Process for converting olefins to jet fuels using parallel reactors

Abstract

A process for oligomerizing an olefin stream, the process comprising: feeding a light olefin vapor stream comprising ethylene to a first oligomerization reactor loaded with a first oligomerization catalyst to produce a first oligomerized stream. A heavy olefin liquid stream comprising the C3-C8 olefin liquid stream is passed to a second oligomerization reactor loaded with a second oligomerization catalyst to produce a second oligomerization stream. The two oligomeric streams may be processed to recover the fuel stream together.

Description

[0001] This field relates to the conversion of olefins into distillates. In particular, this field can involve the oligomerization of olefins into distillate fuels. Background Technology

[0002] Molecular sieves, such as microporous crystalline zeolites and non-zeolite catalysts, particularly aluminosilicate silica (SAPO), are known to promote the conversion of oxygen-containing compounds such as methanol into light olefins. The efficient methanol-to-olefins (MTO) process can convert oxygen-containing compounds into light olefins, which are typically considered for use in plastics production. The light olefins produced by the MTO process are highly concentrated with ethylene and propylene.

[0003] Ethanol dehydration processes involve the dehydration of ethanol molecules to produce ethylene and water. The process of converting ethanol to ethylene is inherently endothermic, with the heat typically provided by a flame heater that is inherently adiabatic. The dehydration reactor is adiabatic. Adiabatic reactor systems may have disadvantages such as selectivity for undesirable products, potential underutilization of the catalyst, higher utility consumption, and a larger footprint.

[0004] 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. Olefin oligomerization is the process of oligomerizing smaller olefins into larger olefins. More specifically, it can convert olefins (including dimer olefins) into distillates (including products from the jet fuel and diesel range). Oligomeric distillates can be saturated for use as transportation fuels.

[0005] Jet fuel is one of the few petroleum fuels that cannot be easily replaced by electric motor systems because it requires high energy output to fuel aircraft, which electric motors cannot provide. Jet fuel is required to have a final boiling point below 300°C, as determined by ASTM D86. In some regions, there are currently significant incentives for green jet fuel.

[0006] There is a need for an efficient method to convert renewable feedstocks into distillate fuels. Summary of the Invention

[0007] We have devised a method for oligomerizing olefin streams into distillate fuels. A method for oligomerizing an olefin stream includes: feeding a light olefin vapor stream containing ethylene into a first oligomerization reactor equipped with a first oligomerization catalyst to produce a first oligomerized stream; and feeding a heavy olefin liquid stream containing C3-C8 olefin liquid into a second oligomerization reactor equipped with a second oligomerization catalyst to produce a second oligomerized stream. Both oligomerized streams can be processed to recover fuel streams together. Attached Figure Description

[0008] Figure 1This is a schematic diagram of the oligomerization section of the method and apparatus of this disclosure.

[0009] Figure 2 This is a schematic diagram of the hydrogenation section of the method and apparatus of this disclosure.

[0010] Definitions

[0011] The term "connectivity" refers to the operative permission for fluid flow between enumerated components, which can be characterized as "fluid connectivity".

[0012] The term "downstream connectivity" means that in downstream connectivity, at least a portion of the fluid flowing toward the body can be operatively flowed from the object with which it is fluidly connected.

[0013] The term "upstream connectivity" means that at least a portion of the fluid flowing out of the main body can be operatively directed to an object in fluid communication with it.

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

[0015] The term "indirect connection" refers to fluid flow from an upstream component entering a downstream component after passing through an intermediary container.

[0016] The term "bypass" means that an object is disconnected from the downstream entity at least within the scope of the bypass.

[0017] As used herein, the terms “major” or “most” mean greater than 50%, appropriately greater than 75%, and preferably greater than 90%.

[0018] The term "tower" refers to one or more distillation columns used to separate one or more components with different volatility. Unless otherwise specified, each column includes a condenser at the top of the column for condensing a portion of the overhead feed and returning it to the top of the column, and a reboiler at the bottom of the column for vaporizing a portion of the bottom feed and returning it to the bottom of the column. The feed to the column can be preheated. The top pressure is the pressure of the vapor at the top of the column at the vapor outlet. The bottom temperature is the liquid temperature at the bottom outlet. Top and bottom lines refer to the net lines from any downstream reflux or reboiler to the column. Stripping columns may omit the reboiler at the bottom of the column and instead provide the heating requirement and power for separating from a fluidized inert medium such as steam. Stripping columns typically feed from the top tray and remove the main product from the bottom.

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

[0020] As used herein, the term “true boiling point” (TBP) refers to a test method conforming to ASTM D-2892 for determining the boiling point of a substance. ASTM D-2892 is used to produce standardized masses of liquefied gases, distillate fractions, and residues for which analytical data are available, and to determine the yield of the aforementioned fractions by both mass and volume, based on which a graph of distillation temperature versus mass% is obtained using fifteen theoretical plates in a column with a reflux ratio of 5:1.

[0021] As used herein, the terms “T5,” “T90,” or “T95” refer to the temperatures at which 5%, 90%, or 95% of a sample by mass (as the case may be) boil using ASTM D-86 or TBP, respectively.

[0022] As used in this article, the term “initial boiling point” (IBP) refers to the temperature at which a sample begins to boil using ASTM D-7169, ASTM D-86, or TBP (as applicable).

[0023] As used in this article, the term “endpoint” (EP) refers to the temperature at which a sample is brought to a complete boil using ASTM D-7169, ASTM D-86, or TBP (as applicable).

[0024] As used herein, the term "diesel" means hydrocarbons that boil in the following ranges: IBP between about 125°C (257°F) and about 175°C (347°F), or T5 between about 150°C (302°F) and about 200°C (392°F), and "diesel fractionation point," including T95 between about 343°C (650°F) and about 399°C (750°F) using the TBP distillation method, or T90 between 280°C (536°F) and about 340°C (644°F) using ASTM D-86. The term "green diesel" means diesel containing hydrocarbons not derived from fossil fuels.

[0025] As used herein, the term "jet fuel" refers to a hydrocarbon that boils in the T10 range between about 190°C (374°F) and about 215°C (419°F) and whose endpoint is between about 290°C (554°F) and about 310°C (590°F). The term "green jet fuel" refers to a jet fuel that contains hydrocarbons not derived from fossil fuels. Detailed Implementation

[0026] The method disclosed in this invention relates to dimerizing a light olefin stream containing ethylene and oligomerizing a heavy olefin stream containing C3+ olefins, followed by combined downstream processing of the oligomerized streams. This method takes into account that ethylene and higher olefins require different treatment methods to achieve efficient oligomerization. The method and apparatus may include, for example... Figure 1 The oligomer segment 10 shown and Figure 2 The hydrogenation section 110 is shown.

[0027] Go to Figure 1 In the oligomerization section 10, the initial olefin stream in feed line 12 can be obtained from a methanol-to-olefins process or an ethanol dehydration process. The initial olefin stream in feed line 12 can be deethanized by fractionation in deethanizer 14 to provide an ethylene stream in the net column overhead line 16 and a fractionated C3+ olefin stream in the net column bottom line 18. The operating bottom temperature of deethanizer 14 can be from approximately 43°C (-110℉) to approximately 104°C (220℉), and the top pressure can be from approximately -2.1 MPa (gauge pressure) (300 psig) to approximately 3.5 MPa (gauge pressure) (500 psig).

[0028] The ethylene overhead stream in overhead line 16 can be heated by heat exchange with the concentrated ethylene stream in line 22 and combined with the hydrogen stream from line 23 to provide a combined ethylene overhead stream, which is further heated and fed into acetylene conversion reactor 20. In acetylene conversion reactor 20, acetylene is converted to ethylene by an acetylene conversion catalyst in the presence of hydrogen, thereby producing a concentrated ethylene stream in line 22. The concentrated ethylene stream in line 22 is condensed by heat exchange with the ethylene overhead stream in overhead line 16 and further condensed. The further condensed concentrated ethylene stream is separated in deethanizer receiver 24 to provide a gaseous light olefin stream in net overhead line 25 and a condensate stream in a reflux line from the bottom of deethanizer receiver 24, which can be refluxed back to deethanizer 14. The deethanized feed stream in the bottom line 18 can be diverted into a reboiler stream, which, after being reboiled, returns to the deethanizer 14 in a boiling state via the reboiler line 19 to provide the required heat. The acetylene conversion catalyst can be an alumina-supported palladium and silver catalyst. Acetylene conversion conditions can include a pressure of approximately 1.4 MPa (gauge pressure) (200 psig) to approximately 2.8 MPa (gauge pressure) (400 psig) and a temperature of approximately 38°C (100℉) to approximately 93°C (200℉).

[0029] The fractionated heavy olefin stream in the bottom line 26 of the clean tower may contain oxygen-containing compounds (such as dimethyl ether) and dienes, which can be removed. Oxygen-containing compounds can be removed by water washing and / or adsorption units, while dienes can be removed by selective hydrogenation.

[0030] The light olefin stream in line 25 and the heavy olefin stream in line 26 can be conveyed to the oligomerization section 10. The light olefin stream in line 25 may contain a significant amount of ethylene. The light olefin stream may primarily contain ethylene. In one aspect, the light olefin stream may contain at least 95 mol% ethylene. The light olefin stream in line 25 may be referred to as the ethylene stream. The light olefin stream may be provided by ethanol dehydration or by an MTO unit. The temperature of the light olefin stream may be from about 60°C (140°F) to about 150°C (302°F), preferably from about 80°C (176°F) to about 100°C (212°F), and the pressure may be from about 3.5 MPa (500 psig), preferably from about 5.6 MPa (800 psig) to about 8.4 MPa (1200 psig). The light olefin stream in line 25 may be expanded by a throttle valve 26 to provide an expanded light olefin stream in line 32. The expanded light olefin stream can be separated in separator 34 (such as a gas-liquid separator) to provide a dry light olefin stream in top line 36 and a light liquid stream in bottom line 38. The light liquid stream in line 38 can be added to the heavy olefin stream in line 26 to provide a supplemented heavy olefin stream in line 40.

[0031] The dry light olefin feed stream in the top line 36 can be compressed to oligomerization pressure to provide a compressed light olefin feed stream in the feed line 42 before being fed into the first oligomerization reactor 50.

[0032] The light olefin feed stream can be contacted with the first oligomerization catalyst to oligomerize ethylene into dimers and oligomers. The oligomerization reaction generates a significant amount of exothermic heat. For example, the dimerization of ethylene can generate 612 kcal / kg (1100 BTU / lb) of heat. Therefore, this significant exothermic heat must be managed. Accordingly, the compressed light olefin feed stream in line 42 can be divided into multiple olefin feed streams. Figure 1In this embodiment, the compressed light olefin feed stream is divided into two independent streams: a first light olefin feed stream in a first light olefin line 42a and a second light olefin feed stream in a second light olefin line 42b. More or fewer independent multi-stream olefin feed streams can be used. Up to six light olefin feed streams are readily conceivable. The light olefin feed stream in line 42 can be divided into equal portions of multi-stream olefin feed streams. Alternatively, the light olefin feed stream in line 42 can be divided into unequal streams. For example, the light olefin feed stream can be divided into streams with decreasing flow rates, wherein the flow rate of the light olefin feed stream fed to the preceding reactor is greater than the flow rate of the light olefin feed stream fed to the following reactor. In one embodiment, the light olefin feed stream is divided into two streams with equal flow rates, each containing 50% of the volume of the light olefin feed stream. In another embodiment, the first light olefin stream in the first light olefin line 42a may constitute about 70% to about 90% of the light olefin stream, and the second light olefin stream in the second olefin line 42b may constitute about 10% to about 30% of the light olefin stream.

[0033] To further manage the exothermic reaction, the light olefin stream can be diluted with a diluent stream to provide a diluted olefin stream to absorb the exothermic reaction. The diluent stream may comprise the alkane stream in diluent line 44. The diluent stream in diluent line 44 can be split into a first diluent stream in line 44a and a second diluent stream in line 44b. The first diluent stream in line 44a can be added to the first light olefin stream in the first light olefin line 42a before it is fed into the first oligomerization reactor 50. Preferably, the first diluent stream is added to the first light olefin stream in line 42a after the light olefin stream in line 42 has been split into multiple olefin streams to provide a first diluted light olefin stream in line 46a, thereby allowing the diluent stream to flow through all catalyst beds 50a and 50b in the first oligomerization reactor 50. Alternatively, the diluent stream can also be split into multiple streams, with each diluent stream added to a corresponding light olefin stream. The volumetric flow rate of the diluent in pipeline 44a can be about 2 to about 8 times, preferably about 3 to about 6 times, the volumetric flow rate of the light olefin in pipeline 42.

[0034] The first diluted olefin feed stream may contain no more than 35% by weight of olefins, suitably no more than 30% by weight of olefins, and preferably no more than 20% by weight of olefins. In one embodiment, the first diluted light olefin feed stream contains about 10% by weight to about 30% by weight of C2 olefins. The first diluted light olefin feed stream may contain no more than 30% by weight of ethylene, suitably no more than 25% by weight of ethylene, and preferably no more than 20% by weight of ethylene. In one embodiment, the first diluted feed olefin feed stream contains no more than about 5% by weight to about 10% by weight of propylene.

[0035] The first oligomerization reactor 50 may include a series of first oligomerization catalyst beds 50a and 50b, each catalyst bed being used to charge diluted light olefin feed streams 46a and 46b, respectively. The first oligomerization reactor 50 preferably includes two fixed first light oligomerization catalyst beds 50a and 50b. It is also conceivable that each first oligomerization catalyst bed 50a and 50b may be located in a dedicated first-stage oligomerization reactor, or that multiple first oligomerization catalyst beds may be located in two or more independent first oligomerization reaction vessels. Up to six first oligomerization catalyst beds are readily conceivable. Figure 1 In this process, a first oligomerization reaction vessel 51 was used, which contained two catalyst beds 50a and 50b.

[0036] When the first oligomer reactor 50 has become deactivated, it can be used in parallel with other first oligomer reactors, during which time the first oligomer reactor 50 is regenerated in situ by burning coke on the catalyst. In another embodiment, each first oligomer reactor may include a main reactor, a secondary reactor, and a backup reactor to facilitate regeneration. Figure 1 These additional reactors are not shown in the diagram.

[0037] The first diluted light olefin feed stream in line 46a can be cooled in a first light feed cooler 47a to provide the cooled diluted first light olefin feed stream in line 48a, which is then fed into the first bed 50a of the first oligomerization catalyst in the first oligomerization reaction vessel 51 of the first oligomerization reactor 50. The feed temperature of the cooled diluted first light olefin feed stream in line 48a can be from about 80°C (176℉) to about 200°C (392℉), and the pressure can be from about 3.5 MPag (500 psig) to about 8.4 MPag (1200 psig). The feed cooler 47a may include a steam generator.

[0038] The diluted first light olefin feed stream can preferably be fed into the first light oligomerizing catalyst bed 50a in line 48a in a downward flow manner. However, an upward flow manner may also be suitable. When ethylene oligomerizes in the first oligomerizing catalyst bed 50a, exothermic reactions occur due to the highly exothermic nature of ethylene oligomerization. The dimerization and oligomerization of the first light olefin feed stream produce the first light oligomer feed stream in the first oligomer feed stream line 52a, whose outlet temperature still rises despite cooling and dilution. The increased outlet temperature is limited to between 150°C (302℉) and approximately 250°C (482℉).

[0039] The second light olefin feed stream in line 42b can be mixed with the first light oligomer feed stream in line 52a, which is drawn from the first light oligomer catalyst bed 50a in the first oligomer reactor 50, to provide a diluted second light olefin feed stream in line 46b. The first light oligomer feed stream in line 52a includes a diluent stream from diluent line 44a, which is added to the first light olefin feed stream in line 42a. The second light olefin feed stream may contain no more than 35% by weight of ethylene, suitably no more than 25% by weight of ethylene, and preferably no more than 20% by weight of ethylene. The diluted second light olefin feed stream in line 46b can be cooled in a second feed cooler 47b (which may be located outside the first oligomer reactor 50) to provide a cooled second light olefin feed stream in line 48b, which is then fed into the second bed 50b of the first oligomer catalyst in the first oligomer reactor 50. The feed cooler 47b may include a steam generator.

[0040] The feed temperature of the diluted second light olefin feed stream in line 48b can be from about 80°C (176℉) to about 200°C (392℉), and the pressure can be from about 3.5 MPa (500 psig) to about 8.4 MPa (1200 psig). The diluted second light olefin feed stream will contain diluent and ethylene from the first oligomer feed stream. Unreacted ethylene from the first oligomer feed stream will undergo dimerization and oligomerization in the second catalyst bed 50b. On the first oligomer catalyst in the second bed 50b, the dimerization and oligomerization of ethylene and oligomers in the second light olefin feed stream produce the second light oligomeric olefin feed stream in line 52b, whose outlet temperature is increased. The increased outlet temperature can be limited to a range of 30°C (54℉) to about 50°C (90℉) higher than the inlet temperature of catalyst bed 50b.

[0041] The first oligomerization reaction mainly takes place in the gas phase or gas-liquid mixture, with an LHSV of 0.5 h for olefin-based reactions. -1 Up to 10h -1 We have found that in all first catalyst beds, ethylene in the light olefin feed stream is typically converted primarily to higher olefins. Ethylene will initially dimerize to butene above the catalyst.

[0042] The first oligomerization catalyst is preferably an amorphous silica-alumina matrix on which metals from Group VIII and / or Group VIB of the periodic table (using the notation of Chemical Abstracts) are supported. In one aspect, the catalyst has a Group VIII metal promoted by a Group VIB metal. Typically, silica and alumina are only in the matrix, so the silica / alumina ratio of the catalyst is the same as that of the matrix. The metal can be impregnated onto the silica-alumina matrix or ion-exchanged with the silica-alumina matrix. Co-milling is also considered. The catalyst used in the present invention can have a low-temperature acidity ratio of at least about 0.15, suitably about 0.2, and preferably greater than about 0.25, as determined by ammonia temperature-programmed desorption (ammonia TPD) as described below. In addition, a suitable catalyst will have an acidity ratio between about 50 m 2 / g and approximately 400m 2 Surface area between / g, as determined by nitrogen BET method.

[0043] The preferred first oligomerizing catalyst comprises an amorphous silica-alumina support. One component of the catalyst support used in this disclosure is alumina. The alumina can be any of various hydrated aluminum oxides or alumina gels such as α-alumina monohydrate with boehmite or boehmite structure, α-alumina trihydrate with trihydrate structure, β-alumina trihydrate with bayerite structure, etc. Particularly preferred alumina can be purchased under the trade name Catapal from Sasol North America Alumina Product Group. This material is a very high purity α-alumina monohydrate (boehmite), which has shown to produce high purity γ-alumina after calcination at high temperatures. The other 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) in Japan.

[0044] In the preparation of the first oligomerizing catalyst used in this invention, another component used is a surfactant. Preferably, the surfactant is mixed with the aforementioned alumina and silica alumina powders. The resulting mixture of surfactant, alumina, 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 invention. Any suitable surfactant can be used according to the invention. Preferred surfactants are those selected from a range 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.

[0045] A suitable silica-alumina mixture is prepared by mixing silica-alumina and alumina in proportional volume to achieve a desired silica to alumina ratio. In one embodiment, about 75% to about 99% by weight of amorphous silica-alumina and about 10% to about 20% by weight of alumina powder will provide a suitable carrier, wherein the silica to alumina ratio is 2.6. In one embodiment, other ratios of amorphous silica-alumina to alumina may be suitable.

[0046] The surfactant can be combined with the mixture of silica alumina and alumina using any convenient method. Preferably, the surfactant is mixed during the mixing and formation of the alumina and silica alumina. A preferred method is to mix an aqueous solution of the surfactant with the blend of alumina and silica alumina prior to the final formation of the carrier. Preferably, the surfactant is present in the paste or dough in an amount from about 0.01% by weight to about 10% by weight, based on the weight of the alumina and silica alumina.

[0047] Monobasic acids such as nitric acid or formic acid can be added to the mixture in the aqueous solution to dissolve the alumina glue in the binder. Additional water can be added to the mixture to provide sufficient moisture, thus forming a dough with a sufficient consistency to be extruded or spray-dried.

[0048] The paste or dough can be prepared in granular form, preferably by extruding a dough mixture of alumina, silica alumina, surfactant, and water through a die having openings of the desired size and shape, followed by breaking the extruded material into extrudates of the desired length and drying. A further calcination step may be employed to impart increased strength to the extrudate. Typically, calcination is carried out in a dry air stream at a temperature of approximately 260°C (500℉) to approximately 815°C (1500℉).

[0049] The extruded particles can have any suitable cross-sectional shape, i.e., symmetrical or asymmetrical, but most commonly have a symmetrical cross-sectional shape, preferably spherical, cylindrical, or lobed. The cross-sectional diameter of the particles can be as small as 40 μm; however, it is typically from about 0.635 mm (0.25 inch) to about 12.7 mm (0.5 inch), preferably from about 0.79 mm (1 / 32 inch) to about 6.35 mm (0.25 inch), and most preferably from about 0.06 mm (1 / 24 inch) to about 4.23 mm (1 / 6 inch).

[0050] The amorphous silica-alumina supports used in this paper are typically characterized by a large total pore volume, average pore size, and surface area sufficient to provide ample space and area for depositing active metal components. As measured by conventional mercury porosimetry, the total pore volume of the support is typically from about 0.2 cc / g to about 2.0 cc / g, preferably from about 0.25 cc / g to about 1.0 cc / g, and most preferably from about 0.3 cc / g to about 0.9 cc / g. Typically, the pore volume of the support in pores with a diameter greater than 100 angstroms is less than about 0.1 cc / g, preferably less than 0.08 cc / g, and most preferably less than about 0.05 cc / g. The surface area, as measured by the BET method, is typically greater than 50 m². 2 / g, for example, above about 200m 2 / g, preferably at least 250m 2 / g, and the optimal value is approximately 300m 2 / g to approximately 400m 2 / g.

[0051] To prepare the first oligomerization catalyst, the support material is composited with one or more precursors of at least one metal component from Group VIII or Group VIB of the periodic table, such as by single or multiple impregnation of calcined amorphous refractory oxide support particles. The Group VIII metal (preferably nickel) should be present at a concentration of about 0.5% to about 15% by weight, and the Group VIB metal (preferably tungsten) should be present at a concentration of about 0 to about 12% by weight. Impregnation can be achieved by any method known in the art (e.g., by spray impregnation), wherein a solution containing the 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% to 10% smaller than the volume that just fills the pore.

[0052] If the active metal precursor is introduced by impregnation, subsequent or secondary calcination at elevated temperatures (such as, for example, between 399°C (750°F) and 760°C (1400°F)) converts the metal into its corresponding oxide form. In some cases, calcination can be performed after each impregnation of the individual active metals. Subsequent calcination produces a catalyst containing the active metal in its corresponding oxide form.

[0053] A preferred first oligomer catalyst disclosed herein has an amorphous silica-alumina matrix impregnated with 0.5% to 15% by weight of nickel, in the form of a 3.175 mm (0.125 inch) extruder, having a density of about 0.45 g / ml to about 0.65 g / ml. It is also envisioned that the metal can be incorporated onto the support by other methods, such as ion exchange and co-milling.

[0054] The first-stage oligomerization catalyst can be regenerated after deactivation. Suitable regeneration conditions include exposing the catalyst (e.g., in situ) in a laboratory setting to hot air containing 7% oxygen at approximately 400°C to approximately 500°C for 24 hours, or to coking in 0.5% oxygen for at least 24 hours under industrial conditions, followed by verification combustion in 7% oxygen until all coke is burned off. To facilitate regeneration without interrupting operations, a swirling bed arrangement can be used in conjunction with an alternative first-stage oligomerization reactor. The regeneration gas may 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.

[0055] Compared to the first light olefin stream in line 25, the concentration of ethylene dimers and oligomers in the light oligomer stream in the second light oligomer stream line 52b is increased. The light oligomer stream in the second light oligomer stream line 52b can be fed into the dealkane tower 70 after depressurization.

[0056] The heavy olefin stream, which includes liquid C3+ olefin stream and replenishes the liquefied heavy olefin stream in line 38, is fed into the second oligomerizing reactor 60 containing the second oligomerizing catalyst to produce a heavy oligomerizing stream. The replenished heavy olefin stream in line 40 can be stored in buffer tank 41. From buffer tank 41, the replenished heavy olefin stream in line 54 can be divided into a first heavy olefin stream in line 54a and a second heavy olefin stream in line 54b to manage the exothermic reaction throughout the second oligomerizing reactor 60. The first heavy olefin stream in the first heavy olefin stream in line 54a can be fed into the first catalyst bed 60a in the second oligomerizing reactor 60, and the second heavy olefin stream in the second heavy olefin stream in line 54b can be fed into the second catalyst bed 60b in the second oligomerizing reactor 60. To further manage the exothermic reaction in the second oligomerization reactor 60, a second diluent stream 44b is added to the first heavy olefin stream in line 54a to provide a diluted heavy olefin stream for feeding into the second oligomerization reactor 60.

[0057] Unreacted olefin stream 78 is recovered in dealkane tower 70 downstream of oligomerization section 10 and then recycled back to second oligomerization reactor 60 in line 56 to produce additional oligomers. The unreacted olefin stream may contain C3 to C8 olefins. The unreacted olefin stream in line 56 can be split into a first unreacted olefin stream in line 56a and a second unreacted olefin stream in line 56b. The first unreacted olefin stream in line 56a can be added to the first heavy olefin stream in the first heavy olefin line 54a before being fed into the second oligomerization reactor 60 to provide an enhanced heavy olefin stream in line 58a. Figure 1 In the illustrated embodiment, the second diluent stream in line 44b and the first unreacted olefin stream in line 56a are merged at the same location with the first heavy olefin stream in line 54a. Therefore, line 58a delivers the enhanced, diluted first heavy olefin stream to the second oligomerizing reactor 60.

[0058] To obtain the most desirable olefin product, the second oligomerization reactor 60 is operated at a temperature of about 180°C (356°F) to about 250°C (482°F). The operating pressure of the second oligomerization reactor 60 is about 2.1 MPa (300 psig) to about 7.6 MPa (1100 psig), and more preferably about 3.5 MPa (500 psig) to about 6.9 MPa (1000 psig).

[0059] The second oligomerization reactor 60 is not downstream connected to the first oligomerization reactor 50 except via diluent line 44 and unreacted olefins line 56. The second oligomerization reactor 60 is preferably operated in a downward flow manner. However, an upward flow manner may also be suitable. The diluted, enhanced first heavy olefin feed stream contacts the second oligomerization catalyst in the first catalyst bed 60a, oligomerizing C3+ olefins to provide distillate-oil range olefins. Regarding the second-stage oligomerization reactor 60, process conditions are selected to produce a higher percentage of spray-range olefins, which, when hydrogenated in subsequent steps described below, produce the desired spray-range hydrocarbon products. The majority of the unconverted olefins fed into the second oligomerization reactor 60 are oligomerized.

[0060] The second oligomer reactor 60 may include a second reactor vessel 61 containing a first bed 60a of a second oligomer catalyst and a second bed 60b of a second oligomer catalyst. A first heavy oligomer feed stream is discharged from the first bed 60a of the second oligomer catalyst in line 62a and cooled. In one embodiment, the cooled first heavy oligomer feed stream in line 62a is mixed with a second heavy olefin feed stream in line 54b to provide a diluted second heavy olefin feed stream in line 58b, which is fed into the second bed 60b of the second oligomer catalyst. In another embodiment, a second unreacted olefin feed stream in line 56b is recycled to the second heavy olefin feed stream in line 54b to provide an enhanced second heavy olefin feed stream in line 58b, which is fed into the second bed 60b of the second oligomer catalyst.

[0061] exist Figure 1 In the illustrated embodiment, the first heavy oligomer stream in line 62a and the second unreacted olefin stream in line 56b are merged at the same location with the second heavy olefin stream in line 54b. Therefore, line 58b delivers the enhanced, diluted second heavy olefin stream to the second catalyst bed 60b in the second oligomer reactor 60. The heavy oligomer stream, having a higher average carbon number than the supplemented heavy olefin feed stream in line 40, exits the second oligomer reactor 60 in line 62b.

[0062] The second oligomerization catalyst may include a zeolite catalyst. The second oligomerization catalyst can be considered a solid acid catalyst. The zeolite may comprise between about 5% by weight and about 95% by weight of the catalyst, for example, between about 5% by weight and about 85% by weight. Suitable zeolites include those with a structure having 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 three-letter codes indicating the class of zeolites are as defined by the Structure Commission of the International Zeolite Association and are maintained at http: / / www.iza-structure.org / databases. UZM-8 is as described in U.S. Patent No. 6,756,030. In a preferred aspect, the first-stage oligomerization catalyst may comprise a zeolite with a framework having a decacyclic porous structure. Examples of suitable zeolites with a decacyclic pore structure include TON, MTT, MFI, MEL, AFO, AEL, EUO, and FER. In a further preferred aspect, the first-stage oligomerizing catalyst comprising a zeolite with a decacyclic pore structure may comprise a one-dimensional pore structure. A one-dimensional pore structure indicates a zeolite containing non-intersecting pores substantially parallel to one of the crystal axes. The pores preferably extend through the zeolite crystal. Suitable examples of zeolites with a decacyclic one-dimensional pore structure may include MTT. In another aspect, the first-stage oligomerizing catalyst comprises MTT zeolite.

[0063] The second oligomerizing catalyst can be formed by combining zeolite with a binder and then shaping the catalyst into granules. The granules may optionally be treated with a phosphorus reagent to produce zeolite with a phosphorus component ranging from 0.5% to 15% by weight of the treated catalyst. The binder is used to impart hardness and strength to the catalyst. Binders include alumina, aluminum phosphate, silica, silica-alumina, zirconium oxide, titanium dioxide, and combinations of these metal oxides, as well as other refractory oxides and clays such as montmorillonite, kaolin, palygorskite, chlorite, and chlorite-palygorskite. Preferred binders are aluminum-based binders, such as alumina, aluminum phosphate, silica-alumina, and clay.

[0064] One component of the catalyst binder used in this disclosure is alumina. The alumina source can be any of various hydrated alumina oxides or alumina gels, such as α-alumina monohydrate with a boehmite or boehmite structure, α-alumina trihydrate with a trihydrate structure, β-alumina trihydrate with a bayonet structure, etc. Suitable alumina can be purchased under the trade name VERSAL from UOP LLC. Preferred alumina can be purchased under the trade name Catapal from Sasol North American Alumina Products Group. The material is a very high purity α-alumina monohydrate (boehmite), which has shown to produce high purity γ-alumina upon calcination at high temperatures.

[0065] A suitable second oligomerizing catalyst is prepared by mixing zeolite and alumina in a volume ratio to achieve the desired zeolite to alumina ratio. In one embodiment, the MTT content can be from about 5% to about 85% by weight, for example, from about 20% to about 82% by weight of MTT zeolite, and the balance of alumina powder will provide a suitably supported catalyst. Silica supports are also considered.

[0066] A monobasic acid, such as nitric acid or formic acid, can be added to the mixture in the aqueous solution to dissolve the alumina gel in the binder. Additional water can be added to the mixture to provide sufficient moisture to form a dough of sufficient consistency for extrusion or spray drying. Extrusion aids, such as cellulose ether powder, can also be added. A preferred extrusion aid is available under the trade name Methocel from Dow Chemical Company.

[0067] The paste or dough can be prepared in granular form, preferably by extruding the dough through a die with openings of the desired size and shape, followed by breaking the extruded material into extrudates of the desired length and drying. A further calcination step may be employed to impart increased strength to the extrudate. Typically, calcination is carried out in an air stream at temperatures from about 260°C (500℉) to about 815°C (1500℉). The MTT catalyst is not selectively treated (e.g., with amines) to neutralize acid sites.

[0068] The extruded particles can have any suitable cross-sectional shape, i.e., symmetrical or asymmetrical, but most commonly have a symmetrical cross-sectional shape, preferably spherical, cylindrical, or lobed. The cross-sectional diameter of the particles can be as small as 40 μm; however, it is typically from about 0.635 mm (0.25 inch) to about 12.7 mm (0.5 inch), preferably from about 0.79 mm (1 / 32 inch) to about 6.35 mm (0.25 inch), and most preferably from about 0.06 mm (1 / 24 inch) to about 4.23 mm (1 / 6 inch).

[0069] In one exemplary embodiment, in one or more catalyst beds of the first-stage oligomer reactor 22, the catalyst provided is an MTT-type zeolite catalyst disposed on a high-purity boehmite alumina matrix, with a loading ratio of about 90 / 10 to about 20 / 80, and preferably between about 20 / 80 and about 50 / 50.

[0070] The second oligomerizing catalyst can be regenerated after deactivation. Suitable regeneration conditions include exposing the second oligomerizing catalyst (e.g., in situ) to hot air containing 7% oxygen at about 400°C to about 500°C for 24 hours. Industrial production may include coking in 0.5% oxygen for at least 24 hours, followed by validation combustion in 7% oxygen until all coke is burned off. For downtime-free regeneration, a switching bed arrangement can be used, equipped with an alternative second oligomerizing reactor. The regeneration gas stream can be fed into the second oligomerizing reactor 60 to be regenerated. The regeneration gas may 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.

[0071] When the second oligomer reactor 60 has become deactivated, it can be used in parallel, during which time the second oligomer reactor 60 is regenerated in situ by burning coke on the catalyst. In another embodiment, the second oligomer reactor 60 may include a main reactor, a secondary reactor, and a backup reactor to facilitate regeneration. Figure 1 These additional reactors are not shown in the diagram.

[0072] Zeolite catalysts are advantageous as a secondary oligomerization catalyst. Zeolite catalysts exhibit relatively low sensitivity to oxygen contamination. Therefore, if produced via an ethanol dehydration process, the olefin feed in line 12 requires only a minimal degree of oxygen-containing compound removal.

[0073] The light oligomer stream in line 52b can be depressurized before being fed into dealkane tower 70. Similarly, the heavy oligomer stream in line 62b can also be depressurized before being fed into dealkane tower 70. In one embodiment, the light and heavy oligomer streams in line 52b are combined to provide a mixed oligomer stream in line 64, which is then depressurized through throttle valve 65 before being fed into dealkane tower 70. The temperature of the mixed oligomer stream in line 64 can be from about 160°C (320℉) to about 190°C (374℉), and the pressure can be from about 3.9 MPa (gauge pressure) (550 psig) to about 7 MPa (gauge pressure) (1000 psig). Additionally, the mixed oligomer stream in line 64 can be directed to a dealkane tower buffer tank (not shown) for appropriate pressure control of the reactor section. The bottom feed flow of the buffer tank is depressurized by a control valve to produce a mixed oligomer flow in line 65.

[0074] We have found that light alkanes (such as ethane and / or propane) are generated in the first oligomerization reactor 50 and / or the second oligomerization reactor 60, and these substances must be removed from the oligomer feed stream. Therefore, the mixed oligomer feed stream in line 64 is subjected to dealkane treatment by fractionation in dealkane tower 70 to provide a light alkane feed stream and a dealkane feed stream. In one embodiment, the light alkane feed stream is an ethane feed stream, in which case the dealkane tower 70 is an ethane removal tower. In another embodiment, the light alkane feed stream is a propane feed stream, in which case the dealkane tower 70 is a propane removal tower. In dealkane tower 70, light alkanes (such as C3- and suitable C2- hydrocarbons) may be separated in the alkane overhead feed stream in overhead line 72 from the dealkane bottom feed stream containing C4+ and suitable C3+ hydrocarbons that may form in bottom line 74. The operating bottom temperature of the dealkane tower 70 can be between 149°C (300°F) and 293°C (-560°F), and the top pressure can be between approximately -345 kPa (gauge pressure) (50 psig) and approximately 1.1 MPa (gauge pressure) (160 psig).

[0075] The alkane overhead stream in the overhead line 72 can be cooled and separated in the dealkane tower receiver 76 to provide dealkane-treated exhaust gas in the exhaust gas line 77. The condensate containing C3-C8 olefins from the dealkane tower receiver 76 can be diverted, with the reflux portion flowing back to the dealkane tower 70 in line 75. A portion of the condensate can be recycled in line 78. The dealkane stream, possibly in the bottom line 74, can be split into a reboiler stream and a net bottom stream: the reboiler stream, in line 80, is reboiled by heat exchange with a first hot diesel fuel stream in line 82, which may be taken from... Figure 1 The jet fuel fractionation column bottom heat exchange feed stream in line 74; the net bottom feed stream in line 84, can be depressurized and heated by heat exchange with the olefin separation column bottom feed stream in line 90, and then fed into the olefin separation column 86. The cooled first stream of hot diesel fuel is removed in line 83. The reboiled bottom feed stream in line 80 can be returned to the dealkane column 70 in a boiling state to provide the required heat. In some embodiments, the feed can be preheated by the olefin separation column bottom feed stream or reboiled by high-pressure steam.

[0076] The dealkane feed stream in the net bottom line 84 of the dealkane tower is fractionated in the olefin separation tower 86 into an overhead olefin stream and a bottom olefin stream: the overhead olefin stream may be in the overhead line 88 of the olefin separation tower, and the bottom olefin stream may be in the bottom line 90 of the olefin separation tower. The overhead olefin stream can be cooled to approximately 31°C (88℉) to approximately 93°C (200℉) to provide a fully condensed stream from the receiver 92 of the olefin separation tower. The overhead olefin condensate from the bottom of the receiver 92 of the olefin separation tower can be divided into a reflux stream and an unreacted olefin recycle stream: the reflux stream is returned to the tower in line 93; the unreacted olefin recycle stream is in recycle line 94 and can be recycled to the second oligomerization reactor 60.

[0077] The first unreacted olefin stream in line 78 and the second unreacted olefin stream in line 94 can be combined to provide a circulating stream in line 96. The circulating unreacted olefin stream in line 96 can be recycled to the second oligomerization reactor 60 to oligomerize C3-C8 olefins. The stream in line 96 can be fed into buffer tank 97 and pumped to split into an entrained stream and an unreacted olefin stream: the entrained stream is in line 98 and is fed into… Figure 2 The hydrogenation unit; unreacted olefin feed stream in line 56.

[0078] The bottom olefin stream in the separator bottom line 90 can be divided into a reboiler stream, which is then reboiled in the separator reboiler line 91 by heat exchange with a second hot diesel stream in line 173 (this second hot diesel stream may be taken from the jet fuel fractionation tower bottom heat exchange stream in line 74), and then returned to the olefin separator 86. The cooled second hot diesel stream in line 162 returns to the separator. Figure 2 The hydrogenation section 110, after reboiling, is returned to the jet fuel fractionation tower 150. The bottom olefin stream in the net bottom line 99 is then transported to... Figure 2 The feed stream, prior to the hydrogenation section 110, can be cooled by heat exchange with the net dealkane bottom stream in line 84. The heavy olefin stream, containing C8+ olefins, can be fed into the hydrogenation section once cooled.

[0079] Go to Figure 2 The hydrogenation section in the middle comes from Figure 2The bottom olefin stream in the net olefins splitter bottom line 99, containing distillate range C8+ oligoolefins, can be hydrogenated in hydrogenation reactor 120 to saturate the olefin bonds and provide fuel. This step ensures that the engine fuel product meets or exceeds the thermal oxidation requirements specified in ASTM D7566-10a for hydrotreated synthetic alkane kerosene (SPK). Additionally, saturating the oligomeric heavy olefins will provide an alkane stream that can be used as a diluent stream in line 44. The bottom olefin stream in line 99 can be cooled to generate vapor and combined with the same... Figure 1 Unreacted olefin streams in line 98, containing C2 to C8 olefins, are combined to produce a combined olefin stream in line 106. The combined olefin stream in line 106 can also be combined with a hydrogen stream in line 108 to provide a combined hydrogenation feed stream in line 112, which, after cooling, is fed into hydrogenation reactor 120 at approximately 125°C (257℉) to 204°C (400℉) and 3.5 MPa (500 psig) to 6.9 MPa (1000 psig). An excess of hydrogen, such as approximately 1.5 to approximately 2.5 stoichiometric amounts, may be used to ensure complete saturation.

[0080] Hydrogenation is typically carried out using conventional hydrogenation or hydrotreating catalysts and may include metal catalysts containing, for example, palladium, rhodium, nickel, ruthenium, platinum, rhenium, cobalt, molybdenum, or combinations thereof, and their supported forms. The catalyst support can be any solid, inert material, including but not limited to oxides such as silica, alumina, titanium dioxide, calcium carbonate, barium sulfate, and carbon. The catalyst support can be in the form of powder, granules, or pellets.

[0081] In one exemplary embodiment, hydrogenation is performed in a hydrogenation reactor 120, which includes an alumina-supported platinum catalyst, for example, about 0.5 wt% to about 0.9 wt% of an alumina-supported platinum catalyst. The hydrogenation reactor 120 converts olefins into alkane products having the same carbon number distribution as the olefins, thereby forming distillate-range alkanes suitable for use as jet fuel and diesel fuel.

[0082] The hydrogenated heavy feed stream discharged from hydrogenation reactor 120 and in line 123 can be separated in thermal separator 122. In thermal separator 122, the hydrogenated heavy feed stream is separated into a hot hydrogenated vapor stream in top line 124 and a hot hydrogenated liquid stream in bottom line 126. The hydrogenated heavy liquid stream in bottom line 126 can be recycled to the diluent stream... Figure 1Before the oligomerization section 10, the liquid is heated by heat exchange with the diluent stream in line 44. The heated hydrogenated heavy liquid stream in the bottom line 126 of the hot tower can be fed into the stripping tower 130. The operating temperature of the heat separator 122 can be from about 204°C (400°F) to about 343°C (650°F), and the pressure can be about 600 psig.

[0083] The hot hydrogenated vapor stream in the top line 124 of the hot tower can be cooled and fed to the cold separator 128. The cold separator 128 separates the hot hydrogenated vapor stream in the top line 124 into a cold vapor stream in the top line 127 and a cold heavy liquid stream in the bottom line 129. These streams can be combined or fed separately to the stripper 130. A purge stream can be drawn from the hydrogenated cold vapor stream in line 127 and fed into line 125; the remainder can be compressed and combined with makeup hydrogen in line 114 to provide the hydrogen stream in line 108. The cold hydrogenated heavy liquid stream in the bottom line 129 can be fed into the stripper 130 at a feed position higher than the feed position of the hot hydrogenated heavy liquid stream in the bottom line 126 of the hot separator. The operating temperature of the cold separator can be approximately 46°C (115°F).

[0084] Stripper 130 may be a flash stripper to remove light gases from the hot hydrogenated liquid stream in hot bottom line 126 and the cold hydrogenated liquid stream in cold bottom line 129. Stripper 130 removes residual gases from the liquid streams to provide a stripper overhead stream in stripper overhead line 132 and a stripped bottom stream in stripper bottom line 134. The stripper overhead stream in stripper overhead line 132 is cooled and separated in stripper receiver 136 to provide a stripper exhaust stream in stripper receiver overhead line 137 and a condensate stream returned to the tower in line 138. The operating bottom temperature of stripping tower 130 can be from about 232°C (450°F) to about 327°C (620°F), and the top pressure can be from about 210 kPa (gauge pressure) (30 psig) to about 700 kPa (gauge pressure) (100 psig).

[0085] The stripped fuel stream in stripping tower bottom line 134, after stripping to remove volatile substances in stripping tower 130, can be fed into jet fuel fractionation tower 150 without further heating. Alternatively, stripping tower 130 can be omitted upstream of jet fuel fractionation tower 150. In jet fuel fractionation tower 150, the stripped fuel stream can be separated into exhaust gas stream in top line 152, green jet fuel stream in side line 154 from one side of jet fuel fractionation tower 150, and green diesel stream in bottom line 156. Liquid product stream from the jet fuel fractionation tower top receiver can also be obtained. The operating bottom temperature of jet fuel fractionation tower 150 can be from about 288°C (550℉) to about 422°C (792℉), and the top pressure can be from about 35 kPa (5 psig) to about 350 kPa (50 psig). The jet fuel fractionation tower overhead stream in overhead line 152 can be cooled, and the resulting condensate is partially returned to jet fuel fractionation tower 150 from jet fuel fractionation tower receiver 158 via line 159. Simultaneously, the net exhaust stream containing C8- hydrocarbons is drawn from receiver tower overhead line 155 of jet fuel fractionation tower receiver 158. The majority of the hydrocarbons in the net exhaust stream in receiver tower overhead line 155 are lighter hydrocarbons and can be used to fuel the reboiler heater 166 of jet fuel fractionation tower 150. To control the flash point of the jet fuel, a net receiver liquid fraction can also be extracted; this fraction can also be used as fuel oil for flame heaters, or alternatively, added to a gasoline pool.

[0086] The green jet fuel stream extracted from side line 154 contains kerosene-range C9-C19 hydrocarbons and can be cooled and extracted as a jet fuel product meeting applicable SPK standards. In an alternative embodiment, instead of returning all condensate to the tower, the green jet fuel stream can be extracted from the condensate stream in line 159 from jet fuel fractionation receiver 158. This green jet fuel stream extracted from line 159 will then require further stripping to remove light fractions. In such embodiments, side line 154 will not be extracted to recover the green jet fuel stream.

[0087] The green diesel bottom feed stream in bottom pipeline 156 can be diverted into a reboiled diesel feed stream in pipeline 157 and a diesel product feed stream in pipeline 164. The reboiled diesel feed stream in pipeline 157 is further divided into a jet fuel bottom heat exchange feed stream 74 and a bypass bottom feed stream in bottom bypass pipeline 161. The jet fuel bottom heat exchange feed stream in jet fuel bottom heat exchange feed stream pipeline 74 can be diverted to provide a first hot diesel feed stream in pipeline 82 and a second hot diesel feed stream in pipeline 173, respectively for supplying... Figure 1The dealkane tower 70 and olefin separation tower 86 provide reboiling heat. The bypass bottom feed stream in line 161 after the valves on it, or from... Figure 1 The second stream of hot diesel fuel, or a combination thereof, cooled in pipeline 162, is taken into jet fuel reboiling pipeline 163, reboiled in flame heater 166, and then sent back to jet fuel fractionation tower 150.

[0088] The diesel product stream in line 164 is split into a diesel product stream in diesel product line 118 and a diluent stream in line 44. Diesel product can also be drawn from the reboiler liquid loop in line 161. The diluent stream in line 44 can be cooled by heat exchange with the hot hydrogenated heavy liquid stream in the bottom line 126 of the thermal separator, and then recycled back to... Figure 1 The oligomerization section 10 is mixed with the light olefin stream in line 42a and the heavy olefin stream in line 54a. The green diesel stream in diluent line 44 is alkane-based, therefore it is inert to the oligomerization and hydrogenation reactions it may undergo. The diesel product stream in diesel product line 118 can be cooled and fed into the diesel pool. The diesel stream will meet the ASTM D975 standard for diesel fuel.

[0089] Starting with olefins, the method disclosed in this invention can efficiently produce green jet fuel and green diesel fuel that meet the requirements of applicable fuels, while managing the generation of exothermic heat. The carbon recovery rate in this method can exceed 95%. Both the jet fuel stream in side line 154 and the diesel product stream in line 118 can be cooled and fed into their respective fuel pools.

[0090] Figure 1

[0091] We simulated the parallel two-stage oligomerization process disclosed in this invention and compared it with a series two-stage process to determine the advantages of the parallel process. The simulated series process first feeds all olefins into the first-stage oligomerization reactor, and then feeds all first-stage oligomers into the second-stage oligomerization reactor. The simulated parallel process feeds all ethylene into the first oligomerization reactor and all C3+ olefins into the second oligomerization reactor. In both cases, the light oligomers are recycled to the second-stage oligomerization reactor.

[0092] With a nickel catalyst weight hourly space velocity of 1 h⁻¹ -1 Simulations revealed a 65% reduction in the volume of nickel catalyst and a 35% reduction in the volume of zeolite catalyst. Additionally, the power consumption of the ethylene compressor decreased by 13%, while the power consumption of the regenerated gas compressor decreased by 11%. The resulting savings from reduced capital expenditure and operating costs were significant and surprising.

[0093] Examples Specific embodiments

[0094] While the following description is presented in conjunction with specific embodiments, it should be understood that the description is intended to be illustrative and not to limit the scope of the foregoing description and the appended claims.

[0095] A first embodiment of this disclosure is a method for oligomerizing an olefin stream, the method comprising: feeding a light olefin stream comprising a gaseous C2 olefin stream into a first oligomerizing reactor containing a first oligomerizing catalyst to produce a light oligomer stream; and feeding a heavy olefin stream comprising a liquid C3+ olefin stream into a second oligomerizing reactor containing a second oligomerizing catalyst to produce a heavy oligomer stream. An embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, wherein the light olefin stream is expanded to provide an expanded light olefin stream, a light liquid stream is separated from the expanded light olefin stream to provide a dried light olefin stream, and the light liquid stream is added to the heavy olefin stream to provide a replenished heavy olefin stream. An embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, wherein the light olefin stream is compressed before being fed into the first oligomerizing reactor to provide a compressed light olefin stream. One embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, wherein the light olefin stream is divided into a first light olefin stream and a second light olefin stream; and the first light olefin stream is fed into a first catalyst bed in the first oligomerization reactor, and the second light olefin stream is fed into a second catalyst bed in the first oligomerization reactor. Another embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, wherein an alkane diluent stream is added to the first light olefin stream to provide a diluted first light olefin stream, and the diluted first light olefin stream is fed into the first oligomerization reactor. One embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, wherein the first light olefin feed stream contacts the first oligomer catalyst in the first catalyst bed of the first oligomer reactor to generate a first light oligomer feed stream; the first light oligomer feed stream is mixed with a second light olefin feed stream to provide a diluted second light olefin feed stream, and the diluted second light olefin feed stream is fed into the second catalyst bed in the first oligomer reactor. Another embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, wherein the supplemented second feed stream is fed into the second oligomer reactor.One embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, wherein the heavy olefin stream is divided into a first heavy olefin stream and a second heavy olefin stream; and the first heavy olefin stream is fed into a first catalyst bed in a second oligomerizing reactor, and the second heavy olefin stream is fed into a second catalyst bed in the second oligomerizing reactor. Another embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, wherein an alkane diluent stream is added to the supplemented heavy olefin stream to provide a diluted heavy olefin stream, and the diluted heavy olefin stream is fed into the second oligomerizing reactor. One embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, comprising: recycling a first unreacted olefin stream to a first heavy olefin stream to provide an enhanced first heavy olefin stream, and feeding the enhanced first heavy olefin stream into the first catalyst bed in a second oligomerizing reactor; and recycling a second unreacted olefin stream to the second heavy olefin stream to provide an enhanced second heavy olefin stream, and feeding the enhanced second heavy olefin stream into the second catalyst bed in the second oligomerizing reactor. Another embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, wherein the first heavy olefin stream contacts a second oligomerizing catalyst in the first catalyst bed of the second oligomerizing reactor to generate a first heavy oligomer stream; the first heavy oligomer stream is mixed with the second heavy olefin stream to provide a diluted second heavy olefin stream, and the diluted second heavy olefin stream is fed into the second catalyst bed in the second oligomerizing reactor. One embodiment of this disclosure is one, any, or all of the embodiments described above to the first embodiment in this paragraph, wherein the first olefin feed stream can be taken from the overhead stream of a deethanizer. One embodiment of this disclosure is one, any, or all of the embodiments described above to the first embodiment in this paragraph, wherein the heavy olefin feed stream can be taken from the bottom stream of a deethanizer. One embodiment of this disclosure is one, any, or all of the embodiments described above to the first embodiment in this paragraph, wherein the first catalyst is a metal catalyst. One embodiment of this disclosure is one, any, or all of the embodiments described above to the first embodiment in this paragraph, wherein the second catalyst is a zeolite catalyst.One embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, and further includes: feeding the light oligomer stream and the heavy oligomer stream into a fractionation column to generate an unreacted olefin stream, and obtaining the first unreacted olefin stream from the unreacted olefin stream. Another embodiment of this disclosure is any or all of the embodiments described above to the first embodiment in this paragraph, and further includes: generating a bottom stream from the fractionation column, fractionating the bottom stream to generate a second unreacted olefin stream, and recycling the second unreacted olefin stream together with the first unreacted olefin stream.

[0096] A second embodiment of this disclosure is a method for oligomerizing an olefin stream, the method comprising: feeding a light olefin stream containing gaseous C2 olefins into a first oligomerization reactor containing a first oligomerization catalyst to generate a light oligomer stream; feeding a heavy olefin stream containing a liquid C3+ olefin stream into a second oligomerization reactor containing a second oligomerization catalyst to generate a heavy oligomer stream; and fractionating the light oligomer stream and the heavy oligomer stream to generate an unreacted olefin stream and a bottom stream from an olefin separation tower. One embodiment of this disclosure is one, any, or all of the embodiments described above to the second embodiment in this paragraph, further comprising recycling the unreacted olefin stream to the heavy olefin stream to provide an enhanced heavy olefin stream, and feeding the enhanced heavy olefin stream into the second oligomerization reactor.

[0097] A third embodiment of this disclosure is a method for oligomerizing an olefin stream, the method comprising: separating a light olefin stream containing gaseous C2 olefins to provide a dry light olefin stream and a light liquid olefin stream; adding the light liquid olefin stream to a heavy olefin stream containing liquid C3+ olefins to provide a supplemented heavy olefin stream; feeding the dry light olefin stream into a first oligomerization reactor containing a first oligomerization catalyst to produce a light oligomer stream; and feeding the supplemented heavy olefin stream into a second oligomerization reactor containing a second oligomerization catalyst to produce a heavy oligomer stream. An embodiment of this disclosure is one, any, or all of the embodiments described above to the third embodiment in this paragraph, further comprising: fractionating the light oligomer stream and the heavy oligomer stream together.

[0098] Although no further detailed description has been provided, it is believed that those skilled in the art can make full use of this disclosure by employing the foregoing description and can readily identify the essential features of this disclosure without departing from the spirit and scope of the invention, and can make various changes and modifications to this disclosure to suit various uses and conditions. Therefore, the foregoing preferred specific embodiments should be understood as illustrative only and not as limiting the remainder of this disclosure in any way, and are intended to cover various modifications and equivalent arrangements included within the scope of the appended claims.

[0099] In the foregoing, all temperatures are expressed in degrees Celsius, and all portions and percentages are by weight unless otherwise specified.

Claims

1. A method for oligomerizing an olefin stream, the method comprising: A light olefin stream containing a gaseous C2 olefin stream is fed into a first oligomerization reactor containing a first oligomerization catalyst to generate a light oligomer stream. as well as A heavy olefin stream containing a liquid C3+ olefin stream is fed into a second oligomerization reactor containing a second oligomerization catalyst to generate a heavy oligomerization stream.

2. The method of claim 1, wherein the light olefin stream is expanded to provide an expanded light olefin stream, a light liquid stream is separated from the expanded light olefin stream to provide a dried light olefin stream, and the light liquid stream is added to the heavy olefin stream to provide a replenished heavy olefin stream.

3. The method of claim 1, wherein the light olefin stream is compressed to provide a compressed light olefin stream before being fed into the first oligomer reactor.

4. The method according to claim 1, wherein the light olefin stream is divided into a first light olefin stream and a second light olefin stream; and the first light olefin stream is fed into a first catalyst bed in the first oligomerization reactor, and the second light olefin stream is fed into a second catalyst bed in the first oligomerization reactor.

5. The method of claim 4, wherein an alkane diluent stream is added to the first light olefin stream to provide a diluted first light olefin stream, and the diluted first light olefin stream is fed into the first oligomerization reactor.

6. The method of claim 5, wherein the first light olefin stream is contacted with the first oligomer catalyst in the first catalyst bed of the first oligomer reactor to generate a first light oligomer stream; the first light oligomer stream is mixed with the second light olefin stream to provide a diluted second light olefin stream, and the diluted second light olefin stream is fed into the second catalyst bed of the first oligomer reactor.

7. The method of claim 2, wherein the supplemented second feed stream is fed into the second oligomer reactor.

8. The method according to claim 1, wherein the heavy olefin stream is divided into a first heavy olefin stream and a second heavy olefin stream; and the first heavy olefin stream is fed into a first catalyst bed in the second oligomerization reactor, and the second heavy olefin stream is fed into a second catalyst bed in the second oligomerization reactor.

9. The method of claim 7, wherein an alkane diluent stream is added to the supplemented heavy olefin stream to provide a diluted heavy olefin stream, and the diluted heavy olefin stream is fed into the second oligomerization reactor.

10. The method of claim 9, the method comprising recycling a first unreacted olefin stream to a first heavy olefin stream to provide an enhanced first heavy olefin stream, and feeding the enhanced first heavy olefin stream into the first catalyst bed in the second oligomerization reactor; and recycling a second unreacted olefin stream to the second heavy olefin stream to provide an enhanced second heavy olefin stream, and feeding the enhanced second heavy olefin stream into the second catalyst bed in the second oligomerization reactor.

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