Process for converting olefins to distillate fuels

Abstract

A process for oligomerizing an olefin stream with an oligomerization catalyst to produce an oligomerized olefin stream. The oligomerization may include a first stage oligomerization step of ethylene followed by a second stage oligomerization of the first stage oligomerized olefins to higher olefins. The oligomerized olefin stream may be separated into a jet fuel stream and a diesel fuel stream. The olefin stream may be obtained by converting oxygenates to olefins with an MTO catalyst.

Description

[Technical field]

[0001] (Statement of priority) This application claims priority to U.S. Provisional Application No. 63 / 325,102, filed March 29, 2022, which is incorporated herein in its entirety.

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

[0003] Carbon dioxide is a so-called greenhouse gas whose concentration is desired by many to be reduced in the atmosphere. Carbon dioxide can be converted to oxygenates such as methanol or dimethyl ether. Molecular sieves, such as microporous crystalline zeolites and non-zeolitic catalysts, particularly silicoaluminophosphate (SAPO), are known to promote the conversion of oxygenates to hydrocarbon mixtures, especially those consisting mostly of light olefins. The highly efficient Methanol to Olefin (MTO) process can convert oxygenates to light olefins that have typically been considered for plastics production. The light olefins produced from the MTO process are highly concentrated in ethylene.

[0004] The oligomerization of light olefins is a process that can effect the conversion of C3-C5 olefins to more desirable products. More specifically, C4 and C5 olefins can be converted to diesel range products or distillates. However, depending on the catalyst, the products from oligomerization can have very poor diesel quality.

[0005] Jet fuel is one of the few petroleum fuels that cannot be easily substituted with electric motor systems because fueling an airplane requires a high energy output that cannot be provided by an electric motor. In certain regions, significant incentives are currently available for green jet fuel.

[0006] An efficient process for converting oxygenates to distillate fuels is desired. Summary of the Invention

[0007] The inventors have developed a process for converting methanol to distillate fuels, comprising contacting an oxygenate stream with an MTO catalyst to produce an olefin stream, and oligomerizing the olefin stream with an oligomerization catalyst to produce an oligomerized olefin stream. The olefin stream may be first oligomerized to provide a first stage oligomerized olefins, which may be further oligomerized to provide a second stage oligomerized olefin stream. The second stage oligomerized olefin stream may be separated into a jet fuel stream and a diesel fuel stream. The olefin stream may be obtained by converting oxygenates to olefins with an MTO catalyst. The ethylene stream may be recycled to the first stage oligomerization step, in embodiments. The oligomerization catalyst may have a silica aluminum oxide support with the aluminum oxide fully dispersed throughout the support. [Brief description of the drawings]

[0008] [Figure 1] FIG. 1 is a schematic diagram of the process and apparatus of the present disclosure. [Diagram 2] FIG. 2 is a schematic diagram of an alternative process and apparatus of the present disclosure. [Diagram 3] FIG. 2 is a schematic diagram of an additional alternative process and apparatus of the present disclosure. [Figure 4] FIG. 2 is a schematic diagram of a further alternative embodiment of FIG. 1; [Diagram 5] FIG. 5 is a schematic diagram of an alternative embodiment of FIG. [Figure 6]FIG. 5 is a schematic diagram of a further alternative embodiment of FIG. 4; DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

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

[0010] The term "downstream communication" means that at least a portion of the fluid flowing to an object with which it is in downstream communication can operatively flow from the object with which it is in fluid communication.

[0011] The term "upstream communication" means that at least a portion of a fluid flowing from an object with which it is in upstream communication can operatively flow to the object with which it is in fluid communication.

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

[0013] The term "indirect communication" means that fluid flow from an upstream component enters the downstream component after passing through an intervening vessel.

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

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

[0016] Detailed Description In the proposed disclosure, the overall process for converting oxygenates to distillate fuels can be divided into several steps: 1. The production of methanol or dimethyl ether can be made from carbon dioxide and hydrogen in a process that is not included in this disclosure but includes a water-gas shift reaction followed by methanol synthesis. 2. Methanol or dimethyl ether can be converted by an MTO process or ethanol can be dehydrated to provide C2-C6 olefins for the following steps. 3. C2-C6 olefins are oligomerized to C9+ distillates including kerosene and diesel. 4. The C9+ olefin distillate can be hydrogenated and split into green jet and green diesel products to meet jet fuel specifications.

[0017] The process and apparatus may include an MTO section 6, an oligomerization section 50, an olefin recovery section 88, and a hydrogenation section 110. Starting with the MTO section 6, the process may include charging an oxygenate stream 10 to an MTO reactor 12 and contacting the oxygenate stream with an MTO catalyst at MTO reaction conditions to convert the oxygenates to olefins and water. The MTO reactor 12 may provide a fluidized catalyst operating at fast flow conditions. The oxygenates may be methanol, dimethyl ether, ethanol, or combinations thereof. The methanol may be derived from the water gas shift reaction of carbon dioxide with hydrogen followed by methanol synthesis.

[0018] The MTO catalyst may be a silicoaluminophosphate (SAPO) catalyst. SAPO catalysts and their formulations are generally taught in U.S. Pat. Nos. 4,499,327(A), 10,358,394, and 10,384,986. The MTO catalyst is preferably SAPO-18, which is more suitable for producing higher olefins than SAPO-34, which is recognized for producing lighter olefins. The increased production of higher olefins with SAPO-18 reduces the throughput of ethylene through oligomerization required to achieve higher distillate yields. The inventors have found that using SAPO-18 at higher pressures increases the butene to ethylene production ratio (B / E) more than SAPO-34, which typically operates at lower pressures.

[0019] MTO reaction conditions include contact with the SAPO catalyst at pressures between 2 MPa and 3.8 MPA, with the partial pressure of methanol being 200-350 kPa for the SAPO-18 catalyst. When using SAPO-34 to catalyze the MTO process, methanol is typically maintained at a relatively low partial pressure of 100-200 kPa. The MTO reaction temperature should be between 325-450°C. The weight hourly space velocity (WHSV) in the MTO reactor should be between 2-15 h. -1 The range is.

[0020] The MTO catalyst is separated from the olefin stream after the MTO reaction, stripped of hydrocarbons with an inert gas such as nitrogen or steam, and transferred in line 16 to a regenerator 14 where air contacts the spent catalyst to burn off the coke from the MTO catalyst.

[0021] The MTO reactor 12 produces a product olefin stream in line 20. The olefin stream in line 20 may be sent to a dehydration column 22 to cool, separate water from the olefin stream, and neutralize acidic compounds to produce a dehydrated olefin stream in line 24 having a reduced water content, and a water stream in line 26 through which the water passes for further processing. The dehydration column 22 may include two columns. The water stream in line 26 may be further processed to separate additional oxygenates that may be recycled to the MTO reactor 12. The dehydrated olefin stream in line 24 may be compressed and sent to a dimethyl ether (DME) recovery unit 30, where it is separated by an absorption column into a deoxygenated olefin stream in line 32 and a DME-rich stream in line 34. The DME-rich stream in line 34 may be recycled to the MTO reactor 10 and converted to olefins over an MTO catalyst. The deoxygenated olefin stream in line 32 may be enriched in C2-C6 olefins.

[0022] The deoxygenated olefin stream in line 32 is oligomerized with an oligomerization catalyst to produce an oligomerized olefin stream comprising C9+ olefins. If the olefin stream has substantial ethylene, in embodiments, the olefin stream may first be contacted with a first stage oligomerization catalyst to oligomerize the ethylene to butenes and higher olefins, and then with a second stage oligomerization catalyst to oligomerize the first stage oligomerized olefins and unconverted feed olefins to C9+ olefins. The contacting order may be reversed.

[0023] The deoxygenated olefin stream in line 32 may be selectively hydrogenated to convert diolefins and acetylene to monoolefins. Hydrogen may be added to the light olefin stream in line 36. The selective hydrogenation reactor 40 typically operates at relatively mild hydrogenation conditions. The light olefin stream is typically maintained under a minimum pressure sufficient to maintain the reactants as liquid phase hydrocarbons. Thus, a wide range of suitable operating pressures range from 2.8 barg (40 psig) to 55 barg (800 psig), or from 3.5 barg (50 psig) to 21 barg (300 psig). Relatively mild temperatures of 25° C. (77° F.) to 350° C. (662° F.), or from 50° C. (122° F.) to 200° C. (392° F.) are typically used. The liquid hourly space velocity of the reactants over the selective hydrogenation catalyst is about 1.0 hr. -1 Super~35.0 hours -1 To avoid undesirable saturation of significant amounts of monoolefinic hydrocarbons, the molar ratio of hydrogen to diolefinic hydrocarbons in the feed entering the selective hydrogenation catalyst bed is maintained between 0.75:1 and 1.8:1.

[0024] Any suitable catalyst capable of selectively hydrogenating the diolefins in the naphtha stream may be used. Suitable catalysts include, but are not limited to, catalysts containing copper and at least one other metal, such as titanium, vanadium, chromium, manganese, cobalt, nickel, palladium, zinc, molybdenum, and cadmium, or mixtures thereof. The metal is preferably supported on an inorganic oxide support, such as, for example, silica and alumina. The selectively hydrogenated olefin stream exits the selective hydrogenation reactor in line 42 and enters a hydroseparator 44 to provide a hydrogen-rich overhead stream in line 46, which may be compressed, possibly supplemented with a make-up hydrogen stream in line 47, and returned as recycled hydrogen in line 36. The selectively hydrogenated liquid feed olefin stream from the bottom of the separator 44 may be transferred to an oligomerization unit 50 in line 48. The feed olefin stream in line 48 may comprise at least 5 wt%, preferably at least 10 wt%, preferably at least 20 wt%, or even at least 25 wt% ethylene. The feed olefin stream in line 48 may comprise ethylene as a major component. The feed olefin stream may also comprise at least 5 wt%, typically at least 10 wt%, preferably at least 20 wt%, or even at least 25 wt% of one or more of C3, C4, C5, and C6 olefins. The feed olefin stream in line 48 may comprise at least 5 wt%, typically at least 10 wt%, preferably at least 20 wt%, or even at least 25 wt% propylene. The feed olefin stream in line 48 may comprise propylene as a major component.

[0025] In other embodiments, the feed olefin stream in line 48 may contain ethylene as a major component and may be referred to as an ethylene stream. In other embodiments, the feed olefin stream in line 48 may contain propylene as a major component and may be referred to as a propylene stream. In other embodiments, the feed olefin stream in line 48 may be a fresh olefin stream that contains ethylene and does not originate from MTO unit 6.

[0026] For example, the feed olefin stream may be provided from an ethanol dehydration process. The ethanol may be derived from any known thermal or biological process. High purity ethanol is not required and aqueous ethanol may be used. For example, the concentration of ethanol may be from 20% to 100%. Ethanol or an ethanol-containing feedstock may be optionally fed to a dehydration reactor, optionally with an inert gas such as nitrogen or steam, preheated to a selected reaction temperature, and passed over a dehydration catalyst (e.g., alumina, modified alumina, silicoaluminate, modified silicoaluminate, and other catalysts) at a temperature and pressure sufficient to carry out the dehydration reaction to form ethylene. The ethanol may be dehydrated at a temperature of 0.1 hr. -1 ~30 o'clock -1 In some embodiments, ethanol may be introduced into the dehydration reactor at a WHSV of 0.5 h. -1 ~5 o'clock -1 The dehydration reactor may be fed at a WHSV of 1000 to 2000° C. The dehydration reactor may be operated at a temperature between 200° C. and 500° C. In some embodiments, the dehydration reactor may be operated at a temperature between 300° C. and 450° C. In some embodiments, the dehydration reactor may be operated at a pressure between 0 barg and 83 barg. In some embodiments, the dehydration reactor may be operated at a pressure between 0 barg and 35 barg. The ethanol conversion may vary between 10% and 100% depending on the operating conditions and the catalyst selected. The ethylene-containing product may be purified to remove water, by-products, oxygen, and other impurities. Purification may include condensing the water and purifying the product through purification adsorbents such as silica, molecular sieves, and carbon. The purified ethanol may be collected or sent directly to the oligomerization unit 50 as a feed olefin stream in line 48.

[0027] The oligomerization unit 50 may include a first stage oligomerization reactor 60 and a second stage oligomerization reactor 70. A light olefin splitter overhead line 106 carrying unconverted feed olefins, a paraffin-containing diluent stream, which may be a net stripped hydrogenation stream in line 148, and a first stage oligomerization recycle stream in line 52 containing first stage oligomerized olefins may be added to the feed olefin stream in line 48 to provide a charge olefin stream in line 54. The charge olefin stream in line 54 may provide a primary charge olefin stream in line 56 and an interbed olefin charge stream in line 58. The primary charge olefin stream in line 56 may be heated and charged to the first stage oligomerization reactor 60.

[0028] The diluent stream may comprise a paraffin stream to absorb heat generated by the oligomerization reaction. The diluent stream may be provided at a diluent to feed ratio of 1:1 to 6:1, preferably 2:1 to 5:1. The diluent stream is preferably a C9+ paraffin stream, which may be taken from the stripper bottoms stream in line 146 downstream of the hydrogenation. The diluent stream may also be a light paraffin stream.

[0029] The main reactor preferably contains two fixed catalyst beds in which most of the ethylene and some of the propylene and higher olefins are also converted. The primary charge olefin stream is charged to the first catalyst bed in line 57, preferably in downflow operation. However, upflow operation may be preferred. As ethylene conversion occurs in the first catalyst bed, heat is generated. Therefore, a suitable heat exchange medium is selected to control the heat of reaction. An interbed olefin charge stream is charged to an interbed location of reactor 60 by line 58 to cool the first stage oligomerization effluent from the first bed. In an embodiment, the effluent from the first catalyst bed may be combined with the interbed olefin charge stream in line 58 and withdrawn from the first catalyst bed, cooled in a heat exchanger such as a steam generator, and returned to the second catalyst bed. The first stage oligomerized olefins exit the first stage oligomerization reactor 60 in line 64.

[0030] The ethylene conversion catalyst is preferably an amorphous silica-alumina base with a metal from either Group VIII and optionally Group VIB of the periodic table using Chemical Abstracts Service notation. In one embodiment, the catalyst has a Group VIII metal promoted with a Group VIB metal. Typically, the silica and alumina are present only in the base, so the ratio of silica to alumina is the same for both the catalyst and the base. The metal can be impregnated onto the silica-alumina base or ion-exchanged into the silica-alumina base. Co-mulling is also contemplated. Additionally, suitable catalysts have a SiO2 content of 300-600 m as determined by nitrogen BET method. 2 / g surface area.

[0031] The most preferred first stage oligomerization catalysts are described below. The preferred ethylene conversion catalyst comprises an amorphous silica-alumina support. The composition of the amorphous silica-alumina ranges from 40 to 99.5 wt% SiO2, suitably at least 70 wt%, preferably 75 wt% to 97.5 wt% SiO2, with the remainder being alumina. More preferably, the composition of the amorphous silica-alumina ranges from 87 to 98 wt% SiO2, with the remainder being alumina.

[0032] Amorphous silica-alumina support for catalysts is most preferably prepared using the oil dripping process, where the synthesis of amorphous silica-alumina and support for catalyst shaping is accomplished in a single continuous manufacturing process as described in U.S. Pat. Nos. 3,909,450, 4,629,717, and 5,139,989. In the oil dripping process, amorphous silica-alumina is synthesized on-line by vigorously mixing silica and alumina precursors to ensure that silica and alumina are atomically interdispersed. An appropriate level of neutralizing agent, such as hexamethylenetetramine (HMT), urea, ammonia, or combinations thereof, is introduced at ambient temperature during the aforementioned mixing step. The resulting premixture is dripped into a hot oil bath via a vibrating dripping head, and the premixture droplets are converted from a sol state to a gel state via the activation of a gelling agent in the hot oil bath. The formed spherical droplets are aged to establish the configuration of the porous structure, followed by water washing under controlled pH to remove residual oil and alkali and alkaline earth metals associated with the silica and alumina precursors that enter the synthesis. Alternatively, the water washing is performed without removing all of the alkali or alkaline earth cations associated with the silica and alumina precursors, providing ion exchange sites for subsequent metal incorporation via an ion exchange procedure. By selecting and controlling the gelling agent level, solids content, and aging conditions including pressure, temperature, and time, the support properties can be tailored to obtain the attributes required for the catalytic process.

[0033] For example, when using amorphous silica-alumina as a support to disperse metals for ethylene conversion, the alumina must be completely dispersed in the silica. This is achieved by oil-drop sol compositions with more than 70-75 wt.% SiO2. If the alumina is not completely dispersed in the silica matrix, the alumina phase will interact with the metal to form spinels, resulting in, for example, loss of ethylene conversion reactivity. Amorphous silica-alumina synthesized according to a given process combines synthesis and shape formation in a single continuous step. Thus, compositional homogeneity is achieved with aluminum oxide completely dispersed in the silica matrix, as shown by the X-ray diffraction pattern indicated by the absence of crystalline alumina phase. In other words, the composition of the support is uniform on the surface of the support throughout the core of the support.

[0034] The size of the spheres can range from 1 / 32 to 1 / 10 inch in diameter, preferably 1 / 20 to 1 / 12 inch in diameter. The porosity of the support as measured by total intrusion volume by mercury porosimetry ranges from 0.40 to 1.4 mL / g, preferably 0.45 to 1.2 mL / g, and most preferably 0.60 to 1.2 mL / g. The porosity should be between 50 and 80%.

[0035] Group VIII, and optionally Group VI and Group IA and Group IIA elements may be incorporated onto the support via ion exchange or impregnation techniques. Group VIII elements preferably include nickel. Group VI elements preferably include chromium, molybdenum, and tungsten. Group IA elements preferably include lithium, sodium, potassium, and combinations thereof. Group IIA elements include magnesium, calcium, strontium, and combinations thereof.

[0036] Another suitable first stage oligomerization catalyst comprises an amorphous silica-alumina support. One of the components of this first stage oligomerization catalyst support utilized in the present disclosure is alumina. The alumina can be any of a variety of hydrous 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, etc. A particularly preferred alumina is available from Sasol North America Alumina Product Group under the trademark "Catapal". This material is an extremely high purity alpha-alumina monohydrate (pseudo-boehmite) that has been shown to yield high purity gamma-alumina after calcination at high temperatures. Another component of the catalyst support is amorphous silica-alumina. The preferred composition of the amorphous silica-alumina ranges from 70 to 99.5% by weight SiO2, with the remainder being alumina. More preferably, the composition of the amorphous silica-alumina ranges from 87-98 wt.% SiO2, with the remainder being alumina. Suitable silica-aluminas having a silica to alumina ratio of 4.0 are available, for example, from CCIC, a subsidiary of JGC, Japan. Suitable silica-aluminas have a silica to alumina molar ratio of 4.0-300 and can be synthesized using a batch or continuous process using cogel or sequential procedures with balanced cations and anions, followed by aging, spray drying, and water washing. Proper aging at a pH of 6-8 is preferred to obtain a pore texture with favorable mass transport properties. The catalyst support has a crystalline alumina phase throughout the support.

[0037] Another component utilized in the preparation of the catalyst utilized in this disclosure is a surfactant. The surfactant is preferably mixed with the alumina and silica-alumina powders described above. The resulting surfactant, alumina, and silica-alumina mixture is then formed, dried, and calcined as described below. Calcination effectively removes the organic components of the surfactant by combustion, but only after the surfactant has faithfully performed its function according to the present invention. Any suitable surfactant may be utilized according to the present invention. A preferred surfactant is a surfactant selected from the series of commercially available surfactants sold under the trademark "Antarox" by Solvay SA. Antarox surfactants are generally characterized as modified linear aliphatic polyethers and are low foaming biodegradable detergents and wetting agents.

[0038] Suitable silica-alumina mixtures are prepared by mixing proportional volumes of silica-alumina and alumina to achieve the desired silica to alumina ratio. In embodiments, 75-95 wt.% amorphous silica-alumina having a silica to alumina ratio of 4.0 to 300 and 10-20 wt.% alumina powder provide a suitable support. In embodiments, other ratios of amorphous silica-alumina to alumina may be suitable.

[0039] Any convenient method may be used to incorporate the surfactant into the mixture of silica-alumina and alumina. The surfactant is preferably mixed during the mixing and formation of the alumina and silica-alumina. A preferred method is to mix an aqueous solution of the surfactant with the alumina and silica-alumina blend prior to final formation of the carrier. The surfactant is preferably present in the paste or dough in an amount of 0.01 to 10% by weight, based on the weight of the alumina and silica-alumina.

[0040] A monobasic acid such as nitric acid or formic acid can be added to the mixture in an aqueous solution to peptize the alumina in the binder. Additional water can be added to the mixture to provide sufficient wetness to form a dough with sufficient consistency to be extruded or spray dried.

[0041] The paste or dough may be prepared in the form of shaped particles, the preferred method being to extrude a dough mixture of alumina, silica-alumina, surfactant, and water through a die having openings of the desired size and shape, after which the extruded material is divided into extrudates of the desired length and dried. A further calcination step may be used to provide additional strength to the extrudates. Generally, calcination is carried out in a stream of dry air at a temperature between 260°C (500°F) and 815°C (1500°F).

[0042] The extruded particles may have any suitable cross-sectional shape, i.e., symmetric or asymmetric, but in most cases will have a symmetric cross-sectional shape, preferably spherical, cylindrical, or polylobal. The cross-sectional diameter of the particles may be as small as 40 μm. However, it will usually be between 0.79 mm (1 / 32 inch) and 6.35 mm (0.25 inch), and most preferably between 0.06 mm (1 / 24 inch) and 4.23 mm (1 / 6 inch).

[0043] Typical properties of the amorphous silica-alumina supports utilized herein are total pore volume, average pore diameter, and surface area large enough to provide substantial space and area for depositing the active metal components. The total pore volume of the support, as measured by conventional mercury porosimetry, is typically 0.2-2.0 cc / g, preferably 0.25-1.0 cc / g, and most preferably 0.3-0.9 cc / g. The surface area, as measured by the BET method, is typically less than 50 m. 2 / g, e.g. 200m 2 / g, preferably at least 250m 2 / g, most preferably 300m 2 g~550m 2 / g.

[0044] The most preferred first stage oligomerization catalyst is a calcined amorphous refractory oxide support particle in the form of oil globules and contains a metal from Group VIII in a molar concentration of M8 / Al2 (M8=Group VIII) of 0.04 to 0.70. The Group VIII element is preferably nickel. The catalyst preferably contains a Group IA element, and optionally a Group IIA element, with M IA / Al(M IA = Group IA) is greater than 0.04 and less than 0.4. Optionally, the catalyst also contains a metal from Group VIB of the periodic table, preferably tungsten, in a concentration of 0 to 12 wt. %.

[0045] The incorporation of the metals can be accomplished by any method known in the art, for example, by ion exchange, evaporation impregnation, or pore filling / spray impregnation. Ion exchange of the Group VIII elements is carried out using an inorganic precursor such as nickel nitrate in the absence or presence of a complexing agent such as ethylene diamine (EDA). Ion exchange is carried out first with an alkali nitrate, followed by ion exchange with nickel nitrate. Ion exchange can be carried out with a solution containing both nickel and an alkali nitrate. Impregnation can be carried out in a sequential or simultaneous manner, preferably with a solution to support ratio of greater than 1.0 by volume to ensure that the ion exchange process takes place in the impregnation operation.

[0046] Following incorporation of the metals, the catalyst is subjected to a heat treatment in a flowing inert gas, such as helium and nitrogen, or an oxidizing gas, such as air, which is carried out at a temperature above 300° C. and below 700° C., preferably above 400° C. and below 550° C., for a period of 30 minutes to 12 hours to decompose the metal precursors and remove physically and chemically adsorbed water.

[0047] The first stage oligomerization catalyst may be regenerated upon deactivation. Suitable regeneration conditions include subjecting the catalyst, for example in situ, to hot air at 450-550° C. for three hours. To facilitate regeneration without downtime, a swing bed configuration is used with an alternative first stage oligomerization reactor 60′. When the first stage oligomerization catalyst in the first stage oligomerization reactor 60 is deactivated, the valves on the primary charge line 57 and the interbed charge line 58 to the first stage oligomerization reactor 60 are closed, and the valves on the alternative primary charge line 57′ and the alternative interbed charge line 58′ are opened to charge an olefin stream to the alternative first stage oligomerization reactor 60′. The regeneration gas stream from line 62 is then introduced into the first stage oligomerization reactor 60 requiring regeneration. The regeneration gas may comprise air with an increased or decreased oxygen concentration. An alternative first stage oligomerization reactor 60' can be regenerated in the reverse manner by closing the valves on lines 57' and 58' and introducing regeneration gas through line 62'. Each first stage oligomerization reactor 60, 60' can include a vent line 63, 63', respectively, for exhausting the regeneration flue gas. The activity and selectivity of the regenerated catalyst are comparable to that of fresh catalyst.

[0048] A first stage oligomerized olefin stream in line 66 with an increased butene concentration compared to the charged olefin stream in line 54 is collected from lines 64 and 64' and split into a first stage oligomerization recycle stream in line 52 and a charged second stage oligomerized olefin stream in line 68. An intermediate olefin stream in a net light olefin splitter bottoms line 108 containing C3 to C8 olefins and an oligomerization recycle stream in line 72 containing oligomerized olefins may be added to the charged second stage oligomerized olefin stream in line 68 to provide a charged second stage oligomerized olefin stream in line 74. The charged second stage oligomerized olefin stream in line 74 may be cooled and charged to the second stage oligomerization reactor 70 in line 75. The second stage oligomerization reactor 70 may be in downstream communication with the first stage oligomerization reactor 60. The second stage oligomerization reactor 70 is preferably operated in downflow operation; however, upflow operation may be preferred. The charged second stage oligomerization olefin stream is contacted with the second stage oligomerization catalyst to oligomerize C2-C8 olefins to provide distillate range olefins. A second stage oligomerization stream having an increased average carbon number over the charged second stage oligomerization olefin stream in line 74 exits the second stage oligomerization reactor 70 in line 76.

[0049] The second stage oligomerization catalyst may comprise a zeolite catalyst. The zeolite may comprise 5 to 95 wt. % of the catalyst, for example 5 to 85 wt. %. Suitable zeolites include zeolites having a structure from one of the following classes: MFI, MEL, ITH, IMF, TUN, FER, BEA, FAU, BPH, MEI, MSE, MWW, UZM-8, UZM-8HS, UZM-37, MOR, OFF, MTW, MRE, MFS, TON, MTT, AFO, ATO, and AEL. Three letter codes for zeotypes are as defined by the Structure Commission of the International Zeolite Association and maintained at http: / / www.iza-structure.org / databases. UZM-8 is as described in U.S. Pat. No. 6,756,030. In a preferred embodiment, the oligomerization catalyst may include a zeolite having a framework with a 10-ring pore structure. Examples of suitable zeolites having a 10-ring pore structure include TON, MTT, MFS, MRE, MFI, MEL, AFO, AEL, EUO, and FER. In a further preferred embodiment, the second stage oligomerization catalyst comprising a zeolite having a 10-ring pore structure may include a one-dimensional pore structure. One-dimensional pore structure refers to a zeolite containing non-intersecting pores that are substantially parallel to one of the axes of the crystal. The pores preferably extend through the zeolite crystal. A suitable example of a zeolite having a 10-ring one-dimensional pore structure may include MTT. In a further embodiment, the oligomerization catalyst comprises MTT zeolite. The preferred silica to alumina ratio of MTT zeolite is 30-100.

[0050] The second stage oligomerization catalyst may be formed by combining the zeolite with a binder and then forming the catalyst into pellets. The pellets may be optionally treated with a phosphorus reagent to produce a zeolite having a phosphorus component of 0.5-15% by weight of the treated catalyst. The binder is used to impart hardness and strength to the catalyst. Binders include alumina, aluminum phosphate, silica, silica-alumina, zirconia, titania, and combinations of these metal oxides, as well as other refractory oxides and clays such as montmorillonite, kaolin, palygorskite, smectite, and attapulgite. Preferred binders are aluminum-based binders such as alumina, aluminum phosphate, silica-alumina, and clays.

[0051] One of the components of the catalyst binder utilized in this disclosure is alumina. The alumina source can be any of a variety of hydrous 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, etc. A suitable alumina is available from UOP LLC under the trademark "VERSAL". A preferred alumina is available from Sasol North America Alumina Product Group under the trademark "Catapal". This material is an extremely high purity alpha-alumina monohydrate (pseudo-boehmite) that has been shown to yield high purity gamma-alumina after calcination at high temperatures.

[0052] A suitable second stage oligomerization catalyst is prepared by mixing proportional volumes of zeolite and alumina to achieve the desired zeolite to alumina ratio. In an embodiment, the zeolite content may be 5-90, e.g., 10-85 wt.%, preferably 25-75 wt.%, zeolite, with the remaining alumina powder providing a suitable supported catalyst. Silica supports are also contemplated. In one exemplary embodiment, an MTT type zeolite catalyst is provided in the oligomerization reactor 70, disposed on a high purity pseudoboehmite alumina substrate in a ratio of 10 / 90-90 / 10, preferably 25 / 75-75 / 25.

[0053] A monobasic acid such as nitric acid or formic acid may be added to the mixture in aqueous solution to peptize the alumina in the binder. Additional water may be added to the mixture to provide sufficient wetness to form a dough with sufficient consistency to be extruded or spray dried. An extrusion aid such as cellulose ether powder may also be added. A preferred extrusion aid is available from The Dow Chemical Company under the trademark "Methocel".

[0054] The paste or dough may be prepared in the form of shaped particles, the preferred method being to extrude the dough through a die having openings of the desired size and shape, followed by dividing the extruded material into extrudates of the desired length and drying. A further calcination step may be used to provide additional strength to the extrudates. Generally, calcination is carried out in a stream of air at a temperature between 260°C (500°F) and 815°C (1500°F). The MTT catalyst is not selective to neutralize acidic sites such as amines.

[0055] The extruded particles may have any suitable cross-sectional shape, i.e., symmetric or asymmetric, but in most cases will have a symmetric cross-sectional shape, preferably spherical, cylindrical, or multilobal. The cross-sectional diameter of the particles may be as small as 40 μm. However, it is preferably between 0.79 mm (1 / 32 inch) and 6.35 mm (0.25 inch), and most preferably between 1.06 mm (1 / 24 inch) and 4.23 mm (1 / 6 inch).

[0056] For oligomerization reactor 70, process conditions are selected to produce a higher percentage of jet range olefins which, when hydrogenated in a subsequent step as described below, result in the desired jet range hydrocarbon products. In an exemplary embodiment, an MTT type zeolite catalyst disposed on a high purity pseudoboehmite alumina substrate in a ratio of 70 / 30 to 90 / 10, preferably 75 / 25 to 85 / 15, is provided in oligomerization reactor 70. The first stage oligomerization olefin stream charged in line 74 is cooled and charged to oligomerization reactor 70. To obtain the most desirable olefin products, oligomerization reactor 70 is operated at a temperature of 100° C. to 270° C., more preferably 111° C. to 230° C. The second stage oligomerization reactor 70 is operated at a pressure between 21 barg (300 psig) and 69 barg (1000 psig), more preferably between 49 barg (710 psig) and 63 barg (900 psig).

[0057] When the second stage oligomerization reaction is carried out according to the above process conditions, a C4 olefin conversion of 95% or more, or even 97% or more, is achieved. The resulting second stage oligomerized olefin stream in line 76 contains a plurality of olefin products that are distillate range hydrocarbons. In general, the nickel-containing catalyst operates at a relatively lower temperature than the zeolite catalyst. It is preferred to operate the nickel-containing catalyst at an inlet temperature that is 50-150°C lower than the inlet temperature of the zeolite-containing catalyst depending on the ethylene content of the feed. Furthermore, the contact time of the nickel-containing catalyst is shorter than the contact time of the zeolite catalyst. Thus, it is preferred that the contact time for the nickel-containing catalyst is less than 45% of the total contact time, and most preferably less than 35%.

[0058] The second stage oligomerization catalyst may be regenerated upon deactivation. Suitable regeneration conditions include subjecting the oligomerization catalyst, for example in situ, to hot air at 500° C. for three hours. To facilitate regeneration without downtime, a swing bed configuration is used with the alternative second stage oligomerization reactor 70′. When the second stage oligomerization reactor 70 is shut down, a valve on the cooled charge line 75 to the second stage oligomerization reactor 70 is closed and a valve on the alternative cooled charge line 75′ is opened to charge the second stage charge oligomerized olefin stream to the alternative second stage oligomerization reactor 70′. A regeneration gas stream from line 78 is then introduced to the second stage oligomerization reactor 70 requiring regeneration. The regeneration gas may include air with increased or decreased oxygen concentration. The alternative second stage oligomerization reactor 70' can be regenerated in the reverse manner by closing the valve on line 75' and introducing regeneration gas through line 78'. Each second stage oligomerization reactor 70, 70' can include a vent line 73, 73', respectively, for exhausting the regeneration flue gas. The activity and selectivity of the regenerated catalyst are comparable to that of fresh catalyst.

[0059] A second stage oligomerization olefins stream in line 80 collected from lines 76, 76' having an increased concentration of C9+ olefins compared to the second stage oligomerization olefins stream charged in line 74 is split into a second stage oligomerization recycle stream in line 72 and a second stage oligomerization product stream in line 84.

[0060] The second stage oligomerization product stream from oligomerization unit 50 is conveyed to olefin recovery section 88 where it is fed to heavy olefin splitter column 90. In heavy olefin splitter column 90, oligomers boiling lower than jet range hydrocarbons, typically C8-hydrocarbons having atmospheric boiling points below 150° C., are separated into a net heavy olefin splitter overhead stream in line 92 from a net heavy olefin splitter bottoms stream in line 94 containing distillate range C9+hydrocarbons, typically C9 to C20 olefins. Heavy olefin splitter column 90 may operate at a bottoms temperature of 300° C. to 500° C. and an overhead pressure of 9 barg to 15 barg. The heavy olefin splitter overhead stream may be condensed and removed as a net vapor overhead stream from heavy olefin splitter receiver 96 in line 92.

[0061] The C8- net vapor heavy olefins splitter overhead stream in line 92 may be fed to light olefins splitter column 100. The light olefins splitter overhead stream may be cooled and separated in light olefins splitter overhead receiver 102 into a net vapor overhead stream in line 104 containing methane and light off-gas, and a net liquid light olefins splitter overhead stream in line 106 containing ethylene. The net liquid light olefins splitter overhead stream in line 106, together with a diluent stream in line 148 and a first stage oligomerization recycle stream in line 52, may recycle unconverted ethylene to the feed olefins stream in line 48 to provide a charge olefins stream in line 54 to the first stage oligomerization reactor 60. The net liquid light olefin splitter overhead stream may comprise primarily ethylene which may be recycled to the first stage oligomerization reactor 60. In an alternative embodiment, the net liquid olefin splitter overhead stream may comprise primarily propylene which may be recycled to the first stage oligomerization reactor 60. In this case, ethylene is still captured in the net liquid light olefin splitter overhead stream in line 106. An intermediate olefin stream in net light olefin splitter bottoms line 108 comprising C3-C8 or C4-C8 olefins may be recycled to combine with the charged second stage oligomerized olefin stream in line 68 along with an oligomerization recycle stream in line 72 to provide a charged second stage oligomerized olefin stream in line 74 which is oligomerized in the second stage oligomerization reactor 70. A drag stream may be removed from lines 106 and 108. The light olefin splitter column 100 may operate at a bottoms temperature between 100° C. and 300° C. and an overhead pressure between 5 barg and 11 barg.

[0062] The net heavy olefin splitter column bottoms stream in line 94 containing distillate range C9+ olefins may be hydrogenated to provide motor fuel for saturating olefinic bonds in hydrogenation reactor 120. This step is performed to ensure that the product motor fuel meets or exceeds the thermo-oxidation requirements set forth in ASTM D7566-10a for hydrotreated synthetic paraffinic kerosene (SPK) and applicable requirements for diesel. Hydrogenation is typically performed using conventional hydrogenation or hydrotreating catalysts, which may include, for example, metal catalysts containing palladium, rhodium, nickel, ruthenium, platinum, rhenium, cobalt, molybdenum, or combinations thereof, and supported metal catalysts thereof. The catalyst support may be any solid, inert material, including, but not limited to, oxides such as silica, alumina, titania, calcium carbonate, barium sulfate, and carbon. The catalyst support may be in the form of a powder, granules, pellets, and the like. A hydrogen stream is provided in line 122 to the hydrogenation reactor 120 as a hydrogen source.

[0063] In an exemplary embodiment, hydrogenation is carried out in hydrogenation reactor 120 comprising an alumina supported platinum catalyst, for example, 0.5% to 0.9% by weight platinum catalyst on alumina. Using this catalyst, hydrogenation suitably occurs at temperatures between 125° C. and 175° C. and pressures between 35 barg (500 psig) and 105 barg (1500 psig). In accordance with these process conditions, hydrogenation reactor 120 converts olefins to paraffinic products having the same carbon number distribution as the olefins, thereby forming paraffins in the distillate range suitable for use as jet and diesel fuels.

[0064] The hydrogenated distillate stream discharged in line 124 from the hydrogenation reactor 120 may be cooled and fed to the hydrogenation separator 130. In the hydrogenation separator 130, the hydrogenated distillate stream is separated into a hydrogenation separator vapor stream in an overhead line 132 and a hydrogenation separator liquid stream in a bottoms line 134. The hydrogenation separator vapor stream in line 132 may be compressed and combined with make-up hydrogen in line 136 to feed the hydrogen stream in line 122 and / or the hydrogen stream in line 36 to the selective hydrogenation reactor 40. The hydrogenation separator liquid stream in bottoms line 134 may be heated by heat exchange with the hydrogenated distillate stream in line 124 and fed to the stripper column 140.

[0065] Stripper column 140 strips light gases from the hydroseparator liquid stream to provide a stripper off-gas stream in off-gas line 142 from stripper overhead receiver 144. The net stripped hydrotreated stream in stripper bottoms line 146 is split into a product fuel stream in line 147 and a diluent stream in line 148. The diluent stream in line 148 may be recycled to the feed olefins stream in line 48 along with the light olefins splitter overhead stream 106 and the first stage oligomerization recycle stream in line 52 to provide the charge olefins stream in line 54. The diluent stream is inert in the oligomerization reactors and serves to absorb exothermic heat in the first stage oligomerization reactors 60, 60' and the second stage oligomerization reactors 70, 70'. The stripper column 140 may operate at a bottoms temperature between 250° C. and 500° C. and an overhead pressure between 2 barg and 8 barg.

[0066] The product fuel stream in line 147 may be fed to a jet fractionation column 150 for separation into an off-gas stream in overhead line 152 from a jet receiver overhead 154, a green jet fuel stream from a jet receiver bottom line 156, and a green diesel stream in a net diesel bottoms line 158. Both the jet fuel stream in line 156 and the diesel stream in line 158 may be fed to respective fuel pools. The jet fractionation column 150 may operate at a bottoms temperature of 350° C. to 600° C. and an overhead pressure of 1 barg to 5 barg.

[0067] In the alternative embodiment in Figure 2, only a single olefin splitter column 90# is used in the olefin recovery section 88#. Elements in Figure 2 that have the same configuration as Figure 1 will have the same reference numbers as Figure 1. Elements in Figure 2 that have a different configuration than the corresponding elements in Figure 1 will have the same reference numbers but are designated with a hash tag symbol (#). The configuration and operation of the embodiment of Figure 2 is essentially the same as Figure 1, with the following exceptions.

[0068] The olefin splitter column 90# in FIG. 2 separates dimers and oligomers boiling lower than the jet range hydrocarbons, typically C8-hydrocarbons having atmospheric boiling points below 150° C., in the net olefin splitter overhead stream in line 92# from the net olefin splitter bottoms stream in line 94# containing the distillate range C9+hydrocarbons, typically C9 to C20 olefins. The olefin splitter column 90# may operate at a bottoms temperature of 100° C. to 500° C. and an overhead pressure of 5 barg to 15 barg. The olefin splitter overhead stream may be condensed and removed as a net liquid overhead stream from the olefin splitter receiver 96 in line 106#. The net liquid olefin splitter overhead stream in line 106#, together with the diluent stream in line 148 and the first stage oligomerization recycle stream in line 52, can recycle unconverted ethylene and oligomers to the feed olefin stream in line 48 to provide a charge olefin stream in line 54 to the first stage oligomerization reactor 60. The net liquid light olefin splitter overhead stream in line 106# can contain at least 10 wt%, suitably at least 20 wt%, preferably at least 25 wt% ethylene, which can be recycled to the first stage oligomerization reactor 60. In the embodiment of FIG. 2, the intermediate olefin stream in the net olefin splitter bottom line 108 containing C3-C8 olefins is not returned to the oligomerization reactor 70 in the oligomerization olefin stream in line 74. 2 embodiment, the light and intermediate olefins in the net liquid overhead stream may be recycled to oligomerization unit 50 in line 106#. Alternatively, olefin splitter column 90# may be operated to recycle light olefins, C2-C3 olefins, or C2-C4 olefins back to oligomerization unit 50 in line 106# while sending heavier C4+ or C5+ oligomers to hydrogenation in bottoms line 94#.

[0069] In an alternative embodiment in FIG. 3, even if only a single olefin splitter column 90# is used in olefin recovery section 88#, the intermediate olefin stream is returned to second stage oligomerization reactor 70. Elements in FIG. 3 that have the same configuration as FIG. 2 will have the same reference numbers as FIG. 2. Elements in FIG. 3 that have a different configuration than the corresponding elements in FIG. 2 will have the same reference numbers but with an asterisk symbol ( * ) The structure and operation of the embodiment of FIG. 3 is essentially the same as that of FIG. 2, with noted exceptions.

[0070] Olefins Recovery Section 88 * Olefin splitter column in 90 * is taken from the side of the column through line 108 * The side cut stream in the olefin splitter side line 108 contains C3 to C8 olefins. * The intermediate olefins stream in line 68 may be recycled to combine with the oligomerization recycle stream in line 72 to provide a second stage oligomerized olefins stream in line 74 that is oligomerized in the second stage oligomerization reactor 70. Additionally, the olefin splitter overhead stream is condensed and passed through line 106. * The olefin splitter receiver 96 may be removed as a net liquid overhead stream from line 106. * The net liquid olefin splitter overhead stream in line 106, together with the diluent stream in line 148 and the first stage oligomerization recycle stream in line 52, may recycle unconverted ethylene to the feed olefin stream in line 48 to provide the charge olefin stream in line 54 to the first stage oligomerization reactor 60. * The net liquid light olefins splitter overhead stream in may comprise primarily ethylene which may be recycled to the first stage oligomerization reactor 60.

[0071] Figure 4 illustrates an embodiment in which oligomerization reactor 60+ contains beds of both a first stage oligomerization catalyst and a second stage oligomerization catalyst. Elements in Figure 4 that have the same configuration as Figure 2 will have the same reference numbers as Figure 2. Elements in Figure 4 that have a different configuration than the corresponding elements in Figure 2 will have the same reference numbers but are designated with a plus sign (+). The configuration and operation of the embodiment of Figure 4 is essentially the same as Figure 2.

[0072] The oligomerization reactor 60+ includes catalyst beds in which both the first stage oligomerization catalyst and the second stage oligomerization catalyst of the previous embodiment are either stacked or intermingled in the same catalyst bed, so that both ethylene conversion and higher olefin conversion occur in the same catalyst bed. In one embodiment, the first stage oligomerization catalyst can be stacked on top of the oligomerization catalyst so that the charge olefins encounter the first stage oligomerization and then the second stage oligomerization. The first stage oligomerization stream, which includes the first stage oligomerized olefins and the unconverted feed olefins, is collected in line 66+ from the reactors 60+ and 60'+. The oligomerization recycle stream in line 52+, which includes the dimers, oligomers, and unconverted feed olefins, is recycled to the oligomerization reactor 60+ via the charge olefins stream in line 54. The oligomerization product stream in line 84+ may be fed to an olefin splitter column 90# for separation. The net liquid light olefin splitter overhead stream in line 106# and / or the oligomerization recycle stream in line 52+ may contain at least 10 wt.%, suitably at least 20 wt.%, and preferably at least 25 wt.%, ethylene, which may be recycled to the first stage oligomerization reactor 60.

[0073] Figure 5 shows an embodiment in which the first stage oligomerized olefin stream in line 66+ is fed to an interstage separator 160 for separation. Elements in Figure 5 having the same configuration as Figure 4 will have the same reference numbers as Figure 4. Elements in Figure 5 having a different configuration than the corresponding elements in Figure 4 will have the same reference numbers but are designated with a prime symbol (-). The configuration and operation of the embodiment of Figure 5 is essentially the same as Figure 4, with noted exceptions.

[0074] Reactor 60+ may contain a first stage oligomerization catalyst, similar to the embodiment of FIG. 1. The first stage oligomerization stream in line 66- may be fed to interstage separator 160 to separate an overhead olefins stream containing unreacted ethylene in overhead line 162 from a bottoms first stage oligomerized olefins stream in bottoms line 164. Separator 160 may operate at a temperature of 250-300° C. and a pressure of 40-90 barg. The ethylene in the overhead olefins stream in line 162 may be mixed with the primary charge olefins stream in line 56- and charged to first stage oligomerization reactor 60+ or ​​60'+ in line 57 or 57', respectively. The overhead olefins stream in line 162 comprises ethylene as a major component. The bottoms first oligomerized olefins stream in bottoms line 164 may be mixed with unoligomerized recycle olefins in olefin splitter overhead line 106- and the oligomerized recycle stream in line 72- containing oligomerized olefins 72 to provide a second stage oligomerized olefins stream charged in line 74- which may be charged to the second stage oligomerization olefins reactor 70 or 70'-. The second stage oligomerization olefins reactor 70-, 70'- may contain a second stage oligomerization catalyst. The second stage oligomerization product stream in line 84- may be fed to the olefin splitter column 90# for separation.

[0075] Figure 6 shows an embodiment in which the input olefin stream in line 54^ is charged to downstream catalyst beds 65, 67, 65', 67' while recycled ethylene is charged to catalyst beds 61, 61' in upstream reactors 60+, 60'+. Elements in Figure 6 having the same configuration as Figure 4 will have the same reference numbers as Figure 4. Elements in Figure 6 having a different configuration from the corresponding elements in Figure 4 will have the same reference numbers but are designated with a carat symbol (^). The configuration and operation of the embodiment of Figure 6 is essentially the same as Figure 4.

[0076] The charge olefin stream in line 54^ is delivered to the downstream catalyst bed 65, 67, 65', 67' in the interbed reactor 60+, 60'+ by the interbed olefin charge stream in line 58^. The downstream catalyst bed 65, 67, 65', 67' may contain a second stage oligomerization catalyst or may contain both a first stage oligomerization catalyst and a second stage oligomerization catalyst stacked or intermixed. The oligomerization olefin stream in line 64 or 64' containing dimers, oligomers, and unconverted feed olefins is collected in line 66^. The oligomerization recycle stream in line 52^ containing dimers, oligomers, and unconverted feed olefins removed from the oligomerization olefin stream in line 66^ is recycled to the oligomerization reactor 60+, 60'+ via the charge olefin stream in line 54^. An oligomerization product stream in line 84^, taken from and having the same composition as the oligomerized olefin stream in line 66^, may be fed to an olefin splitter column 90# for separation.

[0077] In the olefin separation column 90#, fractionation is performed to separate light and intermediate olefins in the net liquid overhead stream for recycling to the oligomerization unit 50^ in line 106^. Preferably, the olefin splitter column 90# may be operated to recycle unconverted ethylene back to the first stage oligomerization unit 50^ in line 106^ while sending the heavier oligomers to hydrogenation in the bottom line 94#. The net liquid light olefin splitter overhead stream in line 106^ may contain primarily ethylene which may be recycled to the oligomerization reactors 60+, 60'+ in the primary charge lines 57^, 57'^, respectively, after mixing with a paraffin-containing diluent stream, which may be the net stripped hydrogenation stream in line 148^.

[0078] The net liquid overhead stream in line 106^ may be heated if necessary and sent by primary charge lines 57^, 57'^ to upstream catalyst beds 61, 61' which may contain only first stage oligomerization catalyst or may contain both first and second stage oligomerization catalysts, stacked or intermixed. By recycling the ethylene dominated stream to the upstream end of reactor 60+, 60+', heat generation may be more easily controlled.

[0079] Starting from methanol, ethanol, and even carbon dioxide, the disclosed oligomerization process can produce jet and diesel fuels from ethylene or propylene that meet applicable fuel requirements. EXAMPLES

[0080] Example 1 We have prepared catalyst supports using the oil drop method described above, which combines synthesis and shaping in a single continuous step to achieve compositional homogeneity with aluminum oxide fully dispersed in the silica matrix. One catalyst, Example 2, was made using CCIC powder that contains a crystalline alumina phase dispersed throughout the support. Examples of catalyst supports prepared by a given method are shown in Table 1 below.

[0081] [Table 1]

[0082] Example 2 Metals were incorporated onto the supports in Table 1 using evaporative impregnation or ion exchange techniques to provide the catalysts in Table 2.

[0083] [Table 2]

[0084] Example 3 To demonstrate the utility of a given catalyst, the selective catalysts in Table 2 were tested for the conversion of ethylene to dimers and oligomers. In one unit, tests are run at 6.2 MPa(g) (900 psig) using chemical grade ethylene feed. In this test, normal butane, with a nominal ratio of n-butane to ethylene of 4 by weight, is used as a diluent to reduce the heat of reaction. Tests are run at catalyst loadings ranging from 15 to 40 grams and at a total feed rate of 5.0 hours. -1 The combined feed rates were run to achieve WHSV. The catalysts were pre-dried in flowing nitrogen at temperatures between 300 and 400°C until the effluent dew point stabilized at about -40 to -50°C. Normal butane and ethylene were pre-dried using 3A and 13X molecular sieves in series. The product effluent was passed through a high pressure separator and the resulting liquid stream was analyzed by on-line GC gas chromatography using PONA and alumina columns. The overhead gas was analyzed by alumina column for paraffin / olefin / naphthene / aromatic ratios in the light paraffin, light olefin, naphtha and distillate ranges. The test results for example catalysts 5.1 and 6.1 are shown in Tables 3 and 4 below.

[0085] [Table 3]

[0086] [Table 4]

[0087] When tested with n-butane co-feed as a thermally mediated diluent at substantial space velocities, example catalysts 5.1 and 6.1 exhibited stable ethylene conversions of over 90%, C8-C16 distillate yields of 30-50 wt% and 10-15 wt%, respectively, approaching 400 hours and 200 hours, respectively.

[0088] Example Catalyst 3.2 was prepared on the oil-dropped spherical support of Example 3, but with a size of 1 / 32 inch diameter. When Example Catalyst 3.2 was tested according to the prescribed test method, it achieved over 90% ethylene conversion with 20-40% C8-C16 yield, as shown in Table 5 below.

[0089] [Table 5]

[0090] Comparative Example Catalyst 1.1 was prepared on the oil-dropped spherical support of Comparative Example 1. When Comparative Example Catalyst 1.1 was tested according to the prescribed test method, the catalytic activity measured by ethylene conversion and C8-C16 yield was very low, especially under the condition of suitable diluent to ethylene ratio for thermal mediation as shown in Table 6 below. The catalyst contains a crystalline alumina phase by X-ray diffraction and a nickel aluminate spinel phase as shown by X-ray diffraction and confirmed by XANES (X-Ray Absorption Near Edge Structure).

[0091] [Table 6]

[0092] Example 4 In other tests, we evaluated various catalysts using 100% ethylene and a C2-C4 olefin feed blend containing 30-40 wt% ethylene, 40-55 wt% propylene, and 15 wt% butylene, simulating the olefin mixture coming out of an MTO unit, respectively. These tests also used n-butane as a diluent to mitigate the heat of reaction, and the operating pressure was maintained at 60 bar(g) (870 psig) to 62 bar(g) (900 psig). In one test, the C2-C4 olefin mixture was tested with the MTT catalyst, Example 4.1, and in another test, the ethylene feed was tested with a nickel on amorphous silica alumina (ASA) catalyst, Example 4.2. In the third test, Example 4.3, the nickel on amorphous silica alumina and MTT catalysts were tested in a stacked configuration using a given C2-C4 olefin blend. Nickel was incorporated onto ASA oil-dropped sphere support Example 3 of 75 SiO2 and 25% Al2O3 with a diameter of 1 / 16 inch using an evaporation impregnation technique to obtain Example Catalyst 3.1. MTT catalyst was prepared by extruding zeolite MTT of 45 Si / Al2 and boehmite alumina of 25 / 75 zeolite / alumina formation and calcining at 550°C for 2 hours in flowing air. In the former, nickel was incorporated onto amorphous silica-alumina sphere support of 75 mol% SiO2 and 25 mol% Al2O3 made using a given oil-dropped process. Normal butane was used as a diluent in all three tests. A total of 10 grams of catalyst was loaded in Test Examples 4.1 and 4.2 using a given C2-C4 olefin mixture and 100% ethylene, respectively. In test example 4.3, stacked loading of 5 grams of example catalyst 3.1 made of amorphous silica-alumina and 10 grams of zeolite catalyst made of 25% MTT and 75% alumina was carried out.The results are shown in Table 7 below.

[0093] [Table 7]

[0094] As shown in Table 7, the zeolite catalyst with 25% MTT and 75% alumina easily converted propylene and mixed butenes to levels above 90% and ethylene to moderate levels of 40-70% when operated at an average bed temperature of 190-230°C in Test Example 4.1. On the other hand, the nickel catalyst on amorphous silica-alumina oil droplet spheres, Example Catalyst 3.1, converted ethylene to nearly 100% in Test Example 4.2. Test Examples 4.1 and 4.2 appear to show that nickel on amorphous silica-alumina oil droplet spheres can achieve ethylene conversions of over 90% over a wide temperature range of 60-210°C in the presence of a hydrocarbon diluent. As shown in Test Example 4.3, when nickel on an amorphous silica-alumina support is combined with a zeolite catalyst, conversions of over 90% can be achieved for ethylene, propylene, and mixed butenes without increasing the overall contact time (see Table 7).

[0095] Example 5 SAPO-18 powder synthesis was carried out in a 2 liter autoclave by first adding the desired amount of a structure directing template such as diisopropylethylamine (DIPEA) to a pre-made solution of phosphoric acid in cold / ice water. The reaction mixture was stirred while the desired amount of silica and alumina precursors (Ludox colloidal silica and Versal 251 alumina, respectively) were added in sequential steps separated by 5 minute intervals. The reaction slurry was then transferred to a reactor, which was then sealed and heated from room temperature to 175°C at a rate of 15°C / hr and an agitation speed of 300 RPM. Once the set point temperature was reached, the reactor was stirred for 48 hours, at which point the reactor was cooled and the product powder was washed and isolated via centrifugation. The synthesized MTT was NH4+ ion exchanged, dried, bound with an alumina binder, and calcined. Alternatively, the synthesized MTT can be calcined prior to ion exchange and addition of binder. The typical yield, which is equal to the ratio of the mass of the calcined powder to the total mass of the synthesis gel, was 17%.

[0096] Specific embodiments While the following will be described in conjunction with specific embodiments, it will be understood that this description is illustrative, but is not intended to limit the scope of the foregoing description and appended claims.

[0097] A first embodiment of the present disclosure is a process for oligomerizing an olefin stream, comprising: oligomerizing the olefin stream with a first stage oligomerization catalyst to provide a first stage oligomerized olefin stream; oligomerizing the first stage oligomerized olefin stream with an oligomerization catalyst to provide a second stage oligomerized olefin stream; and recycling an ethylene stream to the first stage oligomerization process. An embodiment of the present disclosure is one, any, or all of the preceding embodiments to the first embodiment of this paragraph, in which the first stage oligomerization catalyst comprises a metal on amorphous silica-alumina. An embodiment of the present invention is one, any, or all of the preceding embodiments to the first embodiment of this paragraph, in which the first stage or second stage oligomerization catalyst comprises one or more of MRE, TON, MTT, MFS, MFI, MEL, AFO, AEL, EUO, and FER. An embodiment of the present disclosure is any one, any or all of the preceding embodiments through the first embodiment of this paragraph, wherein the first stage oligomerization catalyst or the second stage oligomerization catalyst is regenerated by contact with air. An embodiment of the present disclosure is any one, any or all of the preceding embodiments through the first embodiment of this paragraph, wherein the first stage oligomerization catalyst or the second stage oligomerization catalyst is regenerated by contact with air. An embodiment of the present disclosure is any one, any or all of the preceding embodiments through the first embodiment of this paragraph, wherein the first stage oligomerization catalyst or the second stage oligomerization catalyst is regenerated by contact with air. An embodiment of the present disclosure is any one, any or all of the preceding embodiments through the first embodiment of this paragraph, wherein the first stage oligomerization catalyst or the second stage oligomerization catalyst is regenerated by contact with air. An embodiment of the present disclosure is any one, any or all of the preceding embodiments of this paragraph through the first embodiment of this paragraph, further comprising separating an ethylene stream from the first stage oligomerized olefin stream and recycling the ethylene stream to the first stage oligomerization process. An embodiment of the present disclosure is any one, any or all of the preceding embodiments of this paragraph through the first embodiment of this paragraph, wherein the second stage oligomerization process is downstream of the first stage oligomerization process.An embodiment of the present disclosure is any one, any or all of the preceding embodiments of this paragraph through the first embodiment of this paragraph, further comprising separating an ethylene stream from the second stage oligomerized olefin stream and recycling the ethylene stream to the first oligomerization step. An embodiment of the present disclosure is any one, any or all of the preceding embodiments of this paragraph through the first embodiment of this paragraph, further comprising separating an intermediate olefin stream from the second stage oligomerized olefin stream and recycling the intermediate olefin stream to the second stage oligomerization step. An embodiment of the present disclosure is any one, any or all of the preceding embodiments of this paragraph through the first embodiment of this paragraph, further comprising separating a distillate olefin stream from the second stage oligomerized olefin stream and hydrogenating the distillate olefin stream. An embodiment of the present disclosure is any one, any or all of the preceding to first embodiments of this paragraph, further comprising separating the hydrogenated distillate stream into a jet stream and a diesel stream. An embodiment of the present disclosure is any one, any or all of the preceding to first embodiments of this paragraph, further comprising removing a diluent stream from the hydrogenated distillate stream and charging the diluent stream to a first stage oligomerization process. An embodiment of the present disclosure is any one, any or all of the preceding to first embodiments of this paragraph, further comprising removing a diluent stream from the hydrogenated distillate stream and charging the diluent stream to a first stage oligomerization process. An embodiment of the present disclosure is any one, any or all of the preceding to first embodiments of this paragraph, further comprising separating the hydrogenated distillate stream into a jet stream and a diesel stream.

[0098] A second embodiment of the present disclosure is a process for converting methanol to distillate fuels, comprising contacting an oxygenate stream with an MTO catalyst to produce an olefin stream, and oligomerizing the olefin stream with an oligomerization catalyst to produce an oligomerized olefin stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiment of this paragraph through the second embodiment of this paragraph, wherein the oligomerization step further comprises first oligomerizing the olefin stream with a first stage oligomerization catalyst to produce a first stage oligomerized olefin stream, and oligomerizing the first stage oligomerized olefin stream. An embodiment of the present disclosure is one, any, or all of the preceding embodiment of this paragraph through the second embodiment of this paragraph, further comprising separating an ethylene stream and an intermediate olefins stream from the oligomerized olefins stream, recycling the ethylene stream to the first stage oligomerization process, and recycling the intermediate olefins stream to the second stage oligomerization process.

[0099] 1. A third embodiment of the present disclosure is a process for converting ethylene to distillate fuels, comprising: oligomerizing an ethylene stream with a first stage oligomerization catalyst to produce a first stage oligomerized olefin stream; oligomerizing the first stage oligomerized olefin stream with a second stage oligomerization catalyst to produce a second stage oligomerized olefin stream; separating a distillate olefin stream from the second stage oligomerized olefin stream; and saturating the distillate olefin stream to provide a distillate fuel. An embodiment of the present disclosure is one, any, or all of the preceding to third embodiments of this paragraph, further comprising separating an ethylene stream and an intermediate olefin stream from the second stage oligomerized olefin stream; recycling the ethylene stream to the first stage oligomerization process; and recycling the intermediate olefin stream to the second stage oligomerization process. An embodiment of the present disclosure is any one, any or all of the preceding to third embodiments of this paragraph, where the ethylene stream is obtained by contacting an oxygenated stream with an MTO catalyst to produce an olefin stream. An embodiment of the present disclosure is any one, any or all of the preceding to third embodiments of this paragraph, where the olefin stream comprises at least 5 wt.% ethylene.

[0100] A fourth embodiment of the present disclosure is a process for converting ethylene, comprising oligomerizing an ethylene stream with an oligomerization catalyst comprising a Group VIII metal on a silica aluminum oxide support, the support comprising 70-99.5 wt. % SiO2. An embodiment of the present disclosure is any one, any or all of the preceding through fourth embodiments of this paragraph, where the support comprises 70-99.5 wt. % SiO2. An embodiment of the present disclosure is any one, any or all of the preceding through fourth embodiments of this paragraph, where the aluminum oxide is fully dispersed in the silica matrix. An embodiment of the present disclosure is any one, any or all of the preceding through fourth embodiments of this paragraph, where the aluminum oxide is non-crystalline. An embodiment of the present disclosure is any one, any or all of the preceding through fourth embodiments of this paragraph, where the support is synthesized and formed into a shape in a single step. The present disclosure is directed to any one, any or all of the preceding through fourth embodiments of this paragraph, wherein the support is an oil droplet sphere. The present disclosure is directed to any one, any or all of the preceding through fourth embodiments of this paragraph, wherein the support has a surface area of ​​300-600 m2 / g as determined by nitrogen BET. The present disclosure is directed to any one, any or all of the preceding through fourth embodiments of this paragraph, wherein the support has a porosity of 0.40-1.4 mL / g as measured by total mercury intrusion volume. The present disclosure is directed to any one, any or all of the preceding through fourth embodiments of this paragraph, wherein the Group VIII metal is nickel. The present disclosure is directed to any one, any or all of the preceding through fourth embodiments of this paragraph, wherein the molar ratio of Group VIII metal to Al2 in the catalyst is 0.04-0.70. An embodiment of the present disclosure is any one, any or all of the preceding embodiment of this paragraph through the fourth embodiment of this paragraph, wherein the catalyst also includes a Group IA metal, and the molar ratio of the Group IA metal to aluminum is from 0.04 to 0.4.

[0101] Without further elaboration, it is believed that, using the preceding description, one skilled in the art can utilize the present disclosure to its fullest extent and easily ascertain the essential characteristics of the present disclosure without departing from the spirit and scope of the present disclosure, and can make various changes and modifications to the present disclosure to adapt it to various uses and conditions. Therefore, the preceding preferred specific embodiments are to be construed as merely illustrative and not limitative of the remainder of the disclosure in any way, and are intended to cover various modifications and equivalent arrangements falling within the scope of the appended claims.

[0102] Above, all temperatures are set forth in degrees Celsius and all parts and percentages are by weight unless otherwise indicated.

Claims

1. 1. A process for oligomerizing an olefin stream comprising: oligomerizing said olefin stream with a first stage oligomerization catalyst to provide a first stage oligomerized olefin stream; oligomerizing said first-stage oligomerized olefin stream with an oligomerization catalyst to provide a second-stage oligomerized olefin stream; and recycling said ethylene stream to said first stage oligomerization step.

2. 10. The process of claim 1, wherein the first stage or the second stage oligomerization catalyst comprises a metal on amorphous silica-alumina.

3. 2. The process of claim 1, wherein the first stage or the second stage oligomerization catalyst comprises one or more of MRE, TON, MTT, MFS, MFI, MEL, AFO, AEL, EUO, and FER.

4. 10. The process of claim 1, wherein the first stage oligomerization catalyst or the second stage oligomerization catalyst is regenerated by contact with air.

5. 10. The process of claim 1 further comprising recycling oligomerized olefins from said second stage oligomerization stream to said second stage oligomerization step.

6. 10. The process of claim 1 further comprising recycling said first-stage oligomerized olefin to said first-stage oligomerization step.

7. 10. The process of claim 1 further comprising separating an ethylene stream from said first stage oligomerized olefin stream and recycling said ethylene stream to said first stage oligomerization step.

8. 10. The process of claim 1 further comprising separating an ethylene stream from said second stage oligomerized olefin stream and recycling said ethylene stream to the first oligomerization step.

9. 10. The process of claim 1 further comprising separating an intermediate olefin stream from said second-stage oligomerized olefin stream and recycling said intermediate olefin stream to said second-stage oligomerization step.

10. 1. A process for converting ethylene, comprising oligomerizing an ethylene stream with an oligomerization catalyst comprising a Group VIII metal on an amorphous silica aluminum oxide support, said support comprising 70-99.5 wt. % SiO2.

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