A process for converting olefins into jet fuel through stripping.

Abstract

A process for oligomerizing olefins into distillate fuels, in which a recycle olefin stream is oligomerized with a feed olefin stream in a first-stage oligomerization reactor. The first-stage oligomerized product is oligomerized in a second-stage oligomerization reactor to provide a second-stage oligomerized stream. The second-stage oligomerized stream is hydrogenated to produce fuel. A stripping column is used to prepare a hydrogenated fuel stream for fractionation.

Description

TECHNICAL FIELD

[0001] The present field relates to the conversion of olefins to distillates. In particular, the present field may relate to the oligomerization of olefins to distillate fuels. BACKGROUND ART

[0002] Molecular sieves, such as microporous crystalline zeolites and non-zeolitic catalysts, particularly silicoaluminophosphates (SAPO), are known to promote the conversion of oxygenates, such as methanol, to light olefins. Highly efficient Methanol to Olefin (MTO) processes are capable of converting oxygenates to light olefins, and have typically been investigated for plastics production. The 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 into olefins such as C4, C6, and C8 olefins. Propylene can be dimerized and oligomerized into olefins such as C6, C9, and C12 olefins. Ethylene and propylene can be co-oligomerized into olefins such as C5 and C7 olefins. Olefin oligomerization is an exothermic process that can oligomerize smaller olefins into larger olefins. More specifically, olefins, including oligomerized olefins, can be converted to distillates comprising products in the jet fuel and diesel range. 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 aircraft requires a high energy density that batteries cannot provide. Significant incentives are currently available for renewable jet fuel in certain regions. In other regions, planned mandates for renewable jet fuel, to be implemented gradually over the next few decades, have been announced to meet carbon emission reduction targets.

[0005] An efficient process is needed to convert renewable olefin feedstocks into distillate fuels. [Overview of the project]

[0006] The inventors have devised a process for oligomerizing an olefin flow into distillate fuel. This process includes oligomerizing the input olefin flow in the gas phase and liquid phase, as well as the recirculated olefin flow, with a first-stage oligomerization catalyst to produce a first-stage oligomerized olefin flow. The first-stage oligomerized olefin flow is oligomerized with a second-stage oligomerization catalyst to provide a second-stage oligomerized flow. The second-stage oligomerized flow is hydrogenated to produce fuel. A hydrogenated fuel flow for rectification is prepared using a stripping column. [Brief explanation of the drawing]

[0007] [Figure 1] This is a schematic diagram of the oligomerization section of the process and apparatus of this disclosure. [Figure 2] This is a schematic diagram of the hydrogenation section of the process and apparatus of this disclosure. [Figure 3] This is a plot of ethylene and propylene conversion over time. [Figure 4] This is a plot of product selectivity over time. [Figure 5] This is a plot of carbon loss over time for ethane, propane, and total C2-C7 paraffins. [Modes for carrying out the invention]

[0008] definition The term "communication" means that fluid flow is operably permitted between the enumerated components, and this can be characterized as "fluid communication."

[0009] The term "downstream communication" means that at least a portion of the fluid flowing to the downstream-communicating object can be operably flowed from the fluid-communicating object.

[0010] The term "upstream communication" means that at least a portion of the fluid flowing from an upstream communication object can flow operably into a fluid-communicated object.

[0011] The term "direct communication" means that the fluid flow from the upstream component enters the downstream component without passing through any other intervening vessel.

[0012] The term "indirect communication" refers to the flow of fluid from an upstream component entering a downstream component after passing through an intervening container.

[0013] The term "bypass" means that an object is removed from downstream communication with the object it is bypassing, at least to the extent that it is bypassing it.

[0014] As used herein, the terms “main” or “major” mean more than 50%, preferably more than 75%, and more preferably more than 90%.

[0015] The term “column” refers to a distillation column (singular or plural) for separating one or more different volatile components. Unless otherwise indicated, each column includes a condenser at the top of the column to condense and reflux a portion of the top flow returning to the top of the column, and a reboiler at the bottom of the column to vaporize a portion of the bottom flow and return it to the bottom of the column. The feed into the column may be preheated. The top pressure is the pressure of the top vapor at the vapor outlet of the column. The bottom temperature is the temperature at the liquid bottom outlet. The top line and bottom line 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 heating requirements and the propulsion for separation from fluidizing inert media such as vapor. A stripping column typically feeds to a top tray and removes the main product from the bottom.

[0016] 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 which may also have an aqueous outlet from a boot. A flash drum is a type of separator that can be downstream-communicated with separators that can operate at higher pressures. As used herein, the term “boiling temperature” means the atmospheric equivalent boiling point (AEBP) calculated from the observed boiling point and distillation pressure using the formulas provided in ASTM D1160 Appendix A7, entitled “Practice for Converting Observed Vapor Temperatures to Atmospheric Equivalent Temperatures”.

[0017] As used herein, the term "True Boiling Point (TBP)" refers to a test method for determining the boiling point of a substance, which corresponds to ASTM D-2892 for producing standardized quality liquefied gas, distillate fractions, and residual oil from 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 generated using 15 theoretical plates in a column with a 5:1 reflux ratio.

[0018] As used herein, the terms "T5", "T90", or "T95" mean the temperature at which 5 mass percent, 90 mass percent, or in some cases 95 mass percent of a sample boils, respectively, using ASTM D-86 or TBP.

[0019] As used herein, the term "initial boiling point (IBP)" means, as the case may be, the temperature at which a sample begins to boil using ASTM D-7169, ASTM D-86, or TBP.

[0020] As used herein, the term "end point (EP)" means, in some cases, the temperature at which all of the sample has evaporated using ASTM D-7169, ASTM D-86, or TBP.

[0021] As used herein, the term "diesel" means hydrocarbons boiling within the "diesel cut point" range, which includes an IBP of from about 125°C (257°F) to about 175°C (347°F) or a T5 of from about 150°C (302°F) to about 200°C (392°F), and a T95 of from about 343°C (650°F) to about 399°C (750°F) using the TBP distillation method or a T90 of from 280°C (536°F) to about 340°C (644°F) using ASTM D-86. The term "green diesel" means diesel comprising hydrocarbons not derived from fossil fuels.

[0022] As used herein, the term "jet fuel" means hydrocarbons boiling in the range of a T10 of about 190°C (374°F) to about 215°C (419°F) and an end point of about 290°C (554°F) to about 310°C (590°F). The term "green jet fuel" means jet fuel comprising hydrocarbons not originating from fossil fuels.

[0023] Detailed Description The disclosed process comprises oligomerizing an olefin stream comprising ethylene and / or propylene. The olefin stream may comprise larger C4 to C8 olefins. The present inventors have found that C2 to C3 alkanes can be produced in an oligomerization process that requires removal in a dealkanizer column. An olefin splitter column can be used to recover light olefins for recycle to the oligomerization process. This process and apparatus may include the oligomerization section 10 shown in Figure 1 and the hydrogenation section 110 shown in Figure 2.

[0024] Referring to the oligomerization section 10 of Figure 1, a pre-vaporized olefin stream in line 1 comprising C2 olefins is fed to a pre-separator 6. The pre-vaporized 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, which may be an olefin pre-separator, can operate at a temperature of about 16°C (60°F) to about 38°C (100°F) and a pressure of about 2.1 MPa(g) (300 psig) to about 2.8 MPa(g) (400 psig). The vapor olefin stream in line 3 can be compressed in compressor 8 to the oligomerization pressure in line 12.

[0025] A preliminary liquid olefin flow in line 2, containing C3+ olefins, is supplied to the oligomerization section 10. The preliminary olefin flow may contain C3-C8 olefins. The preliminary olefin flow in line 2 may be combined with a liquefied olefin flow in line 4, which may contain C3-C8 olefins, and supplied to the liquid supply surge drum 7. The liquid olefin flow in line 5 from the liquid supply surge drum 7, containing C3-C8 liquid olefins, may be combined with a compressed vapor olefin flow in line 11 to provide a vapor olefin flow in line 12.

[0026] The vapor and liquid olefin streams may contain substantial ethylene and propylene. The vapor and liquid olefin streams may contain mainly ethylene and / or propylene. In some embodiments, the vapor and liquid olefin streams may contain at least 95 mol% 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 in the range of C4 to C8 can be expected to be present in the feed stream. The light olefin streams may be supplied by dehydration of ethanol or from an MTO unit. The light olefin streams may be at temperatures of about 20°C (68°F) to about 150°C (302°F) and pressures of about 2.16 MPag (350 psig), preferably about 3.5 MPag (500 psig) to about 8.4 MPag (1200 psig).

[0027] The light olefin stream may first be brought into contact with a first-stage oligomerization catalyst to oligomerize ethylene and propylene into oligomers, and then brought into contact with a second oligomerization catalyst to oligomerize the ethylene and propylene that have not been converted from the first-stage oligomerization.

[0028] Oligomerization reactions generate significant heat. For example, the dimerization of ethylene can generate 612 kcal / kg (1100 BTU / lb) of heat. As a result, this large amount of heat must be managed. Therefore, the light olefin flows in lines 5 and 12 may be divided into multiple olefin flows, respectively. In Figure 1, the light olefin flows are each divided into two separate flows. The compressed vapor olefin flow in line 12 can be divided into a first vapor olefin flow in line 12a and a second vapor olefin flow in line 12b. The liquid olefin flow in line 5 is divided into a first liquid olefin flow in line 5a and a second liquid olefin flow in line 5b. The first vapor olefin flow in line 12a is mixed with the first liquid olefin flow in line 5a to provide the first input olefin flow in the first input olefin line 13a. The second vaporized olefin flow in line 12a is mixed with the second liquid olefin flow in line 5b to provide a second input olefin flow in the second input olefin line 13b. Multiple separate olefin flows, more or less than the number of other flows, can also be used. Up to six input olefin flows are readily conceivable.

[0029] The compressed vapor olefin flow in line 12 may be divided into multiple olefin flows of equal proportions in lines 12a and 12b. The liquid olefin flow in line 5 may be divided into multiple olefin flows of equal proportions in lines 5a and 5b. Alternatively, the compressed vapor olefin flow in line 12 may be divided into uneven flows. Similarly, the liquid olefin flow in line 5 may be divided into uneven flows. For example, either or both of the vapor olefin flow or the liquid olefin flow may be divided into a flow with a gradually decreasing flow rate, where the flow rate of the olefin flow fed into the preceding reactor is greater than the flow rate of the olefin flow fed into the subsequent reactor. In this embodiment, both the vapor olefin flow or the liquid olefin flow may be divided into two flows of equal flow rates, each containing 50 volume percent of the input olefin flow.

[0030] In another embodiment, the first input olefin flow in the first input olefin line 13a may constitute about 70 to about 90 volume percent of the input olefin flow, and the second input olefin flow in the second olefin line 13b may constitute about 10 to about 30 volume percent of the input olefin flow. In another embodiment, the input flows are divided in different proportions. For example, the liquid olefin flow in line 5 may be divided into two flows such that the first liquid flow in line 5a constitutes 70 to 90% of the liquid olefin flow in line 5, and the second liquid flow in line 5b constitutes 10 to 30% of the liquid olefin flow, while the compressed vapor flow in line 12 may be divided equally into two flows such that both the first vapor olefin flow in line 12a and the second olefin flow in line 12b constitute 50% of the compressed vapor olefin flow.

[0031] To control the heat generation, the input olefin stream may be diluted with a diluent stream to provide a diluted olefin stream for heat absorption. The diluent stream may include a paraffin stream in the diluent line 14. The diluent stream in the diluent line 14 may be added to the first input olefin stream in the first input olefin line 13a and then introduced into the first-stage oligomerization reactor 22. Preferably, the input olefin streams in lines 5 and 12 are divided into multiple olefin streams, and then 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 the entire first-stage oligomerization reaction. Alternatively, the diluent stream may be divided into multiple streams, each of which is added to one or more of the corresponding input olefin streams. The diluent flow may have a mass flow rate of approximately 2 to approximately 8 times, preferably approximately 3 to approximately 6 times, the combined mass flow rate of the first input olefin flow in the first input olefin line 13a and the second input olefin flow in the second input olefin line 13b.

[0032] The recycled olefin stream in the recycling line 26, which contains C4-C8 olefins, can be mixed with the input olefin stream and oligomerized in the first-stage oligomerization reactor 22. In one embodiment, the recycled olefin stream in line 26 is divided into several recycled olefin streams 26a-26d. The recycled olefin stream in the first recycled olefin line 26a can be mixed with the first input olefin stream in line 13a and fed into the first-stage oligomerization reactor 22. In a further embodiment, the first recycled olefin stream in the first recycled 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.

[0033] The first diluted olefin stream may contain 50% by weight or less of olefin, preferably 30% by weight or less of olefin, and more preferably 20% by weight or less of olefin. In embodiments, the first diluted olefin stream contains about 10 to about 35% by weight of C2 to C8 olefins. The first diluted olefin stream may contain 50% by weight or less of ethylene, preferably 25% by weight or less of ethylene, and more preferably 20% by weight or less of ethylene. In embodiments, the first diluted olefin stream contains about 10 to about 20% by weight of propylene. The first diluted olefin stream may contain 50% by weight or less of propylene, preferably 25% by weight or less of propylene, and more preferably 20% by weight or less of propylene. In embodiments, the first diluted olefin stream contains about 10 to about 20% by weight of propylene.

[0034] The first-stage oligomerization reactor 22 may comprise a series of first-stage oligomerization catalyst beds 22a, 22b, 22c, and 22d, each for which an olefin feed stream is fed. Preferably, the first-stage oligomerization reactor 22 includes four fixed first-stage oligomerization catalyst beds 22a, 22b, 22c, and 22d. 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 reaction vessels. Up to six first-stage oligomerization catalyst beds are readily conceivable. In Figure 1, two first-stage oligomerization reaction vessels 21a and 21b are utilized.

[0035] A parallel first-stage oligomerization reactor may be used when the first-stage oligomerization reactor 22 is deactivated, during which time the first-stage oligomerization reactor 22 is regenerated in situ by the combustion of coke from the catalyst. In another embodiment, each first-stage oligomerization reactor may include a reed reactor, a lag reactor, and a pre-reactor to facilitate regeneration. Figure 1 shows only two reaction vessels 21a and 21b.

[0036] The diluted first input olefin stream in line 16a may be cooled in the first input cooler 18a to provide the cooled and diluted first input olefin stream in line 20a, which is then introduced into the first bed 22a of the first-stage oligomerization catalyst in the first-stage oligomerization reaction vessel 21a of the first-stage oligomerization reactor 22. The cooled and diluted first input olefin stream in line 20a may be introduced at a temperature of about 180°C (356°F) to about 260°C (500°F) and a pressure of about 3.5 MPag (500 psig) to about 8.4 MPag (1200 psig). The input cooler 18a may be equipped with a steam generator.

[0037] The diluted first input olefin stream can preferably be introduced into the first first stage catalyst bed 22a in line 20a in a downward flow operation. However, upward flow operation may be preferable. The diluted first input olefin stream is in a mixed vapor-liquid phase in which the vapor phase mainly consists of ethylene. As oligomerization of ethylene, propylene, and recycled olefins occurs in the first first stage oligomerization catalyst bed 22a, heat is generated due to the high exothermic nature of the olefin oligomerization reaction. Once the first input olefin stream is oligomerized, a first oligomerized effluent flow is generated in the first oligomerized effluent line 24a at an elevated outlet temperature despite cooling and dilution. The elevated outlet temperature is limited to 150°C (302°F) to approximately 260°C (500°F).

[0038] The second input olefin stream in line 13b may be mixed with the second recycled olefin stream in the second recycled olefin line 26b and the first oligomerized effluent stream in the first oligomerized effluent line 24a, which is removed from the first oligomerized catalyst bed 22a in the first first stage reactor 21a, to provide a mixed second input olefin stream in line 16b. The first oligomerized effluent stream in line 24a includes a diluent stream from diluent line 14 that is added to the first input olefin stream in line 13a. The second input olefin stream may contain 35% by weight or less of C2-C8 olefins, preferably 25% by weight or less of C2-C8 olefins, and more preferably 20% by weight or less of C2-C8 olefins. The second input olefin stream may contain 30% by weight or less of ethylene, preferably 25% by weight or less of ethylene, and more preferably 20% by weight or less of ethylene. The second input olefin stream may contain 30% by weight or less of propylene, preferably 25% by weight or less of propylene, and more preferably 20% by weight or less of propylene. The second mixed input olefin stream in line 16b may be cooled in a second input cooler 18b, which may be located outside the first first-stage oligomerization reactor 21a, to provide a cooled second input olefin stream in line 20b, which can then be introduced into the second bed 22b of the first-stage oligomerization catalyst in the first first-stage oligomerization reactor 21a. The input cooler 18b may include a steam generator.

[0039] A second cooled input olefin stream can be introduced into line 20b at a temperature of approximately 180°C (356°F) to approximately 230°C (446°F) and a pressure of approximately 3.5 MPag (500 psig) to approximately 8.4 MPag (1200 psig). The second cooled input olefin stream contains a diluent and olefins from the first oligomerization stream. The diluted second input olefin stream is in a mixed vapor-liquid phase in which the vapor phase mainly consists of ethylene. The olefins from the first oligomerization stream are oligomerized in the second first-stage catalyst bed 22b. The oligomerization of ethylene, propylene, recycled olefins, and oligomers in the second olefin stream within the second first-stage oligomerization catalyst bed 22b generates a second oligomerized olefin effluent stream into the second oligomerized effluent line 24b at an elevated outlet temperature. The increased outlet temperature can be limited to a temperature approximately 30°C (54°F) to 50°C (90°F) higher than the inlet temperature to the catalyst bed 22b.

[0040] The second oligomerization outflow in line 24b, removed from the second first-stage oligomerization catalyst bed 22b in the first first-stage reaction vessel 21a, can be mixed with the third recycled olefin flow in the third recycled olefin line 26c to provide the first recycled olefin input flow in line 16c. In this embodiment, neither the first input olefin flow in line 13a nor the second input olefin flow in line 13b is directly added to the first recycled olefin input flow in line 16c. Alternatively, a portion of the input olefin flows in lines 13a and 13b may be added to the second oligomerization outflow together with the first recycled olefin input flow in line 16c. The second oligomerization outflow in line 24a includes a diluent flow from diluent line 14 added to the first input olefin flow in line 13a. The first recycled olefin feed stream in line 16c may contain 30% by weight or less of ethylene, preferably 25% by weight or less of ethylene, and more preferably 20% by weight or less of ethylene. The first recycled olefin feed stream may contain 30% by weight or less of propylene, preferably 25% by weight or less of propylene, and more preferably 20% by weight or less of propylene. The first recycled olefin feed stream in line 16c may contain 30% by weight or less of C2-C8 olefins, preferably 25% by weight or less of C2-C8 olefins, and more preferably 20% by weight or less of C2-C8 olefins. The first recycled olefin feed stream in line 16c may be cooled in a third feed cooler 18c located outside the oligomerization reactor 22 to provide a cooled first recycled olefin feed stream in line 20c, which can then be fed into the third bed 22c of the first-stage oligomerization catalyst in the first-stage oligomerization reactor 22. In this embodiment, the third bed 22c of the first-stage oligomerization catalyst is provided in the second first-stage oligomerization reaction vessel 21b. The input cooler 18c may include a steam generator.

[0041] The cooled first recycled olefin feed stream in line 20c may be fed at a temperature of approximately 180°C (356°F) to approximately 230°C (446°F) and a pressure of approximately 3.5 MPag (500 psig) to approximately 8.4 MPag (1200 psig). The first recycled olefin feed stream includes a diluent and olefins from the second oligomerized olefin stream and the third recycled olefin stream. The olefins are oligomerized in the third catalyst bed 22c. The oligomerization of ethylene and propylene in the third bed 22c of the first-stage oligomerizing catalyst, as well as the oligomerization of oligomers in the first recycled olefin feed stream, generates a third oligomerized effluent flow in the third oligomerized effluent line 24c at an elevated outlet temperature. In this embodiment, the third oligomerized effluent logistics is the second-to-last oligomerized effluent logistics, and the third oligomerized effluent line 24c is the second-to-last oligomerized effluent line 24c. The elevated outlet temperature is limited to a temperature approximately 30°C (54°F) to approximately 50°C (90°F) higher than the inlet temperature to the catalyst bed 22c.

[0042] The third oligomerized effluent in line 24c, removed from the second first-stage oligomerization reaction vessel 21b of the first-stage oligomerization reactor 22, can be mixed with the fourth recycled olefin stream in line 26d to provide the second recycled olefin feed stream in line 16d. The third oligomerized effluent in line 24a includes a diluent stream from diluent line 14 added to the first olefin stream in line 13a. Neither of the feed olefin streams in lines 13a and 13b are directly added to the second recycled olefin feed stream in line 16d. In embodiments, the third oligomerized effluent in line 24c can also be mixed with a portion of the feed olefin streams in lines 13a and 13b and oligomerized with them. The second recycled olefin feed stream may contain 35% by weight or less of C2-C8 olefins, preferably 30% by weight or less of C2-C8 olefins, and preferably 25% by weight or less of C2-C8 olefins. The second recycled olefin feed stream may contain 30% by weight or less of ethylene, preferably 25% by weight or less of ethylene, and more preferably 20% by weight or less of ethylene. The second recycled olefin feed stream may contain 30% by weight or less of propylene, preferably 25% by weight or less of propylene, and more preferably 20% by weight or less of propylene. The second recycled olefin feed stream in line 16d is cooled in a fourth feed cooler 18d, which may be located outside the second vessel 21b of the first-stage oligomerization reactor 22, to provide the cooled second recycled olefin feed stream in line 20d, which can then be fed into the fourth bed 22d of the first-stage oligomerization catalyst in the second vessel of the first-stage oligomerization reactor 22. The feed cooler 18d may include a steam generator.

[0043] The cooled second recycled olefin feed stream in line 20d can be fed at a temperature of approximately 180°C (356°F) to approximately 230°C (446°F) and a pressure of approximately 3.5 MPa(g)(500 psig) to approximately 8.4 MPa(g)(1200 psig). The cooled second recycled olefin feed stream in line 20d contains a diluent, and olefins from the third or second-to-last oligomerized effluent stream, and C4-C8 olefins from the fourth recycled olefin stream. The olefins are oligomerized on the fourth catalyst bed 22d. The oligomerization of ethylene and propylene in the second recycled olefin feed stream on the fourth bed 22d of the first-stage oligomerization catalyst generates a fourth oligomerized stream in the fourth oligomerized effluent line 24d at an elevated outlet temperature. The elevated outlet temperature is limited to a temperature approximately 30°C (54°F) to 50°C (90°F) higher than the inlet temperature to the catalyst bed 22d.

[0044] The fourth oligomerized effluent in line 24d exits the second reaction vessel 21b of the first-stage oligomerization reactor 22. In this embodiment, the fourth oligomerized effluent in line 24d is the final oligomerized effluent, and the fourth oligomerized effluent line 24d is the final oligomerized effluent line 24d.

[0045] The first stage of the oligomerization reaction takes 0.5 to 10 hours based on olefins. -1 In WHSV, this occurs mainly in the liquid phase or gas-liquid mixed phase. The inventors have found that across the first-stage oligomerization catalyst bed, typically 10 to 50% by weight of ethylene in the olefin stream is converted to higher olefins. The ethylene is first dimerized on the catalyst to butene. The majority of the propylene and butene in the olefin stream introduced into the first-stage oligomerization catalyst bed is oligomerized. In embodiments, at least 99 mol% of the propylene and butene in the olefin stream are oligomerized.

[0046] The first-stage oligomerization catalyst may include a zeolite catalyst. The first-stage oligomerization catalyst can be considered as a solid acid catalyst. The zeolite may constitute about 5 to about 95% by weight of the catalyst, for example, about 5 to about 85% by weight. Suitable zeolites include those having structures 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. The three-letter codes indicating the zeotype are as defined by the Structure Commission of the International Zeolite Association and are managed 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 include a zeolite having a skeleton with a 10-membered ring pore structure. Examples of suitable zeolites having a 10-membered ring pore structure include TON, MTT, MFI, MEL, AFO, AEL, EUO, and FER. In a further preferred embodiment, the first-stage oligomerization catalyst, which includes a zeolite having a 10-membered ring pore structure, may include a one-dimensional pore structure. The one-dimensional pore structure represents a zeolite containing non-crossing pores substantially parallel to one of the crystal axes. The pores preferably extend through the zeolite crystal. A suitable example of a zeolite having a 10-membered ring one-dimensional pore structure is MTT. In a further embodiment, the first-stage oligomerization catalyst includes an MTT zeolite.

[0047] The first-stage oligomerization catalyst can be formed by combining a zeolite with a binder, and then forming the catalyst into a pellet. The pellet may be optionally treated with a phosphorus reagent to produce a zeolite having 0.5 to 15% by weight of phosphorus components of the treated catalyst. The binder is used to impart hardness and strength to the catalyst. Examples of binders include alumina, aluminum phosphate, silica, silica-alumina, zirconia, titania, and combinations thereof 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 clay.

[0048] One of the components of the catalyst binder used in the present invention is alumina. The alumina source can be any of various hydrated 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. Preferred alumina is available from UOP LLC under the trademark VERSAL. Preferred alumina is available from Sasol North America Alumina Product Group under the trademark CATAPAL. This material is extremely high-purity alpha-alumina monohydrate (pseudo-boehmite) and has been shown to yield high-purity gamma-alumina after calcination at high temperatures.

[0049] A suitable first-stage oligomerization catalyst is prepared by mixing proportionally volume amounts of zeolite and alumina to achieve a desired zeolite-alumina ratio. In embodiments, the MTT content may be about 5 to about 85% by weight, for example, about 20 to about 82% by weight of MTT zeolite, with the remaining alumina powder providing a suitably supported catalyst. Silica supports are also considered.

[0050] A monobasic acid such as nitric acid or formic acid may be added to the mixture in an aqueous solution to dissolve the alumina in the binder. Additional water may be added to the mixture to provide sufficient wettability to form a dough with sufficient viscosity for extrusion or spray drying. Extrusion aids such as cellulose ether powder may also be added. Preferred extrusion aids are available from The Dow Chemical Company under the trademark Methocel.

[0051] The paste or dough may be prepared in the form of molded particles, a preferred method of which is to extrude the dough through a die having an opening of the desired size and shape, and then divide the extruded material into extruders of the desired length and dry them. Further calcination steps may be used to provide additional strength to the extruders. Generally, calcination is carried out in an airflow at a temperature of about 260°C (500°F) to about 815°C (1500°F). The MTT catalyst does not have the selectivity to neutralize acidic sites such as amines.

[0052] Extruded particles may have any suitable cross-sectional shape, i.e., symmetric or asymmetric, but in most cases have a symmetrical cross-sectional shape, preferably spherical, cylindrical, or multi-lobed. The cross-sectional diameter of the particles can be as small as about 40 μm. However, it is typically about 0.635 mm (0.25 inches) to about 12.7 mm (0.5 inches), preferably about 0.79 mm (1 / 32 inch) to about 6.35 mm (0.25 inches), and most preferably about 0.06 mm (1 / 24 inch) to about 4.23 mm (1 / 6 inch).

[0053] In one exemplary embodiment, an MTT-type zeolite catalyst, arranged on a high-purity pseudo-boehmite alumina substrate in a ratio of approximately 90 / 10 to approximately 20 / 80, preferably approximately 20 / 80 to approximately 50 / 50, is provided to the catalyst bed, or more often, to the first-stage oligomerization reactor 22.

[0054] The first-stage oligomerization catalyst can be regenerated when it becomes inactive. Preferred regeneration conditions include, for example, exposing the first-stage oligomerization catalyst to hot air at approximately 400°C to 500°C in situ. To promote regeneration without downtime, a rocking bed apparatus may be used in conjunction with an alternative first-stage oligomerization reactor. The regeneration gas stream may be introduced into the first-stage oligomerization reactor 22 that requires regeneration. The regeneration gas may contain air with increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to those of a fresh catalyst.

[0055] Zeolite catalysts are advantageous as first-stage oligomerization catalysts. Zeolite catalysts have relatively low susceptibility to oxygenate contamination. As a result, when produced from an alcohol dehydration process, the degree of oxygenate removal required for the olefin feedstock in line 1 is reduced.

[0056] The final first-stage oligomerization flow in the last first-stage oligomerization effluent line 24d has increased concentrations of ethylene and propylene oligomers compared to the light olefin flows in lines 5 and 12. The final first-stage oligomerization flow in the last first-stage oligomerization effluent line 24d is cooled by steam generation in the steam generator 18e or by other heat exchange, further cooled by heat exchange with the second-stage oligomerization flow in line 34, and possibly further cooled by a steam generator, etc., to provide the input first-stage oligomerization flow, which is fed into the second-stage oligomerization reactor 32 in the second-stage oligomerization input line 28. To achieve the most desirable olefin product, the second-stage oligomerization reactor 32 operates at a temperature of approximately 80°C (176°F) to approximately 200°C (392°F). The second-stage oligomerization reactor 32 is operated at a pressure of approximately 2.1 MPa (300 psig) to approximately 7.6 MPa (1100 psig), more preferably approximately 3.5 MPa (500 psig) to approximately 6.9 MPa (1000 psig). The oligomerization feed stream in the second stage is oligomerized in the mixed steam-liquid phase, mainly to C4+ olefins.

[0057] 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 a downward flow manner. However, an upward flow manner may be preferable. The second-stage oligomerization feed stream is brought into contact with the second-stage oligomerization catalyst, dimerizing and trimerizing the unconverted ethylene from the first-stage oligomerization reactor 22, while also dimerizing, trimerizing, and tetramerizing the higher olefins to provide olefins in the distillate range. With respect to 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, yield desired 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% by weight of the ethylene in the second-stage oligomerized input logistics is usually converted to butene.

[0058] The second-stage oligomerization reactor 32 may include a first reaction vessel 31a containing a first bed 32a of the second-stage oligomerization catalyst and a second reaction vessel 31b containing a second bed 32b of the second-stage oligomerization catalyst. The first second-stage oligomerization stream is discharged from the first second-stage reaction vessel 31a, cooled, and introduced into the second second-stage reaction vessel 31b. A second-stage oligomerization stream having a larger average carbon number than the input second-stage oligomerization stream in line 28 exits the second-stage oligomerization reactor 32 in line 34.

[0059] The first-stage oligomerization reactor 22 and the second-stage oligomerization reactor 32 utilize a steam-liquid distribution tray to mix and disperse ethylene vapor with liquid olefin and liquid paraffin, thereby promoting heat transfer and controlling heat generation.

[0060] 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 by a Group VIB metal. Typically, since silica and alumina are present only in the base, the ratio of silica to alumina is the same for both the catalyst and the base. The metal can be impregnated onto the silica-alumina base or ion-exchanged in the silica-alumina base. Co-mulling is also considered. The catalysts of the present invention may have a low-temperature acidity ratio of at least about 0.15, preferably about 0.2, and preferably about 0.25 or greater, as determined by Ammonia Temperature Programmed Desorption (Ammonia TPD) as described below. In addition, preferred catalysts have a low-temperature acidity ratio of about 50 to about 400 m as determined by the nitrogen BET method. 2 It has a surface area of ​​ / g.

[0061] A preferred second-stage oligomerization catalyst comprises an amorphous silica-alumina support. One of the components of the catalyst support used in the present invention is alumina. The alumina can be any of various hydrated aluminum oxides or alumina gels, such as alpha-alumina monohydrate with a boehmite or pseudoboehmite 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 trademark CATAPAL. This material is an extremely high-purity alpha-alumina monohydrate (pseudoboehmite) and has been shown to yield high-purity gamma-alumina after calcination at high temperatures. Another component of the catalyst support is amorphous silica-alumina. A preferred silica-alumina with a silica-to-alumina ratio of 2.6 is available, for example, from CCIC, a subsidiary of JGC in Japan.

[0062] Another component used in the preparation of the second-stage oligomerization catalyst used in the present invention is a surfactant. The surfactant is preferably mixed with the above-mentioned alumina and silica-alumina powders. The resulting mixture of surfactant, alumina, and silica-alumina 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 according to the present invention. Any suitable surfactant can be used according to the present invention. Preferred surfactants are selected from a range of commercially available surfactants marketed by Solvay SA under the trademark "Antarox". "Antarox" surfactants are generally characterized as modified linear aliphatic polyethers and are low-foaming biodegradable detergents and wetting agents.

[0063] A suitable silica-alumina mixture is prepared by mixing proportional volumes of silica-alumina and alumina to achieve a desired silica-to-alumina ratio. In one embodiment, amorphous silica-alumina with a silica:alumina ratio of 2.6 at about 75–99% by weight and alumina powder at about 10–20% by weight provides a suitable carrier. In another embodiment, other ratios of amorphous silica-alumina to alumina may be preferred.

[0064] Any convenient method can be used to incorporate the surfactant into the mixture of silica-alumina and alumina. Preferably, the surfactant is mixed during the mixing and formation of alumina and silica-alumina. A preferred method is to mix an aqueous solution of the surfactant with the alumina and silica-alumina blend before the final formation of the support. Preferably, the surfactant is present in the paste or dough in an amount of about 0.01 to about 10% by weight, based on the weight of alumina and silica-alumina.

[0065] A monobasic acid such as nitric acid or formic acid may be added to the mixture in an aqueous solution to dissolve the alumina in the binder. Additional water may be added to the mixture to provide sufficient wettability to form a dough with sufficient viscosity for extrusion or spray drying.

[0066] The paste or dough can be prepared in the form of molded particles. A preferred method involves extruding a dough mixture of alumina, silica-alumina, a surfactant, and water through a die having an opening of the desired size and shape, then dividing the extruded material into extruders of the desired length and drying them. A further calcination step may be used to provide additional strength to the extruders. Generally, 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).

[0067] Extruded particles may have any suitable cross-sectional shape, i.e., symmetric or asymmetric, but in most cases have a symmetrical cross-sectional shape, preferably spherical, cylindrical, or multi-lobed. The cross-sectional diameter of the particles can be as small as about 40 μm. However, it is typically about 0.635 mm (0.25 inches) to about 12.7 mm (0.5 inches), preferably about 0.79 mm (1 / 32 inch) to about 6.35 mm (0.25 inches), and most preferably about 0.06 mm (1 / 24 inch) to about 4.23 mm (1 / 6 inch).

[0068] Typical properties of the amorphous silica-alumina supports used herein are total pore volume, average pore diameter, and surface area large enough to provide substantial space and area for depositing active metal components. The total pore volume of the support, as measured by conventional mercury porosimeters, 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 pore volume of the support in pores larger 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 BET, is typically 50 m². 2 / grams or more, for example, about 200m 2 / grams or more, preferably at least 250m 2 / gram, most preferably about 300ml 2 Grams ~ approximately 400m 2 It is / grams.

[0069] To prepare the second-stage oligomerization catalyst, the carrier material is compounded 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 carrier particles. The Group VIII metal, preferably nickel, must be present at a concentration of about 0.5 to about 15% by weight, and the Group VIB metal, preferably tungsten, must be present at a concentration of about 0 to about 12% by weight. Impregnation can be achieved by any method known in the art, for example, by spray impregnation, in which a solution containing the metal precursor in a dissolved form is sprayed onto the carrier particles. Another method is a multi-dip procedure, in which the carrier material is repeatedly brought into contact with the impregnation solution with or without intermittent drying. Yet another method includes immersing the carrier in a large volume of impregnation solution or circulating the carrier therein, and yet another method is a pore volume or pore saturation technique, in which the carrier particles are introduced into an impregnation solution of just enough volume to fill the pores of the carrier. In some cases, the pore saturation technique may be modified to utilize an impregnation solution having a volume between approximately 10 percent less and approximately 10 percent more than the volume required to just fill the pores.

[0070] When the active metal precursor is incorporated by impregnation, the metal is converted to its respective oxide form by subsequent or second calcination at a high temperature, for example, 399°C (750°F) to 760°C (1400°F). In some cases, calcination may be performed after each impregnation of the individual active metals. Subsequent calcination yields a catalyst containing the active metals in their respective oxide forms.

[0071] A preferred second-stage oligomerization catalyst of the present invention has an amorphous silica-alumina substrate impregnated with about 0.5 to about 15% by weight of nickel in the form of a 3.175 mm (0.125 inch) extruded material, having a density of about 0.45 to about 0.65 g / mL. It is also intended that the metal may be incorporated onto the support by other methods such as ion exchange and co-miscibility.

[0072] The second-stage oligomerization catalyst can be regenerated when deactivated. Suitable regeneration conditions include, for example, exposing the catalyst to hot air at approximately 400-500°C in situ. To facilitate regeneration without downtime, a rocking bed apparatus may be used in conjunction with an alternative second-stage oligomerization reactor. The regeneration gas may contain air with increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to that of a fresh catalyst.

[0073] The second-stage oligomerization reaction is also inherently exothermic. The final oligomerized olefin stream in line 24d includes a diluent stream from diluent line 14, which is added to the first input olefin stream in the first input olefin line 13a and carried through the first-stage oligomerization catalyst beds 22a-22d. The diluent stream is then transported in line 28 to the second-stage oligomerization reactor 32 to absorb the heat generated in the second-stage oligomerization reactor. A dedicated diluent line to the second-stage oligomerization reactor 32 is also intended for rapid control of the heat rise or for cooling the second-stage oligomerization reactor 32.

[0074] When the oligomerization reaction is carried out according to the above process conditions, a C4 olefin conversion rate of approximately 95% or more, or 97% or more, is achieved. The second-stage oligomerization stream obtained in line 34 contains multiple olefin products, which are hydrocarbons in the distillate range.

[0075] The oligomerized olefin flow in line 34, which has a higher C8+ olefin concentration compared to the second-stage oligomerized flow in line 28, is subjected to heat exchange with the first-stage oligomerized flow in line 24d to reduce its pressure, and then subjected to heat exchange with the olefin splitter column bottom flow in line 30 before being supplied to column 40 of the dealkane column. The oligomerized olefin flow in line 34 has a temperature of approximately 160°C (320°F) to approximately 190°C (374°F) and a pressure of approximately 3.9 MPa (gauge) (550 psig) to approximately 7 MPa (gauge) (1000 psig).

[0076] The inventors have found that light alkanes such as ethane and / or propane are produced in the first-stage oligomerization reactor 22 and / or the second-stage oligomerization reactor 32, and that these must be removed from the second-stage oligomerization reactor for fuel production, particularly to facilitate the recycling of light olefins to the first-stage oligomerization reactor 22. Light alkanes are inert and accumulate in the recycling loop. Therefore, the second-stage oligomerization stream in line 34 is dealkaneized by rectification in the dealkane column 40 to provide a light alkane stream and a dealkaneized stream. In one embodiment, the light alkane stream is an ethane stream, in which case the dealkane column 40 is a deethane column. In another embodiment, the light alkane stream is a propane stream, in which case the dealkane column 40 is a depropane column. The light alkane stream may contain ethane and / or propane, or it may be a mixture of ethane and propane.

[0077] In the dealkane column 40, light alkanes such as C3- and preferably C2- hydrocarbons are separated from the dealkane bottom flow in the bottom line 44, which probably contains C4+ and preferably C3+ hydrocarbons, into the light alkane top flow in the top line 42. Olefins can be recycled from the dealkane top flow in the top line 42 into the first-stage oligomerization reaction product 22. When operating as a dealkane column, the dealkane column 40 may operate at a bottom temperature of about 177°C (-350°F) to about 302°C (575°F) and a top pressure of about 207-kPa (gauge) (30 psig) to about -690kPa (gauge) (100 psig). When operating as a dealkane tower, tower 40 can operate at a bottom temperature of approximately 194°C (-381°F) to approximately 333°C (630°F) and a top pressure of approximately 207 kPa (gauge) (30 psig) to approximately 1.38 MPa (gauge) (200 psig).

[0078] The light alkane top flow in the top line 42 is cooled and separated in the dealkane receiver 46 and supplied to the off-gas line 47, where it is cooled and can be fed for further processing, such as being removed as combustion gas in line 48 along with the net vapor flow in the receiver top line 68. The condensate from the dealkane receiver 46 may be refluxed back to the column 40 of the dealkane column in the dealkane top liquid line 49. The dealkane off-gas flow can be used as fuel to provide a heating load in process 10. In embodiments, a portion of the condensate from the dealkane receiver 44 in line 49 may be taken up by the first-stage oligomerization reactors in lines 72 and 26 as olefin recycling in line 51.

[0079] The dealkane flow in the bottom line 44 may be divided into a reboiling flow in line 50, which is reboiled by a heat exchanger with a first high-temperature diesel flow in line 52, which is taken from the bottom heat exchanger flow of the jet rectification column in the jet bottom heat exchanger line 74, and a net bottom flow in line 54, which is supplied directly to the olefin splitter column 60 without heating. The reboiled bottom flow in line 50 may be returned boiling to the column 40 of the dealkane column to provide heating requirements. In another embodiment, the feed to the column 40 of the dealkane column is not preheated by the olefin splitter bottom flow in line 30, but the feed to the olefin splitter column 60 in the net bottom line 54 is preheated by heat exchange with the olefin splitter bottom flow in line 64.

[0080] The dealkane flow in the net bottom line 54 of the dealkane column is divided by rectification in the olefin splitter column 60 into a light olefin flow, possibly in the olefin splitter top line 62, and possibly in the heavy olefin flow, possibly in the olefin splitter bottom line 64. The olefins can be recycled from the olefin splitter top flow in the olefin splitter top line 62 to the first-stage oligomerization reaction 22. The olefin splitter top flow may be cooled to approximately 19°C (66°F) to approximately 93°C (200°F), and the resulting condensate is refluxed from the olefin splitter receiver 66 back to the olefin splitter column 60. The net vapor flow from the olefin splitter receiver 66 in the receiver top line 68 can be further processed as fuel gas in line 48, etc., together with the off-gas flow in the off-gas line 47. Light olefin condensates from the bottom of the olefin splitter receiver in line 70 can be divided into a reflux flow that is refluxed back into the column in line 71 and a light olefin recycling flow in recycling line 72 that can be recycled to the first-stage oligomerization reactor 22 or the second-stage oligomerization reactor 32. The light olefin flow in line 72 may constitute about 1 to about 15% by weight or perhaps the majority of the light olefin flow in line 70. The light oligomer flow in line 72 may contain about 40 to about 80% by weight of C4-C8 olefins. In embodiments, the light olefin flow in line 72 may be flashed in a knockout drum 75 to remove vapor from the light olefin vapor flow that can be transported to the hydrogenation section in the top column line 77, and the liquid recycled olefin oligomer flow in line 26 may be recycled to the first-stage oligomerization reactor 22 to oligomerize the C4-C8 olefins.

[0081] The heavy olefin flow in the splitter bottom line 64 may be split between a reboiling flow in the splitter reboiling line 65, which is reboiled by heat exchange with a second high-temperature diesel flow in line 73, which is taken from the bottom heat exchange flow of the jet rectification column in the jet bottom heat exchange line 74 in Figure 2 and returned to the olefin splitter column 60. The cooled second high-temperature diesel flow in line 112 is returned to the hydrogenation section 110 in Figure 2 for reboiling and then returned to the jet rectification column 100. The heavy olefin flow in the net bottom line 30 is cooled by heat exchange with the second-stage oligomerized olefin flow in line 34 and then transported to a hydrogenation section (not shown). Purge of the heavy olefin flow can be done in line 33. The heavy olefin flow contains C9+ olefins, which, once cooled, can be transported to the hydrogenation section 110.

[0082] Turning to the hydrogenation section 110 in Figure 2, the heavy olefin stream in the net olefin splitter bottom line 30 from Figure 1, containing C9+ oligomerized olefins in the distillate range, can be hydrogenated in the hydrogenation reactor 80 to saturate the olefin bonds and provide fuel. This process 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). Furthermore, hydrogenation of the oligomerized heavy olefin yields 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 generate steam, which can be combined with the light olefin liquid stream containing C2-C8 olefins in line 77 in Figure 1 to produce a composite olefin stream in line 79. Alternatively, the composite olefin flow in line 79 may be combined with the hydrogen flow in line 76 to provide a composite hydrogenation feed stream in line 81, which is cooled and fed into the hydrogenation reactor 80 at 125°C (257°F) to approximately 204°C (400°F) and 2.8 MPa (400 psig) to approximately 6.9 MPa (1000 psig). Excess hydrogen, such as approximately 1.5 to approximately 5.0 stoichiometric hydrogen, can be used to ensure complete saturation.

[0083] Hydrogenation is typically carried out using conventional hydrogenation or hydrogenation 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. Catalyst supports may be oxides of silica, alumina, titania, calcium carbonate, barium sulfate, and carbon, but may be any solid inert material. Catalyst supports may be in the form of powders, granules, pellets, etc.

[0084] In exemplary embodiments, hydrogenation is carried out in a hydrogenation reactor 80 containing an alumina-supported platinum catalyst, for example, about 0.1% to about 2% by weight, preferably about 0.5% to about 0.9% by weight. In another embodiment, the hydrogenation catalyst contains about 5% to about 30% by weight of nickel catalyst. The hydrogenation reactor 80 converts the olefin into a paraffin product having the same carbon number distribution as the olefin, thereby forming paraffins with a distillate range suitable for use as jet and diesel fuels.

[0085] The hydrogenated heavy mass flow discharged from the hydrogenation reactor 80 to line 83 can be separated in a high-temperature separator 82 that provides hydrocarbon splits. In the high-temperature separator 82, the hydrogenated heavy mass flow is separated into a high-temperature hydrogenated steam flow in the top line 84 and a high-temperature hydrogenated liquid flow in the high-temperature separator bottom line 86. The hydrogenated heavy mass liquid flow in the bottom line 86 may be heated by heat exchange with the diluent flow in line 14 before the diluent flow is recycled to the first-stage oligomerization reactor 22 in Figure 1. The heated hydrogenated heavy mass liquid flow in the high-temperature bottom line 86 may be supplied to the stripper column 90. The high-temperature separator can operate at bottom temperatures of approximately 204°C (400°F) to approximately 343°C (650°F) and pressures of approximately 2.8 MPa (400 psig) to approximately 6.9 MPa (1000 psig).

[0086] The high-temperature hydrogenation vapor flow in the high-temperature top line 84 can be cooled and supplied to the cryogenic separator 88. The cryogenic separator separates the cooled high-temperature hydrogenation vapor flow in the high-temperature top line 84 into a low-temperature steam hydrogenation flow in the cryogenic top line 87 and a low-temperature heavy hydrogenation liquid flow in the cryogenic bottom line 89. The purge flow in the purge line 85 can be taken from the low-temperature steam hydrogenation flow in the cold-temperature top line 87, and the remainder can be compressed and combined with the feed hydrogen in line 88 to provide a hydrogen flow in line 76. The low-temperature heavy hydrogenation liquid flow in the bottom line 89 may be supplied to the stripping column 90 at a supply position higher than the supply position of the high-temperature heavy hydrogenation liquid flow in the high-temperature separator bottom line 86. The cryogenic separator can operate at temperatures ranging from approximately 32°C (90°F) to approximately 71°C (150°F) and pressures ranging from approximately 2.8 MPa (400 psig) to approximately 4.5 MPa (650 psig).

[0087] The stripping column 90 may be a flash stripper for removing light gases from the high-temperature hydrogenated liquid flow in the high-temperature bottom line 86 and the low-temperature hydrogenated liquid flow in the low-temperature bottom line 89. Both of these flows may be supplied to the stripping column 90 together, or separately as shown in the figure. The stripping column 90 removes residual light gases from the liquid hydrogenated flow and provides the stripper top flow to the stripper top line 92 and the stripped bottom flow to the stripper bottom line 94. The stripper top flow in the stripper top line 92 is cooled and separated in the stripper receiver 96 to provide the stripper off-gas flow in the stripper receiver top line 97 and the condensed flow in line 98 which is returned to the column. The stripping tower 90 can operate at a bottom temperature of approximately 232°C (450°F) to approximately 316°C (600°F) and a top pressure of approximately 207 kPa (30 psig) to approximately 689 kPa (100 psig).

[0088] After being stripped in the stripping column 90 to remove volatiles, the stripped fuel flow in the stripper column bottom line 94 may be supplied to the jet rectification column 100 without further heating. Alternatively, the stripping column 90 upstream of the jet rectification column 100 may be omitted. In the jet rectification column 100, the stripped fuel flow can be separated into a jet-off gas flow in the top line 102, a green jet flow in the side line 104 from the side of the jet rectification column 100, and a green diesel flow in the bottom line 106. The jet rectification column 100 can operate at a bottom temperature of approximately 288°C (550°F) to approximately 400°C (750°F) and a top pressure of approximately 35kPa (5psig) to approximately 350kPa (50psig).

[0089] The jet rectification top flow in the top line 102 may be cooled, and a portion of the resulting condensate may be refluxed from the jet rectification receiver 108 to the jet rectification column 100 in the jet rectification top liquid line 109, while the remaining condensate is taken into line 119 as naphtha range product flow. Furthermore, a net off-gas flow containing C4-C8 hydrocarbons is taken into the receiver top line 105 from the jet rectification receiver 108. The majority of the hydrocarbons in the net off-gas flow in the receiver top line 105 are lighter hydrocarbons and can be used to fuel the reboiling heater 116 for the jet rectification column 100.

[0090] The green jet stream, taken into the tower-side line 104, contains C8-C18 hydrocarbons in the kerosene range and can be cooled and extracted as a jet fuel product that meets the applicable SPK standards. In an alternative embodiment, the green jet stream can be extracted from the condensate flow in line 109 from the jet rectifier receiver 108, instead of refluxing all of the condensate back into the tower. This green jet stream extracted from line 109 may need to be further stripped to remove light fractions. In such an embodiment, the tower-side line 104 for recovering the green jet fuel stream is not provided.

[0091] The green diesel bottom flow in bottom line 106 can be split into a reboiling diesel flow in line 107 and a diesel product flow in line 114. The reboiling diesel flow in line 107 can be split into a jet bottom heat exchange flow 74 and a bypass bottom flow in bottom bypass line 111. As shown in Figure 1, the jet bottom heat exchange flow in jet bottom heat exchange line 74 can be split into a first high-temperature diesel flow in line 52 and a second high-temperature diesel flow in line 73 to provide reboiling heat to the dealkane column 40 and the olefin splitter column 60, respectively. The bypass bottom flow in line 111 is taken into the jet reboiling line 113 through a valve above it, or a portion of the cooled second high-temperature diesel flow in line 112, or both, is taken into the jet reboiling line 113, reboiled in the furnace 116, and returned to the jet rectification column 100.

[0092] The diesel product stream in line 114 is divided into a diesel product stream in diesel product line 118 and a diluent stream in line 14. The diluent stream in line 14 is cooled by heat exchange with the high-temperature hydrogenated heavy liquid stream in the high-temperature separator bottom line 86, recycled and returned to mix with the input olefin streams in lines 13a and 13b in the oligomerization section 10 in Figure 1, preferably the first input olefin stream in line 13a, to provide the first diluted input olefin stream in line 16a, thereby absorbing the heat generated in the 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 supplied. The diesel product stream in diesel product line 118 is cooled and can be supplied to the diesel pool.

[0093] The disclosed process can efficiently produce sustainable aviation fuel that meets applicable fuel requirements, starting from ethylene and / or propylene, while controlling heat generation due to exothermic reactions. The process can also produce green gasoline and diesel by-products, which can be used as-is or as blending components in fuel that meet applicable fuel requirements depending on the specific application. The carbon recovery rate in the process can exceed 95%. Both the jet fuel flow in tower-side line 104 and the diesel product flow in line 118 can be cooled and supplied to their respective fuel pools. [Examples]

[0094] A first-stage oligomerization catalyst containing zeolite and a second-stage oligomerization catalyst containing metal were loaded into a multi-bed pilot plant reactor. The reactor was supplied primarily with ethylene, propylene, and paraffin diluents, along with small amounts of C4+ olefin. The test was conducted at a pressure of 6.2 MPa (gauge) (900 psig), varying the inlet temperature to the first oligomerization catalyst from 160 to 250°C, and varying the inlet temperature to the second oligomerization catalyst from 130 to 250°C, while controlling the WHSV of fresh olefin input for 0.25 to 2.0 hours. -1 The experiment was conducted by varying the conditions. Depending on the test conditions, the results showed ethylene and propylene conversion rates exceeding 97% by weight, as shown in Figure 3, and high jet and recyclable light olefin selectivity, as shown in Figure 4. Figure 5 shows the light paraffin selectivity, which is approximately 3-4% by weight in a steady state under these conditions. The mixture consists mostly of ethane and propane.

[0095] Next, the oligomerized product generated in this pilot plant experiment was hydrogenated and rectified. Table 1 shows the physical properties of the jet fuel produced from the hydrogenated oligomerized product.

[0096] [Table 1]

[0097] Specific Embodiments The following will be explained in conjunction with specific embodiments, but it should be understood that this explanation is intended to illustrate the scope of the preceding explanation and the attached claims, and is not intended to limit them.

[0098] A first embodiment of the present disclosure is a process for oligomerizing an olefin flow, comprising: oligomerizing a gas-phase and liquid-phase input olefin flow on a first-stage oligomerization catalyst to produce a first-stage oligomerized flow; oligomerizing the first-stage oligomerized flow on a second-stage oligomerization catalyst to provide a second-stage oligomerized flow; hydrogenating a heavy olefin flow extracted from the second-stage oligomerized flow to provide a hydrogenated flow; and stripping a light gas from the hydrogenated flow to provide a stripped olefin flow. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraphs up to the first embodiment of this paragraph, further comprising: stripping a light gas from the hydrogenated flow; separating the hydrogenated flow to provide a hydrogenated liquid flow; and stripping a light gas from the hydrogenated liquid flow. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraphs to the first embodiment of this paragraph, further comprising stripping a light gas from a hydrogenated liquid flow, separating the hydrogenated flow into a high-temperature vapor flow and a high-temperature hydrogenated liquid flow, separating the high-temperature vapor flow into a low-temperature vapor flow and a low-temperature hydrogenated liquid flow, and stripping a light gas from the low-temperature hydrogenated liquid flow and the high-temperature hydrogenated liquid flow. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraphs to the first embodiment of this paragraph, further comprising rectifying the stripped hydrogenated flow to provide a naphtha flow, a jet fuel flow, and a diesel flow. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraphs to the first embodiment of this paragraph, further comprising recirculating a portion of the diesel flow to an oligomerization process. The process according to claim 1, further comprising: dividing the oligomerization flow of the second stage into a light olefin flow and a heavy olefin flow; and re-boiling the heavy olefin flow by heat exchange with a portion of the diesel flow.Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraph to the first embodiment of this paragraph, further comprising splitting the second-stage oligomerized flow into a light olefin flow and a heavy olefin flow, depropanating the second-stage oligomerized flow to provide a light alkane flow and a depropanated flow, and splitting the depropanated flow. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraph to the first embodiment of this paragraph, further comprising re-boiling the depropanated flow by heat exchange with a portion of the diesel flow. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraph to the first embodiment of this paragraph, further comprising oligomerizing the input olefin flow with a recycled olefin flow on a first-stage oligomerization catalyst, the recycled olefin flow being removed from the light olefin flow. Embodiments of the present disclosure further include oligomerizing an input olefin stream with a recycled olefin stream on a first-stage oligomerization catalyst, wherein the recycled olefin stream is removed from the light olefin stream, and are any or all of the embodiments described above to the first embodiment described above.

[0099] A second embodiment of the present disclosure is a process for oligomerizing an olefin flow, comprising: oligomerizing a gas-phase and liquid-phase input olefin flow on a first-stage oligomerization catalyst to produce a first-stage oligomerized flow; oligomerizing the first-stage oligomerized flow on a second-stage oligomerization catalyst to provide a second-stage oligomerized flow; hydrogenating a heavy olefin flow extracted from the second-stage oligomerized flow to provide a hydrogenated flow; stripping a light gas from the hydrogenated flow to provide a stripped flow; and rectifying the stripped flow to provide a naphtha flow, a jet fuel flow, and a diesel flow. Embodiments of the present disclosure are any or all of the embodiments described above to the second embodiment described above, further comprising: stripping a light gas from the hydrogenated flow; separating the hydrogenated flow to provide a hydrogenated liquid flow; and stripping a light gas from the hydrogenated liquid flow. Embodiments of the present disclosure are any or all of the embodiments from the earlier embodiments of this paragraph to the second embodiments of this paragraph, further comprising stripping light gas from a hydrogenated liquid flow, separating the hydrogenated flow into a high-temperature vapor flow and a high-temperature hydrogenated liquid flow, separating the high-temperature vapor flow into a low-temperature vapor flow and a low-temperature hydrogenated liquid flow, and stripping light gas from the low-temperature hydrogenated liquid flow and the high-temperature hydrogenated liquid flow. Embodiments of the present disclosure are any or all of the embodiments from the earlier embodiments of this paragraph to the second embodiments of this paragraph, further comprising recirculating a portion of the diesel flow to the oligomerization process. Embodiments of the present disclosure are any or all of the embodiments from the earlier embodiments of this paragraph to the second embodiments of this paragraph, further comprising splitting the second-stage oligomerization flow into a light olefin flow and a heavy olefin flow, and re-boiling the heavy olefin flow by heat exchange with a portion of the diesel flow.

[0100] A third embodiment of the present disclosure is a process for oligomerizing an olefin flow, comprising: oligomerizing a gas-phase and liquid-phase input olefin flow on a first-stage oligomerization catalyst to generate a first-stage oligomerized flow; oligomerizing the first-stage oligomerized flow on a second-stage oligomerization catalyst to provide a second-stage oligomerized flow; hydrogenating a heavy olefin flow extracted from the second-stage oligomerized flow to provide a hydrogenated flow; separating the hydrogenated flow to provide a hydrogenated liquid flow; and stripping a light gas from the hydrogenated liquid flow to provide a stripped flow. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraphs to the third embodiment of this paragraph, further comprising stripping light gases from a hydrogenated liquid flow, separating the hydrogenated flow into a high-temperature vapor flow and a high-temperature hydrogenated liquid flow, separating the high-temperature vapor flow into a low-temperature vapor flow and a low-temperature hydrogenated liquid flow, and stripping light gases from the low-temperature hydrogenated liquid flow and the high-temperature hydrogenated liquid flow. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraphs to the third embodiment of this paragraph, further comprising rectifying the stripped hydrogenated flow to provide a naphtha flow, a jet fuel flow, and a diesel flow. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraphs to the third embodiment of this paragraph, further comprising recirculating a portion of the diesel flow to an oligomerization process. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraph to the third embodiment described in this paragraph, further comprising: dividing the second-stage oligomerization flow into a light olefin flow and a heavy olefin flow; and re-boiling the heavy olefin flow by heat exchange with a portion of the diesel flow.

[0101] Without further detail, it is expected that those skilled in the art will be able to utilize the Disclosure to the fullest extent without departing from the spirit and scope of the Disclosure, readily identify its essential characteristics, and make various changes and modifications to adapt it to various uses and conditions. Accordingly, the prior preferred specific embodiments should be interpreted as merely illustrative examples and not in any way limiting the remainder of the Disclosure, but are intended to cover various modifications and equivalent configurations that fall within the scope of the appended claims.

[0102] In the above, all temperatures are given in degrees Celsius, and all parts and percentages are based on weight unless otherwise indicated.

Claims

1. This is a process for oligomerizing olefin flows. The gas phase and liquid phase olefin streams are oligomerized on the first-stage oligomerization catalyst to generate the first-stage oligomerized stream, The first-stage oligomerized flow is oligomerized on the second-stage oligomerization catalyst to provide the second-stage oligomerized flow. The heavy olefin stream extracted from the oligomerized stream in the second stage is hydrogenated to provide a hydrogenated stream. This includes stripping light gas from the hydrogenation flow to provide a stripped hydrogenation flow, The first-stage oligomerization catalyst comprises a zeolite catalyst, and the second-stage oligomerization catalyst comprises an amorphous silica-alumina base and at least one metal selected from group VIII and group VIB metals. process.

2. This is a process for oligomerizing olefin flows. The gas phase and liquid phase olefin streams are oligomerized on the first-stage oligomerization catalyst to generate the first-stage oligomerized stream, The first-stage oligomerized flow is oligomerized on the second-stage oligomerization catalyst to provide the second-stage oligomerized flow. The heavy olefin stream extracted from the oligomerized stream in the second stage is hydrogenated to provide a hydrogenated stream. To strip the light gas from the aforementioned hydrogenation flow and provide a stripped hydrogenation flow, This includes rectifying the stripping flow to provide a naphtha flow, a jet fuel flow, and a diesel flow. The first-stage oligomerization catalyst comprises a zeolite catalyst, and the second-stage oligomerization catalyst comprises an amorphous silica-alumina base and at least one metal selected from group VIII and group VIB metals. process.

3. This is a process for oligomerizing olefin flows. The gas phase and liquid phase olefin streams are oligomerized on the first-stage oligomerization catalyst to generate the first-stage oligomerized stream, The first-stage oligomerized flow is oligomerized on the second-stage oligomerization catalyst to provide the second-stage oligomerized flow. The heavy olefin stream extracted from the oligomerized stream in the second stage is hydrogenated to provide a hydrogenated stream. To separate the hydrogenation flow and provide a hydrogenated liquid flow, This includes stripping a light gas from the aforementioned hydrogenated liquid flow to provide a stripped flow, The first-stage oligomerization catalyst comprises a zeolite catalyst, and the second-stage oligomerization catalyst comprises an amorphous silica-alumina base and at least one metal selected from group VIII and group VIB metals. process.

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