably up to 25 wt.% propylene, and preferably up to 20 wt.% propylene. The second recycle olefins feed stream in line 16d is cooled in fourth feed cooler 18d, which may be located external to second vessel 21b of first-stage oligomerization reactor 22, to provide a cooled second recycle olefins feed stream in line 20d, which may be fed to fourth bed 22d of first-stage oligomerization catalyst in the second vessel of first-stage oligomerization reactor 22. Feed cooler 18d may include a steam generator.

[0044] The cooled second recycle olefins feed stream in line 20d may be fed at a temperature of from about 180°C (356°F) to about 230°C (446°F) and a pressure of from about 3.5 MPa(g) (500 psig) to about 8.4 MPa(g) (1200 psig). The cooled second recycle olefins feed stream in line 20d comprises diluent and olefins from the third or penultimate oligomerization effluent stream and C4 to C8 olefins from the fourth recycle olefins stream. The olefins oligomerize over fourth catalyst bed 22d. Oligomerization of ethylene and propylene in the second recycle olefins feed stream in fourth bed 22d of first-stage oligomerization catalyst produces a fourth oligomerization stream in fourth oligomerization effluent line 24d at an elevated outlet temperature. The elevated outlet temperature is limited to about 30° C. (54° F.) to about 50° C. (90° F.) above the inlet temperature to catalyst bed 22d.

[0045] A fourth oligomerization effluent stream in line 24d exits second reaction vessel 21b of first-stage oligomerization reactor 22. In an embodiment, the fourth oligomerization effluent stream in line 24d is the final oligomerization effluent stream, and fourth oligomerization effluent line 24d is the final oligomerization effluent line 24d.

[0046] The first stage oligomerization reaction takes 0.5 to 10 hours based on olefins. -1 The WHSV is primarily in the liquid phase or mixed gas-liquid phase. The inventors have found that, across the first-stage oligomerization catalyst bed, typically 10 to 50 weight percent 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 propylene and butenes in the olefin stream introduced into the first-stage oligomerization catalyst bed are oligomerized. In embodiments, at least 99 mole percent of the propylene and butenes in the olefin stream are oligomerized.

[0047] 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 about 5 to about 95 wt. % of the catalyst, for example, about 5 to about 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.

[0048] 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.

[0049] 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 calcination at high temperatures.

[0050] 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 can be from about 5 to about 85 weight percent, e.g., from about 20 to about 82 weight percent, MTT zeolite, with the remainder being alumina powder to provide a suitably supported catalyst. Silica supports are also contemplated.

[0051] 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.

[0052] 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 about 260°C (500°F) to about 815°C (1500°F). The MTT catalyst is not selective to neutralize acidic sites such as amines.

[0053] 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 from about 0.635 mm (0.25 inch) to about 12.7 mm (0.5 inch), preferably from about 0.79 mm (1 / 32 inch) to about 6.35 mm (0.25 inch), and most preferably from about 0.06 mm (1 / 24 inch) to about 4.23 mm (1 / 6 inch).

[0054] In one exemplary embodiment, a catalyst bed or more of an MTT-type zeolite catalyst disposed on a high purity pseudoboehmite alumina substrate in a ratio of about 90 / 10 to about 20 / 80, preferably about 20 / 80 to about 50 / 50, is provided in the first-stage oligomerization reactor 22.

[0055] 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 about 400°C to about 500°C. To facilitate regeneration without downtime, a swinging 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 an increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to that of fresh catalyst.

[0056] 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 1, when produced from an alcohol dehydration process, requires less oxygenate removal.

[0057] The final first stage oligomerization stream in final first stage oligomerization effluent line 24d has an increased concentration of ethylene and propylene oligomers compared to the light olefin streams in lines 5 and 12. The final first stage oligomerization stream in final first stage oligomerization effluent line 24d is cooled by steam generation in steam generator 18e or by other heat exchange, further cooled by heat exchange with the second stage oligomerization stream in line 34, possibly further cooled by a steam generator or the like, to provide an input first stage oligomerization stream that is input 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 from about 80°C (176°F) to about 200°C (392°F). Second-stage oligomerization reactor 32 is operated at a pressure of from about 2.1 MPa (300 psig) to about 7.6 MPa (1100 psig), more preferably from about 3.5 MPa (500 psig) to about 6.9 MPa (1000 psig). The second-stage oligomerization input stream oligomerizes in a mixed vapor-liquid phase to primarily C4+ olefins.

[0058] 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 may be selected to produce a higher proportion of jet range olefins, which, when hydrogenated in subsequent steps as described below, result in 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 input stream is converted mostly to butenes.

[0059] Second-stage oligomerization reactor 32 may include a first reactor vessel 31a containing a first bed 32a of second-stage oligomerization catalyst and a second reactor vessel 31b containing a second bed 32b of second-stage oligomerization catalyst. A first second-stage oligomerization stream is discharged from first second-stage reactor vessel 31a, cooled, and charged to second second-stage reactor vessel 31b. A second-stage oligomerization stream having a higher average carbon number than the input second-stage oligomerization stream in line 28 exits second-stage oligomerization reactor 32 in line 34.

[0060] The first stage oligomerization reactor 22 and the second stage oligomerization reactor 32 may utilize vapor-liquid distribution trays to mix and distribute the ethylene vapor with the liquid olefins and liquid paraffins to enhance heat transfer and manage the exotherm.

[0061] 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 about 0.15, suitably about 0.2, and preferably greater than about 0.25, as determined by ammonia temperature programmed desorption (TPD) as described below. Additionally, suitable catalysts have a low temperature acidity ratio of about 50 to about 400 m as determined by the nitrogen BET method. 2 / g of surface area.

[0062] 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.

[0063] 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.

[0064] Suitable silica-alumina mixtures are prepared by mixing proportional volumes of silica-alumina and alumina to achieve the desired silica to alumina ratio. In embodiments, about 75 to about 99 weight percent amorphous silica-alumina having a silica:alumina ratio of 2.6 and about 10 to about 20 weight percent alumina powder provide a suitable support. In embodiments, other ratios of amorphous silica-alumina to alumina may be suitable.

[0065] 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 from about 0.01 to about 10% by weight, based on the weight of the alumina and silica-alumina.

[0066] 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.

[0067] 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 about 260°C (500°F) to about 815°C (1500°F).

[0068] 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 from about 0.635 mm (0.25 inch) to about 12.7 mm (0.5 inch), preferably from about 0.79 mm (1 / 32 inch) to about 6.35 mm (0.25 inch), and most preferably from about 0.06 mm (1 / 24 inch) to about 4.23 mm (1 / 6 inch).

[0069] Typical characteristics 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 about 0.2 to about 2.0 cc / gram, preferably about 0.25 to about 1.0 cc / gram, and most preferably about 0.3 to about 0.9 cc / gram. Typically, the amount of pore volume of the support in pores with diameters greater than 100 angstroms is less than about 0.1 cc / gram, preferably less than 0.08 cc / gram, and most preferably less than about 0.05 cc / gram. The surface area, as measured by the BET method, is typically less than 50 m 2 / g or more, e.g., about 200m 2 / g, preferably at least 250m 2 / gram, most preferably about 300m 2 grams ~ approx. 400m 2 / gram.

[0070] 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 about 0.5 to about 15 wt. %, and the Group VIB metal, preferably tungsten, should be present in a concentration of about 0 to about 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 about 10 percent less and about 10 percent more than the volume that would just fill the pores.

[0071] 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.

[0072] A preferred second-stage oligomerization catalyst of the present invention has an amorphous silica-alumina substrate impregnated with about 0.5 to about 15 weight percent nickel in the form of 3.175 mm (0.125 inch) extrudates and has a density of about 0.45 to about 0.65 g / mL. It is also contemplated that the metals can be incorporated onto the support by other methods, such as ion exchange and co-mixing.

[0073] 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 about 400 to about 500°C. 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.

[0074] The second-stage oligomerization reaction is also exothermic in nature. The final oligomerized olefin stream in line 24d comprises a diluent stream from diluent line 14, added to the first input olefin stream in first input olefin line 13a 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 exotherm in the second-stage oligomerization reactor. A dedicated diluent line to second-stage oligomerization reactor 32 is also contemplated for rapid control of the exotherm rise or for cooling second-stage oligomerization reactor 32.

[0075] When the oligomerization reaction is carried out according to the above process conditions, C4 olefin conversions of about 95% or greater, or even 97% or greater, are achieved. The resulting second-stage oligomerization stream in line 34 contains multiple olefin products that are distillate range hydrocarbons.

[0076] An oligomerized olefin stream in line 34, having a higher C8+ olefin concentration compared to the input second-stage oligomerization stream in line 28, is heat exchanged with the first-stage oligomerization stream in line 24d to reduce its pressure and then heat exchanged with the olefin splitter bottoms stream in line 30 and fed to dealkanizer column 40. The oligomerized olefin stream in line 34 is at a temperature of from about 160°C (320°F) to about 190°C (374°F) and a pressure of from about 3.9 MPa (gauge) (550 psig) to about 7 MPa (gauge) (1000 psig).

[0077] The inventors have discovered that light alkanes, such as ethane and / or propane, produced in the first stage oligomerization reactor 22 and / or the second stage oligomerization reactor 32 must be removed from the second stage oligomerization reaction stream for fuel production, particularly to facilitate the recycle of light olefins to the first stage oligomerization reactor 22. The light alkanes are inert and accumulate in the recycle loop. Therefore, the second stage oligomerization stream in line 34 is dealkanized by fractionation in dealkanizer column 40 to provide a light alkane stream and a dealkanized stream. In an embodiment, the light alkane stream is an ethane stream, in which case dealkanizer column 40 is a deethanizer column. In another embodiment, the light alkane stream is a propane stream, in which case dealkanizer column 40 is a depropanizer column. The light alkane stream may contain ethane and / or propane, or may be a mixture of ethane and propane.

[0078] In dealkanizer column 40, light alkanes, such as C3- and preferably C2- hydrocarbons, are separated from a dealkanizer bottoms stream in bottoms line 44, possibly containing C4+ and preferably C3+ hydrocarbons, and possibly in a light alkane overhead stream in overhead line 42. Olefins may be recycled to first-stage oligomerization reactant 22 from the dealkanizer overhead stream in overhead line 42. Dealkanizer column 40, when operating as a dealkanizer column, may operate at a bottoms temperature of from about 177°C (350°F) to about 302°C (575°F) and an overhead pressure of from about 207 kPa (gauge) (30 psig) to about -690 kPa (gauge) (100 psig). Dealkaniser column 40, when operating as a dealkanizer column, may operate at a bottoms temperature of from about 194°C (381°F) to about 333°C (630°F) and an overhead pressure of from about 207 kPa (gauge) (30 psig) to about 1.38 MPa (gauge) (200 psig).

[0079] The light alkane overhead stream in overhead line 42 is cooled and separated in dealkanizer receiver 46 to provide a dealkanizer off-gas stream in off-gas line 47, where it may be cooled and fed for further processing, such as being removed as a combustion gas in line 48 along with a net vapor stream in receiver overhead line 68. The condensate from dealkanizer receiver 46 may be refluxed back to dealkanizer column 40 in dealkanizer overhead liquid line 49. The dealkanizer off-gas stream may be used as a fuel to provide a heating duty in process 10. In an embodiment, a portion of the condensate from dealkanizer receiver 44 in line 49 may be taken as olefin recycle in line 51 to the first-stage oligomerization reactor in lines 72 and 26.

[0080] The dealkanized stream, possibly in bottoms line 44, may be split into a reboiled stream in line 50 that is reboiled via a heat exchanger with a first hot diesel stream in line 52, possibly taken from a jet fractionator bottoms heat exchanger stream in jet bottoms heat exchanger line 74, and a net bottoms stream in line 54 that is fed directly to olefin splitter column 60, possibly without heating. The reboiled bottoms stream in line 50 may be boiled back to dealkanizer column 40 to provide the heating requirements. In another embodiment, the feed to dealkanizer column 40 is not preheated by the olefin splitter bottoms stream in line 30, but the feed to olefin splitter column 60 in net bottoms line 54 is preheated by heat exchange with the olefin splitter bottoms stream in line 64.

[0081] The dealkanized stream in dealkanizer column net bottoms line 54 is divided by fractionation in olefin splitter column 60 into a light olefin stream, possibly in olefin splitter overhead line 62, and a heavy olefin stream, possibly in olefin splitter bottoms line 64. Olefins may be recycled to first-stage oligomerization reaction 22 from the olefin splitter overhead stream in olefin splitter overhead line 62. The olefin splitter overhead stream may be cooled to about 19°C (66°F) to about 93°C (200°F), and the resulting condensate portion is refluxed from olefin splitter receiver 66 back to olefin splitter column 60. The net vapor stream in receiver overhead line 68 from olefin splitter receiver 66, along with the off-gas stream in off-gas line 47, may be further processed, such as as fuel gas in line 48. The light olefins condensate from the bottom of the olefin splitter receiver in line 70 may be split into a reflux stream that is refluxed back to the column in line 71 and a light olefins recycle stream in recycle line 72 that may be recycled to either the first-stage oligomerization reactor 22 or the second-stage oligomerization reactor 32. The light olefins stream in line 72 may comprise from about 1 to about 15 weight percent, or perhaps a majority, of the light olefins stream in line 70. The light olefins stream in line 72 may contain from about 40 to about 80 weight percent C4 to C8 olefins. In embodiments, the light olefins stream in line 72 may be flashed in knockout drum 75 to remove vapors in a light olefins vapor stream that may be transported to the hydrogenation section in overhead line 77, and a liquid recycle olefin oligomer stream in line 26 may be recycled to the first-stage oligomerization reactor 22 to oligomerize C4 to C8 olefins.

[0082] The heavy olefin stream in splitter bottoms line 64 may be split between a reboil stream in splitter reboil line 65, possibly taken from the jet fractionator bottoms heat exchange stream in jet bottoms heat exchange line 74 of FIG. 2, and reboiled by heat exchange with a second hot diesel stream in line 73 returned to olefin splitter column 60. The cooled second hot diesel stream in line 112 is returned to hydrogenation section 110 of FIG. 2 for reboiling and returned to jet fractionator 100. The heavy olefin stream in net bottoms line 30 is cooled by heat exchange with the second-stage oligomerized olefin stream in line 34 and then transported to a hydrogenation section, not shown. A purge of the heavy olefin stream may occur in line 33. The heavy olefin stream comprises C9+ olefins, which, once cooled, may be transported to hydrogenation section 110.

[0083] Turning to the hydrogenation section 110 of FIG. 2, the heavy olefin stream in the net olefin splitter bottoms line 30 from FIG. 1, containing C9+ oligomerized olefins in the distillate range, can be hydrogenated in hydrogenation reactor 80 to saturate the olefinic bonds and provide fuel. This step is carried out to ensure that the product fuel meets or exceeds the thermal oxidation requirements specified in ASTM D7566-20 for Alcohol to Jet Synthesized Paraffinic Kerosene (ATJ-SPK). Further hydrogenation of the oligomerized heavy olefins produces a paraffin stream that can be used as a diluent stream in line 14. The heavy olefin stream in line 30 can be cooled to produce steam and combined with a light olefin liquid stream containing C2-C8 olefins in line 77 of FIG. 1 to produce a mixed olefin stream in line 79. Alternatively, the combined olefin stream in line 79 may be combined with the hydrogen stream in line 76 to provide a combined hydrogenation input stream in line 81 that is cooled and charged to hydrogenation reactor 80 at 125°C (257°F) to about 204°C (400°F) and 2.8 MPa (400 psig) to about 6.9 MPa (1000 psig). An excess of hydrogen, such as about 1.5 to about 5.0 stoichiometric hydrogen, may be used to ensure complete saturation.

[0084] 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.

[0085] In an exemplary embodiment, the hydrogenation is carried out in hydrogenation reactor 80 containing an alumina-supported platinum catalyst, for example, from about 0.1 wt % to about 2 wt %, preferably from about 0.5 wt % to about 0.9 wt %, of an alumina-supported platinum catalyst. In another embodiment, the hydrogenation catalyst contains from about 5 wt % to about 30 wt % nickel catalyst. Hydrogenation reactor 80 converts the olefins 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.

[0086] The hydrogenated heavy stream discharged from hydrogenation reactor 80 in line 83 may be separated in hot separator 82, which provides a hydrocarbon split. In hot separator 82, the hydrogenated heavy stream is separated into a hot hydrogenated vapor stream in overhead line 84 and a hot hydrogenated liquid stream in hot separator bottoms line 86. The hydrogenated heavy liquid stream in bottoms line 86 may be heated by heat exchange with the diluent stream in line 14 before the diluent stream is recycled to first-stage oligomerization reactor 22 of FIG. 1. The heated hydrogenated heavy liquid stream in hot bottoms line 86 may be fed to stripping column 90. The hot separator may operate at a bottoms temperature of from about 204°C (400°F) to about 343°C (650°F) and a pressure of from about 2.8 MPa (400 psig) to about 6.9 MPa (1000 psig).

[0087] The hot hydrogenated vapor stream in hot overhead line 84 may be cooled and fed to cold separator 88. The cold separator separates the cooled hot hydrogenated vapor stream in hot overhead line 84 into a cold vapor hydrogenated stream in cold overhead line 87 and a cold heavy hydrogenated liquid stream in cold bottoms line 89. A purge stream in purge line 85 may be removed from the cold vapor hydrogenated stream in cold overhead line 87, and the remainder may be compressed and combined with make-up hydrogen in line 88 to provide the hydrogen stream in line 76. The cold hydrogenated heavy liquid stream in bottoms line 89 may be fed to stripping column 90 at a feed location above the feed location of the hot hydrogenated heavy liquid stream in hot separator bottoms line 86. The low temperature separator may operate at a temperature of about 32° C. (90° F.) to about 71° C. (150° F.) and a pressure of about 2.8 MPa (400 psig) to about 4.5 MPa (650 psig).

[0088] Stripping column 90 may be a flash stripper for removing light gases from the hot hydrogenated liquid stream in hot bottoms line 86 and the cold hydrogenated liquid stream in cold bottoms line 89. Both of these streams may be fed to stripping column 90 together, or separately as shown. Stripping column 90 removes residual light gases from the liquid hydrogenated stream and provides a stripper overhead stream in stripper overhead line 92 and a stripped bottoms stream in stripper bottoms line 94. The stripper overhead stream in stripper overhead line 92 is cooled and separated in stripper receiver 96 to provide a stripper off-gas stream in stripper receiver overhead line 97 and a condensate stream in line 98 which is refluxed to the column. Stripping column 90 may operate at a bottom temperature of from about 232° C. (450° F.) to about 316° C. (600° F.) and an overhead pressure of from about 207 kPa (30 psig) to about 689 kPa (100 psig).

[0089] After undergoing stripping to remove volatiles in stripping column 90, the stripped fuel stream in stripper bottoms line 94 may be fed to jet rectifier 100 without further heating. Alternatively, stripping column 90 upstream of jet rectifier 100 may be omitted. In jet rectifier 100, the stripped fuel stream may be separated into a jet off-gas stream in overhead line 102, a green jet stream in side line 104 from the side of jet rectifier 100, and a green diesel stream in bottoms line 106. Jet rectifier 100 may operate at a bottoms temperature of about 288°C (550°F) to about 400°C (750°F) and an overhead pressure of about 35 kPa (5 psig) to about 350 kPa (50 psig).

[0090] The jet rectification overhead stream in overhead line 102 may be cooled and a portion of the resulting condensate may be refluxed from jet rectification receiver 108 to jet rectification tower 100 in jet rectification overhead liquid line 109, while the remaining condensate is taken in line 119 as a naphtha range product stream. Additionally, a net off-gas stream comprising C4 to C8-hydrocarbons is taken from jet rectification receiver 108 in receiver overhead line 105. The majority of the hydrocarbons in the net off-gas stream in receiver overhead line 105 are lighter hydrocarbons and can be used to fuel reboil heater 116 for jet rectification tower 100.

[0091] The green jet stream taken in side line 104 comprises C8 to C18 hydrocarbons in the kerosene range and can be cooled and removed as a jet fuel product meeting applicable SPK specifications. In an alternative embodiment, the green jet stream can be removed from the condensate stream in line 109 from jet rectification receiver 108 instead of returning all of the condensate to the column. This green jet stream removed in line 109 may require further stripping to remove light ends. In such an embodiment, side line 104 is not provided to recover the green jet fuel stream.

[0092] The green diesel bottoms stream in bottoms line 106 may be split into a reboil diesel stream in line 107 and a diesel product stream in line 114. The reboil diesel stream in line 107 may be split into jet bottoms heat exchange stream 74 and a bypass bottoms stream in bottoms bypass line 111. As shown in FIG. 1 , the jet bottoms heat exchange stream in jet bottoms heat exchange line 74 may be split into a first hot diesel stream in line 52 and a second hot diesel stream in line 73 to provide reboil heat to dealkanizer column 40 and olefin splitter column 60, respectively. Either the bypass bottoms stream in line 111, through a valve thereon, or the cooled second hot diesel stream in line 112, or a portion of both, is taken into jet reboil line 113, reboiled in furnace 116, and returned to jet fractionator 100.

[0093] The diesel product stream in line 114 is split into a diesel product stream in diesel product line 118 and a diluent stream in line 14. The diluent stream in line 14 can be cooled by heat exchange with the hot hydrogenated heavy liquid stream in hot separator bottoms line 86 and recycled back to be mixed with the input olefin streams in lines 13a and 13b, preferably the first input olefin stream in line 13a, in oligomerization section 10 of FIG. 1 to provide a first diluted input olefin stream in line 16a to absorb heat generation in oligomerization reactor 22. The green diesel stream in diluent line 14 is paraffinic and therefore inert to the oligomerization and hydrogenation reactions to which it may be subjected. The diesel product stream in diesel product line 118 can be cooled and fed to a diesel pool.

[0094] The disclosed process can efficiently produce sustainable aviation fuel that meets applicable fuel requirements, starting with ethylene and / or propylene, while managing exothermic heat generation. The process can also produce green gasoline and diesel by-products, which can be used as produced or as blending components in fuels that meet applicable fuel requirements depending on the specific application. Carbon recovery in the process can exceed 95%. Both the jet fuel stream in side line 104 and the diesel product stream in line 118 can be cooled and fed to their respective fuel pools. [Example]

[0095] A first-stage oligomerization catalyst containing a zeolite and a second-stage oligomerization catalyst containing a metal were loaded into a pilot plant reactor in a stacked-bed configuration. The reactor was fed primarily with ethylene, propylene, and a paraffinic diluent, along with a small amount of C4+ olefins. Tests were conducted at a pressure of 6.2 MPa (gauge) (900 psig), with the inlet temperature to the first oligomerization stage catalyst varied from 160 to 250°C, the inlet temperature to the second oligomerization stage catalyst varied from 130 to 250°C, and the WHSV of the fresh olefins charge varied from 0.25 to 2.0 hours. -1 Depending on the test conditions, the results show ethylene and propylene conversions of over 97 wt. % as shown in Figure 3 and high jet and recyclable light olefin selectivity as shown in Figure 4. Figure 5 shows that under these conditions, the light paraffin selectivity is about 3-4 wt. % at steady state, consisting mostly of ethane and propane.

[0096] The oligomerization product produced in this pilot plant experiment was then hydrogenated and fractionated. Table 1 shows the physical properties of the jet fuel produced from the hydrogenated oligomerization product.

[0097] [Table 1]

[0098] Specific Embodiments While the following will be described in conjunction with specific embodiments, it will be understood that this description is illustrative, but not intended to limit the scope of the preceding description and appended claims.

[0099] A first embodiment of the present disclosure is a process for oligomerizing an olefin stream, comprising: charging a C2 olefin vapor stream and a C3+ olefin liquid stream to a first-stage oligomerization reactor to produce a first-stage oligomerization stream; and oligomerizing the first-stage oligomerization stream in a second-stage oligomerization reactor to produce a second-stage oligomerization stream. An embodiment of the present disclosure is one, any, or all of the preceding through first embodiments of this section, wherein the C2 olefin vapor stream and the C3+ olefin liquid stream are mixed with a paraffinic liquid diluent to produce a mixed vapor-liquid stream that is sent to the first-stage oligomerization reactor. An embodiment of the present disclosure is one, any, or all of the preceding through first embodiments of this section, wherein the preliminary C2 olefin vapor stream is separated to provide a C2 olefin vapor stream and a liquefied olefin stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this section, further comprising: removing a recycle olefins stream from the second-stage oligomerization stream; and separating the recycle olefins stream with a preliminary C2 olefins vapor stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this section, further comprising: combining the liquefied olefins stream with a preliminary C3+ olefins liquid stream to provide a C3+ olefins liquid stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this section, further comprising compressing the C2 olefins vapor stream prior to injecting it into the first-stage oligomerization catalyst bed. An embodiment of the present disclosure is one, any, or all of the preceding embodiment of this section through the first embodiment of this section, wherein a C2 olefin vapor stream, a C3+ olefin liquid stream, and a paraffinic liquid diluent stream are charged to a first-stage oligomerization reactor to produce a first-stage oligomerization stream in a mixed vapor-liquid phase, wherein the vapor phase comprises primarily ethylene.An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this section, in which the first-stage oligomerization stream is oligomerized in a mixed vapor-liquid phase in a second-stage oligomerization reactor to produce a second-stage oligomerization stream that is primarily C4+ olefins. An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this section, in which the C4+ olefin stream is sent to an olefin splitter column that separates a C4-C8 olefin stream from a C8+ olefin stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this section, in which the C8+ stream is hydrogenated to produce a distillate fuel stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiments through the first embodiment of this section, in which the C3+ olefin liquid stream comprises propylene and / or butenes. An embodiment of the present disclosure is one, any, or all of the preceding through first embodiments of this section, wherein the first stage oligomerization reactor and the second stage oligomerization reactor utilize vapor-liquid distribution trays to mix and distribute the ethylene vapor with the liquid olefins and liquid paraffins to facilitate heat transfer and manage exotherm.An embodiment of the present disclosure is one, any, or all of the preceding through first embodiments of this section, wherein the first stage oligomerization reactor and the second stage oligomerization reactor operate in a downflow configuration.

[0100] A second embodiment of the present disclosure is a process for oligomerizing an olefin stream, comprising: separating a preliminary C2 olefin vapor stream to provide a C2 olefin vapor stream and a liquefied olefin stream; charging the C2 olefin vapor stream and the C3+ olefin liquid stream to a first-stage oligomerization reactor to produce a first-stage oligomerization stream; and oligomerizing the first-stage oligomerization stream in a second-stage oligomerization reactor to produce a second-stage oligomerization stream. An embodiment of the present disclosure is any one, any, or all of the preceding through second embodiments of this section, further comprising: removing a recycle olefin stream from the second-stage oligomerization stream; and separating the recycle olefin stream along with the preliminary C2 olefin vapor stream. An embodiment of the present disclosure is any one, any, or all of the preceding through second embodiments of this section, further comprising compressing the C2 olefin vapor stream prior to charging the first-stage oligomerization catalyst bed. An embodiment of the present invention is one, any, or all of the preceding embodiment to the second embodiment of this section, wherein the liquefied olefins stream is mixed with a preliminary C3+ olefins liquid stream to provide a C3+ olefins liquid stream.

[0101] A third embodiment of the present disclosure is a process for oligomerizing an olefin stream, comprising: separating a preliminary C2 olefin vapor stream to provide a C2 olefin vapor stream and a liquefied olefin stream; combining the liquefied olefin stream with a preliminary C3+ olefin liquid stream to provide a C3+ olefin liquid stream; charging the C2 olefin vapor stream and the C3+ olefin liquid stream to a first stage oligomerization reactor to produce a first stage oligomerization stream; and oligomerizing the first stage oligomerization stream in a second stage oligomerization reactor to produce a second stage oligomerization stream. An embodiment of the present disclosure is one, any, or all of the preceding through third embodiments of this section, further comprising removing a recycle olefin stream from the second stage oligomerization stream and separating the recycle olefin stream along with the preliminary C2 olefin vapor stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiment of this section through the third embodiment of this section, further comprising compressing the C2 olefin vapor stream prior to entering the first stage oligomerization catalyst bed.

[0102] 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.

[0103] 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: The process includes charging a C2 olefins vapor stream and a C3+ olefins liquid stream to a first stage oligomerization reactor to produce a first stage oligomerization stream; and oligomerizing said first-stage oligomerization stream in a second-stage oligomerization reactor to produce a second-stage oligomerization stream.

2. 10. The process of claim 1, wherein said C2 olefin vapor stream and said C3+ olefin liquid stream are mixed with a paraffinic liquid diluent to produce a mixed vapor-liquid stream that is sent to said first-stage oligomerization reactor.

3. 10. The process of claim 1, wherein a preliminary C2 olefin vapor stream is separated to provide said C2 olefin vapor stream and a liquefied olefin stream.

4. 4. The process of claim 3, further comprising removing a recycle olefins stream from said second-stage oligomerization stream and separating said recycle olefins stream along with said preliminary C2 olefins vapor stream.

5. 4. The process of claim 3, wherein the liquefied olefins stream is mixed with a preliminary C3+ olefins liquid stream to provide the C3+ olefins liquid stream.

6. 10. The process of claim 1, further comprising compressing said C2 olefin vapor stream prior to entering said first-stage oligomerization catalyst bed.

7. 10. The process of claim 1, wherein the C2 olefins vapor stream, the C3+ olefins liquid stream, and the paraffinic liquid diluent stream are charged to the first-stage oligomerization reactor to produce a first-stage oligomerization stream in a mixed vapor-liquid phase, the vapor phase comprising primarily ethylene.

8. 4. The process of claim 3, wherein said first-stage oligomerization stream is oligomerized in a mixed vapor-liquid phase in said second-stage oligomerization reactor to produce said second-stage oligomerization stream which is primarily C4+ olefins.

9. 8. The process of claim 7, wherein the C4+ olefin stream is sent to an olefin splitter column that separates a C4 to C8 olefin stream from a C8+ olefin stream.

10. 10. The process of claim 1, wherein the C8+ stream is hydrogenated to produce a distillate fuel stream.

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