Process for converting olefins to distillate fuels

Abstract

A process for dimerizing and oligomerizing an olefin into a distillate fuel manages the dimerization exotherm by diluting the olefin with a paraffin that is inert in the dimerization. The olefin stream can also be separated and fed to multiple dimerization reactors to further reduce the heat generated. The ethylene feed can also be cooled before entering the dimerization reactor. The paraffin is obtainable from a saturated oligomeric effluent.

Description

[0001] Priority declaration

[0002] This application claims priority to U.S. patent application serial number 17 / 872,714 filed on July 25, 2022, the entire contents of which are incorporated herein by reference. Technical Field

[0003] The field is the conversion of olefins to distillates. The field may particularly relate to dimerizing olefins and oligomerizing dimerized olefins to distillate fuels. Background Art

[0004] Ethylene can be dimerized into olefins, such as C4, C6 and C8 olefins. Olefin oligomerization is the process by which smaller olefins can be oligomerized into larger olefins. More specifically, it can convert olefins (including dimerized olefins) into distillates (including jet fuel and diesel range products). The oligomeric distillates can be saturated to be used as transportation fuels.

[0005] The dimerization reaction of ethylene is highly exothermic. The exotherm generated by the dimerization of ethylene can be difficult to manage.

[0006] Jet fuel is one of the few petroleum fuels that cannot be easily replaced by electric motor systems because of the high energy output required to fuel an aircraft, which electric motors cannot provide. In some regions, there are currently large incentives for green jet fuel.

[0007] An efficient process for converting ethylene to distillate fuels is desired. Summary of the invention

[0008] We have developed a method for dimerizing and oligomerizing olefins into distillate fuels that manages the ethylene dimerization exotherm by diluting the olefins with a paraffin stream that is inert in the dimerization. The olefin stream can also be separated and fed to multiple dimerization reactors to further reduce the heat generated. The ethylene feed can also be cooled before entering the dimerization reactor. The paraffin stream can be obtained from a saturated oligomerization effluent. BRIEF DESCRIPTION OF THE DRAWINGS

[0009] Figure 1 is a schematic diagram of the oligomerization section of the method and apparatus of the present disclosure.

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

[0011] definition

[0012] The term "communication" means that fluid flow is operatively allowed between the enumerated components, which can be characterized as "fluid communication".

[0013] The term "downstream communication" means that at least a portion of the fluid flowing to the body in the downstream communication can operatively flow from the object in fluid communication therewith.

[0014] The term "upstream communication" means that at least a portion of the fluid flowing out of the body in the upstream communication can be operatively flowed to the object in fluid communication therewith.

[0015] The term "direct communication" means that the fluid flows from the upstream component into the downstream component without passing through any other intervening container.

[0016] The term "indirect communication" means that the fluid flow from the upstream component enters the downstream component after passing through an intervening container.

[0017] The term "bypass" means that the object loses downstream communication with the bypassed body at least within the scope of the bypass.

[0018] As used herein, the term "predominantly" or "majority" means greater than 50%, suitably greater than 75%, and preferably greater than 90%.

[0019] The term "tower" means one or more distillation towers for separating one or more components with different volatilities. Unless otherwise specified, each tower includes a condenser on the top of the tower for condensing a portion of the overhead stream and refluxing it back to the top of the tower, and a reboiler at the bottom of the tower for vaporizing a portion of the bottom stream and sending it back to the bottom of the tower. The feed to the tower can be preheated. The top pressure is the pressure of the overhead vapor at the vapor outlet of the tower. The bottom temperature is the liquid bottom outlet temperature. The top line and the bottom line refer to the net line from any reflux or reboil to the tower downstream. The stripping tower can omit the reboiler at the bottom of the tower, but provide heating requirements and separation power for the liquefied inert medium (such as steam). The stripping tower is usually fed from the top tray and the main product is taken out from the bottom.

[0020] As used herein, the term "separator" means a container having an inlet and at least one overhead vapor outlet and a bottom liquid outlet, and may also have an outlet for a water-containing stream from a storage tank (boot). A flash tank is a type of separator that can be communicated with a separator downstream that can operate at a higher pressure. As used herein, the term "boiling temperature" means the atmospheric pressure equivalent boiling point (AEBP) calculated from the observed boiling temperature and distillation pressure, such as using the formula provided in ASTM D1160 Appendix A7, entitled "Practice for Converting Observed Vapor Temperatures to Atmospheric Equivalent Temperatures".

[0021] As used herein, the term "true boiling point" (TBP) means a test method for determining the boiling point of a substance in accordance with ASTM D-2892 for producing standardized masses of liquefied gases, distillate fractions, and residues for which analytical data are available, and for determining the yields of the above fractions by both mass and volume from which a plot of distillation temperature versus mass % is obtained in a column having a reflux ratio of 5:1 using fifteen theoretical plates.

[0022] As used herein, the term "T5," "T90," or "T95" means the temperature at which 5 mass percent, 90 mass percent, or 95 mass percent, as the case may be, of a sample boils using ASTM D-86 or TBP, respectively.

[0023] As used herein, the term "initial boiling point" (IBP) means the temperature at which a sample begins to boil using ASTM D-7169, ASTM D-86, or TBP, as appropriate.

[0024] As used herein, the term "endpoint" (EP) means the temperature at which a sample is brought to a complete boil using ASTM D-7169, ASTM D-86, or TBP, as appropriate.

[0025] As used herein, the term "diesel" means hydrocarbons boiling in the following ranges: IBP between 125°C (257°F) and 175°C (347°F), or T5 between 150°C (302°F) and 200°C (392°F), and "diesel cut points" including T95 between 343°C (650°F) and 399°C (750°F) using the TBP distillation method, or T90 between 280°C (536°F) and 340°C (644°F) using ASTM D-86. The term "green diesel" means diesel that contains hydrocarbons not derived from fossil fuels.

[0026] As used herein, the term "jet fuel" refers to hydrocarbons that boil within the T10 range between 190°C (374°F) and 215°C (419°F) and have endpoints between 290°C (554°F) and 310°C (590°F). The term "green jet fuel" means a jet fuel that contains hydrocarbons that are not derived from fossil fuels. DETAILED DESCRIPTION

[0027] The disclosed method relates to dimerizing an olefin stream comprising ethylene, and then further oligomerizing ethylene dimers and ethylene oligomers. The resulting oligomers can be separated to provide a distillate stream, which can be saturated to provide a distillate fuel. The saturated stream can be recycled to the ethylene dimerization as a diluent to absorb the exothermic heat. In addition, the olefin stream can be separated and loaded into several dimerization catalyst beds, also in order to manage the exothermic heat. The dimerization product of the catalyst bed that may contain unconverted olefins is also transferred to the downstream bed, thereby improving the overall per-pass conversion. In addition, the dimerization product of the upstream catalyst bed is used as an additional diluent to absorb the exothermic heat in the downstream catalyst bed.

[0028] The method and apparatus may include Figure 1 The oligomeric segments 10 and Figure 2 The hydrogenation section 110 in the embodiment of the present invention.

[0029] Go to Figure 1 The oligomerization section 10 of the present invention provides the feed olefin stream in the pipeline 12 to the oligomerization section 10. The feed olefin stream may contain a large amount of ethylene. The feed olefin stream may mainly contain ethylene. On the one hand, the feed olefin stream may contain at least 95 mol% ethylene. The feed olefin stream in the pipeline 12 may be designed as an ethylene stream. The olefin stream may be provided by dehydration of ethanol or provided from an MTO unit. The temperature of the feed olefin stream may be 60°C (140°F) to 150°C (302°F), preferably 80°C (176°F) to 100°C (212°F), and the pressure may be 5.6MPag (800psig) to 8.4MPag (1200psig).

[0030] The olefin stream may be initially contacted with a dimerization catalyst to dimerize ethylene into dimers, and then contacted with an oligomerization catalyst to oligomerize the diethylene. The dimerization reaction generates a large amount of heat release. For example, the dimerization of ethylene can generate 612 kcal / kg (1100 BTU / lb) of heat. Therefore, this large amount of heat release must be managed.

[0031] Thus, the olefin stream in line 12 can be separated into a plurality of olefin streams. Figure 1 In the embodiment of the present invention, the olefin feed stream is separated into four independent streams: the first olefin feed stream in feed line 12a, the second olefin feed stream in feed line 12b, the third olefin feed stream in feed line 12c and the fourth or last olefin feed stream in feed line 12d. More or less independent multiple olefin feed streams can be used. It is easy to imagine up to six olefin feed streams. The feed olefin feed stream in pipeline 12 can be separated into multiple olefin feed streams of equal equal parts. Alternatively, the feed olefin feed stream in pipeline 12 can be separated into unequal streams. For example, the feed olefin feed stream can be separated into the stream of rising flow, wherein the subsequent olefin feed stream has a larger flow than the aforementioned stream. In one embodiment, the feed olefin feed stream is separated into four streams of equal flow, and each stream comprises 25 volume % of the feed olefin feed stream.

[0032] In order to manage the exotherm, the olefin stream can be diluted with a diluent stream to provide a diluted olefin stream to absorb the exotherm. The diluent stream can be included in the paraffin stream in the diluent line 14. Before the feed olefin stream is separated into a plurality of olefin streams, the diluent stream in the diluent line 14 can be added to the feed olefin stream in the pipeline 12. Preferably, after being separated into a plurality of olefin streams, the diluent stream is added to the first olefin stream in the pipeline 12a to provide the first diluted olefin stream in the pipeline 16a, so that the diluent stream passes through all dimerization reactions. Alternatively, the diluent stream can also be separated into a plurality of streams, wherein each diluent stream is added to the corresponding olefin stream. The diluent stream can have a volume flow of 2 to 8 times and preferably 3 to 6 times of the feed olefin stream volume flow. The first diluted olefin stream can include no more than 25 weight % of olefins, suitably no more than 10 weight % of olefins and preferably no more than 6 weight % of olefins. The first diluted olefin stream may contain no more than 25 wt% ethylene, suitably no more than 10 wt% ethylene, and preferably no more than 6 wt% ethylene. The first diluted olefin stream in line 16a may be cooled in a first feed cooler 18a to provide a first cooled diluted olefin stream in line 20a and charged to a first bed 22a of a dimerization catalyst in a dimerization reactor 22. The first cooled diluted olefin stream in line 20a may be charged at a temperature of 54°C (130°F) to 165°C (329°F) and a pressure of 5.6 MPag (800 psig) to 8.4 MPag (1200 psig).

[0033] The dimerization reactor 22 may include a series of dimerization catalyst beds 22a, 22b, 22c and 22d for loading with multiple olefin streams 12a, 12b, 12c and 12d, respectively. The dimerization reactor preferably contains four fixed dimerization catalyst beds 22a, 22b, 22c and 22d. It is also envisioned that each dimerization catalyst bed 22a, 22b, 22c and 22d can be in a dedicated dimerization reactor, or that multiple dimerization catalyst beds can be in two or more independent dimerization reactors. It is easy to envision up to six dimerization catalyst beds. When the dimerization reactor 22 has been passivated, a parallel dimerization reactor can be used, during which the dimerization reactor 22 is regenerated in situ by burning coke from the catalyst.

[0034] The first cooled diluted olefin stream can preferably be loaded into the first catalyst bed 22a in the pipeline 20a with a downward flow operation. However, an upward flow operation may be suitable. When the dimerization of ethylene occurs in the first dimerization catalyst bed 22a, an exothermic heat is generated due to the exothermic nature of the ethylene dimerization reaction. Despite cooling and dilution, the dimerization of the first olefin stream produces the first dimerized olefin stream in the first dimerization effluent pipeline 24a at an elevated outlet temperature. The elevated outlet temperature is limited to between 50°C (90°F) and 61°C (110°F), which is higher than the inlet temperature to the catalyst bed 22a.

[0035] The second olefin stream in line 12b can be diluted with the first dimerized olefin stream in the line 24a removed from the dimerization reactor 22 to provide the second diluted olefin stream in line 16b. The first dimerized olefin stream in line 24a includes a diluent stream from the diluent line 14, which is added to the first olefin stream in line 12a. The second diluted olefin stream can include no more than 25% by weight of ethylene, suitably no more than 10% by weight of ethylene, and preferably no more than 6% by weight of ethylene. The second diluted olefin stream in line 16b can be cooled in the second charging cooler 18b positioned outside the dimerization reactor 22 to provide the second cooled diluted olefin stream in line 20b, and loaded into the second bed 22b of the dimerization catalyst in the dimerization reactor 22. The second cooled diluted olefin stream in line 20b can be charged at a temperature of 54°C (130°F) to 165°C (329°F) and a pressure of 5.6MPag (800psig) to 8.4MPag (1200psig). The second diluted olefin stream will contain diluent and olefins from the first dimerized olefin stream. The olefins from the first dimerized olefin stream will dimerize in the second catalyst bed 22b. The dimerization of ethylene in the second olefin stream in the second bed 22b of the dimerization catalyst produces a second dimerized olefin stream in the second dimerization effluent line 24b at an elevated outlet temperature. The elevated outlet temperature can be limited to between 50°C (90°F) and 61°C (110°F), which is higher than the inlet temperature to the catalyst bed 22b.

[0036] The third olefin stream in line 12c can be diluted with the second dimerized olefin stream in line 24b removed from dimerization reactor 22 to provide the third diluted olefin stream in line 16c. The second dimerized olefin stream in line 24b includes a diluent stream from diluent line 14, which is added to the first olefin stream in line 12a. The third diluted olefin stream can include no more than 25% by weight of ethylene, suitably no more than 10% by weight of ethylene, and preferably no more than 6% by weight of ethylene. The third diluted olefin stream in line 16c can be cooled in the third feed cooler 18c positioned outside dimerization reactor 22 to provide the third cooled diluted olefin stream in line 20c, and loaded into the third bed 22c of the dimerization catalyst in dimerization reactor 22. The third cooled diluted olefin stream in line 20c can be charged at a temperature of 54°C (130°F) to 165°C (329°F) and a pressure of 5.6MPag (800psig) to 8.4MPag (1200psig). The third diluted olefin stream will contain diluent and olefins from the second dimerized olefin stream. The olefins from the second dimerized olefin stream will dimerize in the third catalyst bed 22c. The dimerization of ethylene in the third olefin stream in the third bed 22c of the dimerization catalyst produces a third dimerized olefin stream in the third dimerization effluent line 24c at an elevated outlet temperature. In one embodiment, the third dimerized olefin stream is the penultimate dimerized olefin stream, and the third dimerization effluent line 24c is the penultimate dimerization effluent line 24c. The elevated outlet temperature is limited to between 50°C (90°F) and 61°C (110°F), which is higher than the inlet temperature to the catalyst bed 22c.

[0037] The fourth olefin stream in line 12d can be diluted with the third or penultimate dimerized olefin stream in line 24c removed from dimerization reactor 22 to provide the fourth diluted olefin stream in line 16d. The third or penultimate dimerized olefin stream in line 24c includes a diluent stream from diluent line 14, which is added to the first olefin stream in line 12a. The fourth diluted olefin stream can include no more than 25 wt % of ethylene, suitably no more than 10 wt % of ethylene, and preferably no more than 6 wt % of ethylene. The fourth diluted olefin stream in line 16d can be cooled in a fourth feed cooler 18d positioned outside dimerization reactor 22 to provide the fourth cooled diluted olefin stream in line 20d, and loaded into the fourth bed 22d of the dimerization catalyst in dimerization reactor 22. The fourth cooled diluted olefin stream in line 20d can be charged at a temperature of 54°C (130°F) to 165°C (329°F) and a pressure of 5.6MPag (800psig) to 8.4MPag (1200psig). The fourth or last diluted olefin stream will contain diluent and olefins from the third or penultimate dimerized olefin stream. Olefins from the third or penultimate dimerized olefin stream will dimerize in the fourth catalyst bed 22b. The dimerization of ethylene in the fourth olefin stream in the fourth bed 22d of the dimerization catalyst produces a fourth dimerized olefin stream in the fourth dimerization effluent line 24d at an elevated outlet temperature. The elevated outlet temperature is limited to between 50°C (90°F) and 61°C (110°F), which is higher than the inlet temperature to the catalyst bed 22d.

[0038] In one embodiment, the fourth olefin stream is the last olefin stream, the fourth dimerized olefin stream is the last dimerized olefin stream, and the fourth dimerization effluent line 24d is the last dimerization effluent line 24d.

[0039] The dimerization reaction is carried out mainly in the liquid phase or in a mixed liquid and gas phase at a rate of 0.5 to 10 hours based on the olefin. -1 The LHSV of the catalyst occurs. We have found that the major fraction of ethylene in the olefin stream is converted to higher olefins. Typically, at least 90 to 95 mole percent of the ethylene will be dimerized by the dimerization catalyst bed. The ethylene will initially dimerize to butenes over the catalyst.

[0040] The dimerization catalyst is preferably an amorphous silica alumina base with a metal from Group VIII and / or Group VIB of the periodic table using Chemical Abstracts Service symbols. In one aspect, the catalyst has a Group VIII metal promoted with a Group VIB metal. Typically, the silica and alumina are only in the base so that the silica / alumina ratio of the catalyst is the same as the base. The metal can be impregnated onto the silica alumina base or ion exchanged with the silica alumina base. Co-grinding is also contemplated. The catalyst used in the present invention may have a low temperature acidity ratio of at least 0.15, suitably 0.2 and preferably greater than 0.25, as determined by ammonia programmed temperature desorption (ammonia TPD) as described below. In addition, a suitable catalyst will have a low temperature acidity ratio between 50 and 400 m 2 The surface area is between 1000 Å / g as determined by nitrogen BET.

[0041] Preferred dimerization catalysts are described below. Preferred dimerization catalysts include an amorphous silica alumina support. One of the components of the catalyst support for the present invention is alumina. Alumina can be any of various hydrated aluminum oxides or alumina gels such as α-alumina monohydrate of boehmite or pseudo-boehmite structure, α-alumina trihydrate of gibbsite structure, β-alumina trihydrate of bayerite structure, etc. Particularly preferred alumina can be purchased from Sasol North America Alumina Product Group under the trade name Catapal. This material is an extremely high purity α-alumina monohydrate (pseudo-boehmite) that has been shown to produce high purity γ-alumina after calcination at high temperatures. Another component of the catalyst support is amorphous silica alumina. Suitable silica alumina with a silica to alumina ratio of 2.6 is purchased from CCIC, a subsidiary of JGC, Japan.

[0042] Another component used to prepare the catalyst used in the present invention is a surfactant. The surfactant is preferably mixed with the above-mentioned aluminum oxide and silica alumina powder. The resulting mixture of surfactant, aluminum oxide and silica alumina is then formed as described below, dried and calcined. Calcination effectively removes the organic components of the surfactant by combustion, but only after the surfactant has faithfully performed its function according to the present invention. Any suitable surfactant can be used according to the present invention. A preferred surfactant is a surfactant selected from a series of commercial surfactants sold by Solvay SA under the trade name "Antarox". "Antarox" surfactants are generally characterized as modified linear aliphatic polyethers and are low-foaming biodegradable detergents and wetting agents.

[0043] Suitable silica alumina mixtures are prepared by mixing proportions of silica alumina and alumina in volume to achieve the desired silica to alumina ratio. In one embodiment, 75 to 95 weight percent amorphous silica alumina and 10 to 20 weight percent alumina powder will provide a suitable support with a silica to alumina ratio of 2.6. In one embodiment, other ratios of amorphous silica alumina to alumina may be suitable.

[0044] The surfactant may be combined with the silica alumina and alumina mixture using any convenient method. The surfactant is preferably mixed during the mixing and forming of the alumina and silica alumina. A preferred method is to mix an aqueous solution of the surfactant with the blend of alumina and silica alumina prior to final formation of the carrier. Preferably, the surfactant is present in the paste or dough in an amount of 0.01 wt % to 10 wt % based on the weight of the alumina and silica alumina.

[0045] A monoprotic acid such as nitric acid or formic acid may be added to the mixture in aqueous solution to peptize the alumina in the binder.Additional water may be added to the mixture to provide sufficient moisture to form a dough of sufficient consistency to be extruded or spray dried.

[0046] The paste or dough may be prepared in the form of shaped particles, preferably by extruding a dough mixture of alumina, silica alumina, surfactant and water through a die having openings of the desired size and shape therein, after which the extruded mass is broken into extrudates of the desired length and dried. A further calcination step may be employed to impart increased strength to the extrudates. Typically, calcination is carried out in a dry air stream at a temperature of 260°C (500°F) to 815°C (1500°F).

[0047] The extruded particles may have any suitable cross-sectional shape, i.e. symmetrical or asymmetrical, but most typically have a symmetrical cross-sectional shape, preferably spherical, cylindrical or multilobed. The cross-sectional diameter of the particles may be as small as 40 μm; however, it is typically 0.635 mm (0.25 in) to 12.7 mm (0.5 in), preferably 0.79 mm (1 / 32 in) to 6.35 mm (0.25 in), and most preferably 0.06 mm (1 / 24 in) to 4.23 mm (1 / 6 in).

[0048] The typical characteristics of the amorphous silica alumina support used herein are that the total pore volume, average pore size and surface area are large enough to provide a large amount of space and area for depositing active metal components. As measured by conventional mercury intrusion method, the total pore volume of the support is generally 0.2cc / gram to 2.0cc / gram, preferably 0.25cc / gram to 1.0cc / gram, and most preferably 0.3cc / gram to 0.9cc / gram. Typically, the pore volume of the support in pores with a diameter greater than 100 angstroms is less than 0.1cc / gram, preferably less than 0.08cc / gram, and most preferably less than 0.05cc / gram. As measured by the BET method, the surface area is generally above 50m 2 / g, for example, higher than 200m 2 / g, preferably at least 250m 2 / gram, and most preferably 300m 2 g to 400m 2 / gram.

[0049] To prepare the catalyst, the support material is compounded with one or more precursors of at least one metal component from Group VIII or VIB of the periodic table (such as by single or multiple impregnations of calcined amorphous refractory oxide support particles). The Group VIII metal (preferably nickel) should be present in a concentration of 0.5% to 15% by weight, and the Group VIB metal (preferably tungsten) should be present in a concentration of 0 to 12% by weight. Impregnation can be achieved by any method known in the art (such as, for example, by spray impregnation), wherein a solution containing a metal precursor in 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. Other methods involve immersing the support in a large volume of impregnation solution or circulating the support therein, and yet another method is a pore volume or pore saturation technique, in which the support particles are introduced into an impregnation solution of a volume just sufficient to fill the pores of the support. Sometimes, the pore saturation technique can be modified to utilize an impregnation solution having a volume that is 10% less to 10% greater than that which just fills the pores.

[0050] If the active metal precursors are combined by impregnation, a subsequent or second calcination at an elevated temperature, such as, for example, between 399°C (750°F) and 760°C (1400°F), converts the metals to their corresponding oxide forms. In some cases, calcination can be performed after each impregnation of the individual active metals. Subsequent calcinations will produce a catalyst containing the active metals in their corresponding oxide forms.

[0051] The preferred dimerization catalyst of the present invention has an amorphous silica alumina base material impregnated with 0.5 wt % to 15 wt % nickel in the form of 3.175 mm (0.125 inch) extrudates and a density of 0.45 g / ml to 0.65 g / ml. It is also contemplated that the metal may be incorporated into the support by other methods such as ion exchange and co-grinding.

[0052] The dimerization catalyst can be regenerated when passivated. Suitable regeneration conditions include subjecting the catalyst to hot air at 500° C. for 3 hours, for example, in situ. To facilitate regeneration without shutting down the plant, a swing bed arrangement can be employed with an alternative dimerization reactor. The regeneration gas can include air with an increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to those of the fresh catalyst.

[0053] The final dimer olefin stream in the final dimer effluent line 24d having an increased concentration of ethylene dimers and oligomers compared to the feed olefin stream in line 12 is mixed with the oligomer recycle stream in line 26 to provide a feed oligomer stream in line 28. The feed oligomer stream is heated by heat exchange with the net olefin splitter bottoms stream in line 30 and is charged to the oligomerization reactor 32 at a temperature of 204° C. (400° F.) to 265° C. (509° F.) at a pressure of 5.6 MPa (800 psig) to 8.4 MPa (1200 psig).

[0054] Oligomerization reactor 32 can be communicated with dimerization reactor 22 downstream.Oligomerization reactor 32 is preferably operated in a downward flow operation.However, upward flow operation may be suitable.The feed oligomerization stream contacts with an oligomerization catalyst, causing C2-C8 olefin dimerization and trimerization to provide olefins in the distillate range.Most of the butenes in the feed oligomerization stream are oligomerized.In one embodiment, at least 99 mol % of the butenes in the feed oligomerization stream are oligomerized.The average carbon number of the increase is higher than the oligomerization olefin stream of the feed oligomerization olefin stream in line 28 and leaves oligomerization reactor 32 in line 34.

[0055] The oligomerization catalyst may include a zeolite catalyst. Zeolite may account for between 5% and 95% by weight of the catalyst, for example, between 5% and 85% by weight. Suitable zeolites include zeolites having a structure of one of the following categories: MFI, MEL, ITH, IMF, TUN, FER, BEA, FAU, BPH, MEI, MSE, MWW, UZM-8, MOR, OFF, MTW, TON, MTT, AFO, ATO and AEL. The 3-letter code indicating a zeolite is defined by the Structure Commission of the International Zeolite Association and is maintained at http: / / www.iza-structure.org / databases. UZM-8 is as described in U.S. Patent No. 6,756,030. In preferred aspects, the oligomerization catalyst may include a zeolite having a skeleton with a ten-ring pore structure. Examples of suitable zeolites with a ten-ring pore structure include TON, MTT, MFI, MEL, AFO, AEL, EUO and FER. In further preferred aspects, the oligomerization catalyst comprising the zeolite with ten-ring pore structure can comprise a one-dimensional pore structure. The one-dimensional pore structure indicates a zeolite containing non-intersecting pores substantially parallel to one of the crystal axes. The hole preferably extends through the zeolite crystal. Suitable examples of zeolites with ten-ring one-dimensional pore structure can include MTT. In other aspects, the oligomerization catalyst comprises MTT zeolite.

[0056] The oligomerization catalyst can be formed by combining the zeolite with a binder and then forming the catalyst into pellets. The pellets can be optionally treated with a phosphorus reagent to produce a zeolite having a phosphorus component of between 0.5% and 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, zirconium oxide, titanium dioxide 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 clay.

[0057] 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 of boehmite or pseudo-boehmite structure, alpha-alumina trihydrate of gibbsite structure, beta-alumina trihydrate of bayerite structure, and the like. Suitable alumina can be purchased from UOP LLC under the trade name VERSAL. Preferred alumina can be purchased from Sasol North American Alumina Products Group under the trade name Catapal. This material is an extremely high purity alpha-alumina monohydrate (pseudo-boehmite) that has been shown to produce high purity gamma-alumina after calcination at high temperatures.

[0058] Suitable oligomerization catalysts are prepared by mixing proportional volumes of zeolite and alumina to achieve the desired zeolite to alumina ratio. In one embodiment, the MTT content may be from 5 wt % to 85 wt %, such as 20 wt % to 82 wt % MTT zeolite, and the balance of alumina powder will provide a suitably supported catalyst. Silica supports are also contemplated.

[0059] A monoacid such as nitric acid or formic acid may be added to the mixture in aqueous solution to gelatinize the alumina in the binder. Additional water may be added to the mixture to provide sufficient moisture to form a dough of sufficient consistency to be extruded or spray dried. An extrusion aid such as a cellulose ether powder may also be added. A preferred extrusion aid is available from Dow Chemical Company under the trade name Methocel.

[0060] The paste or dough may be prepared in the form of shaped particles, preferably by extruding the dough through a die having an opening of the desired size and shape therein, after which the extruded material is broken into extrudates of the desired length and dried. A further calcination step may be employed to impart increased strength to the extrudates. Typically, calcination is carried out in an air stream at a temperature of 260°C (500°F) to 815°C (1500°F). The MTT catalyst is not selectively neutralized at the acid sites, such as with an amine.

[0061] The extruded particles may have any suitable cross-sectional shape, i.e. symmetrical or asymmetrical, but most typically have a symmetrical cross-sectional shape, preferably spherical, cylindrical or multilobed. The cross-sectional diameter of the particles may be as small as 40 μm; however, it is typically 0.635 mm (0.25 in) to 12.7 mm (0.5 in), preferably 0.79 mm (1 / 32 in) to 6.35 mm (0.25 in), and most preferably 0.06 mm (1 / 24 in) to 4.23 mm (1 / 6 in).

[0062] With respect to the oligomerization reactor 32, process conditions are selected to produce a higher percentage of injection range olefins, which, when hydrogenated in a subsequent step to be described below, produce the desired injection range hydrocarbon product. In an exemplary embodiment, an MTT-type zeolite catalyst disposed on a high purity pseudo-boehmite alumina substrate in a ratio of 90 / 10 to 20 / 80 and preferably between 20 / 80 and 50 / 50 is provided in a catalyst bed or more catalyst beds in the oligomerization reactor 32. The feed oligomerization stream in line 28 is heated and charged to the oligomerization reactor 32. In order to achieve the most desired olefin products, the oligomerization reactor 32 is operated at a temperature of 204°C (400°F) to 260°C (500°F). The oligomerization reactor 70 operates at a pressure of 2.1 MPa (300 psig) to 7.6 MPa (1100 psig) and more preferably 4.9 MPa (710 psig) to 6.9 MPa (1000 psig).

[0063] The oligomerization reaction is also exothermic in nature. The final dimerized olefin stream in line 24d comprises a diluent stream from diluent line 14 which is added to the first olefin stream in line 12a and carried through the dimerization catalyst beds 22a to 22d. The diluent stream is then transported in line 28 to the oligomerization reactor 32 to absorb the exotherm in the oligomerization reactor.

[0064] When the oligomerization reaction is performed according to the above process conditions, a C4 olefin conversion of greater than or equal to 95%, or greater than or equal to 97% is achieved. The resulting oligomerized olefin stream in line 34 includes a variety of olefin products that are distillate range hydrocarbons.

[0065] The oligomerization catalyst can be regenerated when passivated. Suitable regeneration conditions include subjecting the oligomerization catalyst, for example, to hot air at 500° C. for 3 hours in situ. To promote regeneration without shutting down, a swing bed arrangement can be used with an alternative oligomerization reactor. A regeneration gas stream can be passed to the oligomerization reactor 32 that needs to be regenerated. The regeneration gas can include air with an increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to those of the fresh catalyst.

[0066] The oligomeric olefin stream in line 34 having an increased concentration of C9+ olefins compared to the feed oligomeric olefin stream in line 28 is heat exchanged with the olefin splitter bottoms stream in line 30 , reduced in pressure and fed to olefin splitter column 36 .

[0067] In the olefin splitter column 36, oligomers having boiling points below the jet range hydrocarbons (typically C8 hydrocarbons having atmospheric boiling points below 150°C) are separated in an olefin splitter overhead stream in overhead line 38 from a bottoms stream in bottoms line 40 containing distillate range C9+ hydrocarbons (typically C9-C22 olefins). The olefin splitter column 36 can be operated at a bottoms temperature of 400°C (750°F) to 427°C (800°F) and an overhead pressure of 172 kPa (25 psig) to 517 kPa (75 psig). It is contemplated that the olefin splitter column 36 can be two columns. The olefin splitter overhead stream can be cooled and the resulting condensate portion can be refluxed from the olefin splitter receiver 42 back to the olefin splitter column 36. The net vapor stream in the receiver overhead line 44 from the olefin splitter receiver 42 can be compressed up to the oligomerization pressure in the offgas compressor 46 to provide a light oligomer stream in line 48 in a vapor or liquid phase after cooling. Alternatively, the olefin splitter overhead stream in the overhead line 38 can be completely condensed by cooling (perhaps in an external refrigeration loop) to provide a liquid light oligomer stream in line 48. The light oligomer stream in line 48 can be separated between a light olefin drag stream in line 50 and an oligomer recycle stream in line 26 (which is recycled to the oligomerization reactor 32). The light olefin drag stream in line 50 can account for 3 wt% to 15 wt% of the light oligomer stream in line 48. The light oligomer stream in line 48 can contain 50 wt% to 80 wt% light olefins. The oligomer recycle stream in line 26 is mixed with the last dimerized olefin stream in the last dimerization effluent line 24d to provide a feed oligomer stream in line 28 for charging to the oligomerization reactor 32. It is also contemplated that the oligomer recycle stream in line 26 is mixed with the first diluted olefin stream in line 16a or divided between the first through fourth diluted olefin streams in lines 16a through 16d to dimerize unreacted ethylene.

[0068] The heavy olefin stream in splitter bottoms line 40 can be separated between a reboiled stream that is reboiled and fed back to olefin splitter column 36 and a heavy olefin stream in net splitter bottoms line 30. Figure 2 Prior to the hydrogenation section 110 in , the heavy olefin stream is cooled by heat exchange with the oligomerized olefin stream in line 34 and then with the feed oligomer stream in line 28 .

[0069] Go to Figure 2 The hydrogenation section 110 in Figure 1The heavy olefin stream in the net olefin splitter bottoms line 30 containing distillate range C9+ oligomerized olefins can be hydrogenated in the hydrogenation reactor 52 to saturate the olefinic bonds to provide fuel. This step is performed to ensure that the motor fuel product meets or exceeds the thermal oxidation requirements specified in ASTM D7566-10a for hydrotreated synthetic paraffinic kerosene (SPK). In addition, saturating the oligomerized heavy olefins will provide 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 can be mixed with steam also from Figure 1 The light olefin drag stream in line 50 comprising C2 to C8 olefins is combined with a hydrogen stream in line 56 to produce a combined olefin stream in line 54. The combined olefin stream in line 54 may also be combined with a hydrogen stream in line 56 to provide a combined hydrogenation feed stream in line 58, which is cooled and charged to hydrogenation reactor 52 at 125° C. (257° F.) to 204° C. (400° F.) and 3.5 MPa (500 psig) to 6.9 MPa (1000 psig). Excess hydrogen may be used to ensure full saturation, such as 1.5 to 2.5 stoichiometric amounts of hydrogen.

[0070] Hydrogenation is usually carried out using conventional hydrogenation or hydroprocessing catalysts and may include metal catalysts containing, for example, palladium, rhodium, nickel, ruthenium, platinum, rhenium, cobalt, molybdenum or combinations thereof and supported forms thereof. The catalyst support may be any solid, inert material, including but not limited to oxides such as silicon dioxide, aluminum oxide, titanium dioxide, calcium carbonate, barium sulfate and carbon. The catalyst support may be in the form of powders, particles, pellets, etc.

[0071] In an exemplary embodiment, hydrogenation is performed in a hydrogenation reactor 52 that includes a platinum catalyst on alumina, such as 0.5 wt % to 0.9 wt % platinum catalyst on alumina. The hydrogenation reactor 52 converts the olefins to paraffin products having the same carbon number distribution as the olefins, thereby forming distillate range paraffins suitable for use as jet fuel and diesel fuel.

[0072] The saturated heavy stream in line 60 discharged from the hydrogenation reactor 52 can be cooled by heat exchange with the saturated heavy liquid stream in the separator bottoms line 66, and fed to the hydrogenation separator 62. In the hydrogenation separator 62, the saturated heavy stream is separated into a hydrogenation separator vapor stream in the overhead line 64 and a saturated heavy liquid stream in the hydrogenation separator bottoms line 66. A purge in line 65 can be taken from the hydrogenation separator vapor stream in line 64, and the remainder can be compressed and combined with make-up hydrogen in line 68 to provide a hydrogen stream in line 56. The saturated heavy liquid stream in the bottoms line 66 can be heated by heat exchange with the saturated heavy stream in line 60 and the diluent stream in line 14, and fed to the spray fractionation column 70.

[0073] The saturated heavy liquid stream in the bottoms line 66 can be fed to the jet fractionator 70 without prior stripping in a stripper. Alternatively, a stripper can be utilized upstream of the jet fractionator 70. In the jet fractionator 70, the saturated heavy liquid stream can be separated into an off-gas stream in an overhead line 72, a green jet stream in a side line 74 from one side of the jet fractionator 70, and a green diesel stream in a bottoms line 76. The jet fractionator 70 can be operated at a bottoms temperature of 427° C. (800° F.) to 482° C. (900° F.) and an overhead pressure of 35 kPa (5 psig) to 350 kPa (50 psig).

[0074] The jet fractionation overhead stream in overhead line 72 may be cooled and the resulting condensate portion may be refluxed from a jet fractionation receiver 78 back to the jet fractionation column 70 in line 79 while a net off-gas stream comprising C8 hydrocarbons is withdrawn from the jet fractionation receiver 78 in receiver overhead line 80. Most of the hydrocarbons in the net off-gas stream in receiver overhead line 80 are lighter hydrocarbons and may be used to fuel the reboiler of the jet fractionation column 70 and / or the olefin splitter column 36.

[0075] The green jet stream withdrawn in side line 74 contains kerosene range C9-C17 hydrocarbons and can be cooled and withdrawn as a product meeting applicable SPK standards. In an alternative embodiment, a green jet stream can be withdrawn from the condensate stream in line 79 from jet fractionation receiver 78, rather than refluxing all of the condensate to the column. This green jet stream withdrawn from line 79 will have to be stripped to remove light ends. In such embodiments, side line 74 will not be withdrawn to recover a green jet fuel stream.

[0076] The green diesel bottoms stream in bottoms line 76 may be separated between a reboiled stream that is reboiled and fed back to the jet fractionator 70, a green diesel product stream in line 82, and a diluent stream in line 14. The diluent stream in line 14 may be cooled by heat exchange with separator bottoms line 66 and by steam generation and recycled back to the jet fractionator 70. Figure 1 The olefin stream in line 12 in the oligomerization section 10 in the side line 74 (preferably the first olefin stream in line 12a) is mixed to provide a first diluted olefin stream in line 16a to absorb the exotherm in the dimerization reactor 22. The green diesel in diluent line 14 is paraffinic, so it is inert to the dimerization, oligomerization and hydrogenation reactions it may undergo. Both the jet fuel stream in side line 74 and the diesel stream in line 82 can be cooled and fed into their respective fuel pools. The diesel stream will meet the ASTM D975 standard for diesel.

[0077] Starting from ethylene, the disclosed method can efficiently produce green jet fuel and green diesel fuel that meet applicable fuel requirements while managing exothermic generation. The carbon recovery rate in the method can exceed 95%.

[0078] Specific implementation plan

[0079] While the following is described in conjunction with specific embodiments, it should be understood that this description is intended to illustrate and not to limit the scope of the foregoing description and the appended claims.

[0080] The first embodiment of the present invention is a method for oligomerizing an olefin stream, the method comprising: diluting an olefin stream with a paraffin stream to provide a diluted olefin stream; dimerizing the diluted olefin stream with a dimerization catalyst to produce a dimerized olefin stream; oligomerizing the dimerized olefin stream with an oligomerization catalyst to provide an oligomerized olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, which also includes: separating the olefin stream into multiple olefin streams; and diluting the first olefin stream in the multiple olefin streams with a paraffin stream to provide a first diluted olefin stream; and dimerizing the first diluted olefin stream to produce a first dimerized olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, which also includes cooling the first diluted olefin stream before dimerizing the olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, further comprising: diluting the second olefin stream in the plurality of olefin streams with the first dimerized olefin stream to provide a second diluted olefin stream; and dimerizing the second diluted olefin stream to produce a second dimerized olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, further comprising cooling the second diluted olefin stream before dimerizing the second diluted olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, further comprising: diluting the last olefin stream in the plurality of olefin streams with the penultimate dimerized olefin stream to provide the last diluted olefin stream; and dimerizing the last diluted olefin stream to produce a dimerized olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, further comprising: saturating the oligomerized olefin stream to provide a paraffin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, which also includes: taking out a heavy olefin stream from the oligoolefin stream and saturating the heavy olefin stream to produce a saturated heavy stream; and taking out a paraffin stream from the saturated heavy stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, which also includes: taking out a diesel stream from the saturated heavy stream; and taking out a paraffin stream from the diesel stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the olefin stream is mainly ethylene.One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the diluted olefin stream has no more than 6 weight % ethylene. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, and it also includes recycling the oligomeric olefins taken out from the oligomeric olefin stream to the oligomerization step.

[0081] A second embodiment of the present invention is a method for oligomerizing an olefin stream, the method comprising: separating an olefin stream into multiple olefin streams; dimerizing a first olefin stream in the multiple olefin streams to produce a first dimerized olefin stream; dimerizing the last olefin stream in the multiple olefin streams to produce a dimerized olefin stream; and oligomerizing the dimerized olefin stream with an oligomerization catalyst to provide an oligomerized olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, which further comprises: diluting the first olefin stream with a paraffin stream to provide a first diluted olefin stream; and dimerizing the first diluted olefin stream to produce a first dimerized olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, which further comprises cooling the first diluted olefin stream before dimerizing the olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, which also includes: diluting the second olefin stream in the plurality of olefin streams with the first dimerized olefin stream to provide a second diluted olefin stream; and dimerizing the second diluted olefin stream to produce a second dimerized olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, which also includes cooling the second diluted olefin stream before dimerizing the second diluted olefin stream. One embodiment of the present invention is one, any or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, which also includes: taking out a heavy olefin stream from the oligomerized olefin stream and saturating the heavy olefin stream to produce a saturated heavy stream; and taking out a paraffin stream from the saturated heavy stream.

[0082] The third embodiment of the present invention is a method for oligomerizing an olefin stream, the method comprising: separating an olefin stream into a plurality of olefin streams; dimerizing a first olefin stream in the plurality of olefin streams to produce a first dimerized olefin stream; and dimerizing a last olefin stream in the plurality of olefin streams to produce a dimerized olefin stream. An embodiment of the present invention is one, any one, or all of the previous embodiments in this paragraph to the third embodiment in this paragraph, further comprising oligomerizing the dimerized olefin stream with an oligomerization catalyst to provide an oligomerized olefin stream.

[0083] Although there is no further detailed description, it is believed that those skilled in the art can make the most of the present disclosure by using the foregoing description and can easily determine the essential characteristics of the present disclosure without departing from the spirit and scope of the present invention, and can make various changes and modifications of the present disclosure and adapt it to various usages and conditions. Therefore, the aforementioned preferred specific embodiments should be understood as merely illustrative and not to limit the rest of the present disclosure in any way, and are intended to cover various modifications and equivalent arrangements included within the scope of the appended claims.

[0084] In the foregoing, all temperatures are set forth in degrees Celsius and all parts and percentages are by weight unless otherwise indicated.

Claims

1. A process for oligomerizing an olefin stream, the process comprising: diluting the olefin stream with a paraffin stream to provide a diluted olefin stream; dimerizing the diluted olefin stream with a dimerization catalyst to produce a dimerized olefin stream; The dimerized olefin stream is oligomerized with an oligomerization catalyst to provide an oligomerized olefin stream.

2. The method according to claim 1, further comprising: separating the olefin stream into a plurality of olefin streams; and diluting a first olefin stream of the plurality of olefin streams with the paraffin stream to provide a first diluted olefin stream; and dimerizing the first diluted olefin stream to produce a first dimerized olefin stream.

3. The process of claim 2, further comprising cooling the first diluted olefin stream prior to dimerizing the olefin stream.

4. The method according to claim 2, further comprising: diluting a second olefin stream of the plurality of olefin streams with the first dimerized olefin stream to provide a second diluted olefin stream; and dimerizing the second diluted olefin stream to produce a second dimerized olefin stream.

5. The process of claim 4, further comprising cooling the second diluted olefin stream prior to dimerizing the second diluted olefin stream.

6. The method according to claim 4, further comprising: diluting a last olefin stream of the plurality of olefin streams with a penultimate dimerized olefin stream to provide a last diluted olefin stream; and dimerizing the final diluted olefin stream to produce the dimerized olefin stream.

7. The process of claim 1 further comprising saturating the oligoolefin stream to provide the paraffin stream.

8. The method according to claim 7, further comprising: removing a heavy olefin stream from the oligomerized olefin stream and saturating the heavy olefin stream to produce a saturated heavy stream; and removing the paraffin stream from the saturated heavy stream.

9. The method according to claim 8, further comprising: removing a diesel stream from the saturated heavy stream; and removing the paraffin stream from the diesel stream.

10. The process of claim 1 wherein the olefin stream is primarily ethylene.

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