Process for converting olefins into 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] (Priority statement) This application claims priority to U.S. Provisional Application No. 63 / 325,102, filed on 29 March 2022, which is incorporated herein by reference in its entirety.

[0002] (Field of Invention) This field concerns the conversion of oxygenates to distillates. In particular, this field may relate to the conversion of oxygenates to olefins and the oligomerization of olefins into distillate fuels. [Background technology]

[0003] Carbon dioxide is a so-called greenhouse gas, and its concentration in the atmosphere is desired by many. Carbon dioxide can be converted to oxygenated substances such as methanol or dimethyl ether. Molecular sieves, such as microporous crystalline zeolites and non-zeolite catalysts, particularly silicoaluminophosphate (SAPO), are known to promote the conversion of oxygenated substances into hydrocarbon mixtures, especially hydrocarbon mixtures consisting mostly of light olefins. The highly efficient methanol to olefin (MTO) process can convert oxygenated substances into light olefins, which have typically been considered for plastic production. The light olefins produced from the MTO process are highly enriched in ethylene.

[0004] Oligomerization of light olefins is a process that can carry out 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 may have very poor diesel quality.

[0005] Jet fuel is one of the few petroleum fuels and cannot be easily replaced by electric motor systems. This is because fueling an aircraft requires high energy output, which electric motors cannot provide. In certain regions, substantial incentives are currently available for green jet fuel.

[0006] An efficient process is needed to convert oxygenated materials into distillate fuels. [Overview of the Initiative]

[0007] The inventors have developed a process for converting methanol into a distillate fuel, comprising contacting an oxygenated flow with an MTO catalyst to generate an olefin flow, and oligomerizing the olefin flow with an oligomerizing catalyst to generate an oligomerized olefin flow. The olefin flow may be first oligomerized to provide a first-stage oligomerized olefin, which may be further oligomerized to provide a second-stage oligomerized olefin flow. The second-stage oligomerized olefin flow may be separated into a jet fuel flow and a diesel fuel flow. The olefin flow may be obtained by converting oxygenated materials to olefins with an MTO catalyst. The ethylene flow may, in embodiments, be recycled back to the first-stage oligomerization step. The oligomerizing catalyst may have a silica aluminum oxide support in which aluminum oxide is completely dispersed throughout the support. [Brief explanation of the drawing]

[0008] [Figure 1] This is a schematic diagram of the process and apparatus of this disclosure. [Figure 2] This is a schematic diagram of an alternative process and apparatus to the present disclosure. [Figure 3] This is a schematic diagram of additional alternative processes and apparatus as described herein. [Figure 4] This is a schematic diagram of a further alternative embodiment shown in Figure 1. [Figure 5] This is a schematic diagram of an alternative embodiment shown in Figure 4. [Figure 6]Schematic diagram of a further alternative embodiment of FIG. 4. **DETAILED DESCRIPTION OF THE INVENTION**

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

[0010] The term "downstream communication" means that at least a portion of the fluid flowing to the object in downstream communication can flow operably from the object in fluid communication.

[0011] The term "upstream communication" means that at least a portion of the fluid flowing from the object in upstream communication can flow operably into the object in fluid communication.

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

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

[0014] The term "bypass" means that the object is removed from downstream communication with the bypass target at least to the extent of bypassing.

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

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

[0017] This 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 oxygenation flow 10 into an MTO reactor 12 and contacting the oxygenation flow with an MTO catalyst under MTO reaction conditions to convert the oxygenated product into an olefin and water. The MTO reactor 12 may provide a fluid catalyst operating under high-speed fluid conditions. The oxygenated product may be methanol, dimethyl ether, ethanol, or a combination thereof. Methanol may be derived from a water-gas shift reaction of carbon dioxide and 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. Patents 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 known for producing lighter olefins. The increased production of higher olefins with SAPO-18 reduces the amount of ethylene processed 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) compared to SAPO-34, which typically operates at lower pressures.

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

[0020] The MTO catalyst is separated from the olefin stream after the MTO reaction, and the hydrocarbons are stripped off with an inert gas such as nitrogen or water vapor. It is then transferred to a regenerator 14 in line 16, where air comes into contact with the spent catalyst to burn the coke from the MTO catalyst.

[0021] The MTO reactor 12 generates a product olefin stream in line 20. The olefin stream in line 20 can be sent to a dehydration column 22 for cooling, separating water from the olefin stream, neutralizing acidic compounds, and generating a dehydrated olefin stream in line 24 with reduced water content, and a water stream in line 26 through which water passes for further processing. The dehydration column 22 may consist of two columns. The water stream in line 26 can be further processed to separate further oxygenated materials that can be recycled back to the MTO reactor 12. The dehydrated olefin stream in line 24 can 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 can be recycled back to the MTO reactor 10, where it can be converted to olefins on the MTO catalyst. The deoxygenated olefin stream in line 32 may contain a large amount of C2-C6 olefins.

[0022] The deoxygenated olefin stream in line 32 is oligomerized with an oligomerization catalyst to produce an oligomerized olefin stream containing C9+ olefins. If the olefin stream contains substantial ethylene, in the embodiment, the olefin stream may first be contacted with a first-stage oligomerization catalyst to oligomerize the ethylene to butene and higher olefins, and then contacted with a second-stage oligomerization catalyst to oligomerize the first-stage oligomerized olefin and the unconverted feed olefin to C9+ olefins. The contact order may be reversed.

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

[0024] Any suitable catalyst capable of selectively hydrogenating diolefins in a naphtha stream can be used. Suitable catalysts include, but are not limited to, those comprising copper and at least one other metal such as titanium, vanadium, chromium, manganese, cobalt, nickel, palladium, zinc, molybdenum, and cadmium, or mixtures thereof. The metals are preferably supported on inorganic oxide supports such as silica and alumina. The selectively hydrogenated olefin stream may exit the selective hydrogenation reactor in line 42 and enter the hydrogenation separator 44, providing a hydrogen-rich overhead stream in line 46, which may be compressed and possibly supplemented with a supplement hydrogen stream in line 47 and returned as recirculated hydrogen in line 36. The selectively hydrogenated liquid feed olefin stream from the bottom of the separator 44 may be transferred to the oligomerization unit 50 in line 48. The feed olefin stream in line 48 may contain at least 5% by weight, preferably at least 10% by weight, ethylene, preferably at least 20% by weight, and even more preferably at least 25% by weight. The feed olefin stream in line 48 may contain ethylene as a main component. The feed olefin stream may also contain at least 5% by weight, typically at least 10% by weight, preferably at least 20% by weight, and even more preferably at least 25% by weight, one or more of C3, C4, C5, and C6 olefins. The feed olefin stream in line 48 may contain at least 5% by weight, typically at least 10% by weight, preferably at least 20% by weight, and even more preferably at least 25% by weight, propylene. The feed olefin stream in line 48 may contain propylene as a main component.

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

[0026] For example, the feed olefin stream may be provided from an ethanol dehydration process. 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 20% to 100%. Ethanol or ethanol-containing feed material may be optionally supplied to a dehydration reactor 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. Ethanol is 0.1 -1 ~30 o'clock -1 It can be introduced into the dehydration reactor at WHSV. In some embodiments, ethanol is 0.5 -1 ~5 o'clock -1 The WHSV can be supplied to the dehydration reactor. The dehydration reactor can operate at temperatures of 200°C to 500°C. In some embodiments, the dehydration reactor can operate at temperatures of 300°C to 450°C. In some embodiments, the dehydration reactor can operate at pressures of 0 barg to 83 barg. In some embodiments, the dehydration reactor can operate at pressures of 0 barg to 35 barg. The ethanol conversion rate can vary between 10% and 100% depending on the operating conditions and the selected catalyst. The ethylene-containing product can be purified to remove water, by-products, oxygen, and other impurities. Purification may include condensing water and purifying the product through purifying adsorbents such as silica, molecular sieves, and carbon. The purified ethanol can 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 for transferring unconverted feed olefins, a diluent flow containing paraffin which may be a net stripped hydrogenation flow in line 148, and a first-stage oligomerization recirculation flow in line 52 containing the first-stage oligomerized olefins may be added to the feed olefin flow in line 48 to provide the charge olefin flow in line 54. The charge olefin flow in line 54 may supply the primary charge olefin flow in line 56 and the inter-floor olefin charge flow in line 58. The primary charge olefin flow in line 56 may be heated and charged into the first-stage oligomerization reactor 60.

[0028] The diluent stream may include a paraffinic stream that absorbs the heat generated by the oligomerization reaction. The diluent stream may be provided in a diluent-to-feed ratio of 1:1 to 6:1, preferably 2:1 to 5:1. The diluent stream is preferably a C9+ paraffinic stream, which may be taken from a stripper bottom stream in line 146 downstream of the hydrogenation. The diluent stream may also be a light paraffinic stream.

[0029] The main reactor preferably contains two stationary catalyst beds, where most ethylene and some propylene and higher olefins are converted. The primary charge olefin stream is preferably charged into the first catalyst bed in line 57 in a downward flow operation. However, an upward flow operation may be preferable. When the conversion of ethylene occurs in the first catalyst bed, exothermic reactions occur. Therefore, a suitable heat exchange medium is selected to control the heat of reaction. An inter-bed olefin charge stream is charged into the inter-bed position of reactor 60 by line 58 to cool the first-stage oligomerized effluent from the first bed. In embodiments, the effluent from the first catalyst bed may be combined with the inter-bed 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 having 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 by a Group VIB metal. Typically, since silica and alumina are present only in the base, the ratio of silica to alumina is the same for both the catalyst and the base. The metal can be impregnated onto the silica-alumina base or ion-exchanged within the silica-alumina base. Co-mulling is also considered. In addition, preferred catalysts have a concentration of 300-600 m, determined by the nitrogen BET method. 2 It has a surface area of ​​ / g.

[0031] The most preferred first-stage oligomerization catalyst is described below. The preferred ethylene conversion catalyst comprises an amorphous silica-alumina support. The composition of the amorphous silica-alumina is in the range of 40 to 99.5% by weight of SiO2, preferably at least 70% by weight, preferably 75% to 97.5% by weight of SiO2, with the remainder being alumina. More preferably, the composition of the amorphous silica-alumina is in the range of 87 to 98% by weight of SiO2, with the remainder being alumina.

[0032] Amorphous silica-alumina supports for catalysts are most preferably prepared using an oil-dropping process, and the synthesis of amorphous silica-alumina and the support for catalyst morphogenesis are achieved in a single continuous manufacturing process described in U.S. Patents No. 3,909,450, No. 4,629,717, and No. 5,139,989. In the oil-dropping process, amorphous silica-alumina was synthesized online by vigorously mixing silica and alumina precursors to ensure that the silica and alumina were atomically interdispersed. Appropriate levels of neutralizing agents, such as hexamethylenetetramine (HMT), urea, ammonia, or a combination thereof, were introduced at ambient temperature during the aforementioned mixing step. The resulting premixture was dropped into a hot oil bath via a vibrating dropper, and the premixture droplets were converted from a sol state to a gel state via activation of a gelling agent in the hot oil bath. The formed spherical droplets are aged to establish a porous structure, followed by washing with water under controlled pH to remove residual oil and alkali and alkaline earth metals associated with the silica and alumina precursors entering the synthesis. Alternatively, the water washing may be carried out without removing all alkali or alkaline earth cations associated with the silica and alumina precursors, thereby providing ion exchange sites for subsequent metal incorporation via an ion exchange procedure. By selecting and controlling the gelling agent level, solid content, and aging conditions including pressure, temperature, and time, the carrier properties can be adjusted to obtain the attributes required for the catalytic process.

[0033] For example, when amorphous silica-alumina is used as a support for dispersing a metal for ethylene conversion, the alumina must be completely dispersed in the silica. This is achieved by an oil-droplet sol composition having more than 70-75% by weight of SiO2. If the alumina is not completely dispersed in the silica matrix, the alumina phase interacts with the metal to form spinels, resulting in, for example, a loss of ethylene conversion reactivity. Amorphous silica-alumina synthesized according to a predetermined process combines synthesis and morphogenesis in a single continuous step. Thus, homogeneity of composition is achieved with aluminum oxide completely dispersed in the silica matrix, as indicated by the X-ray diffraction pattern shown by the absence of a crystalline alumina phase. In other words, the composition of the support is uniform on the surface of the support throughout the entire core of the support.

[0034] The size of the spheres may be in the range of 1 / 32 to 1 / 10 inch in diameter, preferably 1 / 20 to 1 / 12 inch in diameter. The porosity of the carrier, as measured by the total intrusion volume by mercury intrusion technique, is in the range of 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 50 to 80%.

[0035] Group VIII elements, and optionally Group VI, IA, and IIA elements, can be incorporated onto a carrier 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 preferred first-step oligomerization catalyst comprises an amorphous silica-alumina support. One of the components of this first-step oligomerization catalyst support used in this disclosure is alumina. The alumina may be any of various hydrated aluminum oxide or alumina gels, such as alpha-alumina monohydrate with a boehmite or pseudoboehmite structure, alpha-alumina trihydrate with a gibbsite structure, or beta-alumina trihydrate with a bayerite structure. A particularly preferred alumina is available from Sasol North America Alumina Product Group under the trademark "Catapal." This material is an extremely high-purity alpha-alumina monohydrate (pseudoboehmite) and has been shown to yield high-purity gamma-alumina after calcination at high temperatures. Another component of the catalyst support is amorphous silica-alumina. A preferred composition of amorphous silica-alumina is in the range of 70–99.5% by weight of SiO2, with the remainder being alumina. More preferably, the composition of amorphous silica-alumina is in the range of 87-98% by weight of SiO2, with the remainder being alumina. A suitable silica-alumina having a silica-to-alumina ratio of 4.0 is available, for example, from CCIC, a subsidiary of JGC in Japan. Suitable silica-alumina may have a silica-to-alumina molar ratio of 4.0-300 and can be synthesized using a batch or continuous process with a co-gel or sequential procedure using balanced cations and anions, followed by aging, spray drying, and water washing. Appropriate aging at a pH of 6-8 is preferred to obtain a pore texture with desirable mass transfer properties. This catalyst support has a crystalline alumina phase throughout the support.

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

[0038] A suitable silica-alumina mixture is prepared by mixing proportional volumes of silica-alumina and alumina to achieve a desired silica-to-alumina ratio. In one embodiment, 75-95% by weight of amorphous silica-alumina and 10-20% by weight of alumina powder, having a silica-to-alumina ratio of 4.0-300, provide a suitable carrier. In another embodiment, other ratios of amorphous silica-alumina to alumina may be preferred.

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

[0040] The alumina in the binder can be gelatinized by adding a monobasic acid such as nitric acid or formic acid to the mixture in an aqueous solution. Additional water may be added to the mixture to provide sufficient wettability to form a dough with sufficient viscosity for extrusion or spray drying.

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

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

[0043] Typical properties of the amorphous silica-alumina support used herein are a total pore volume, an average pore diameter, and a surface area large enough to provide substantial space and area for depositing the active metal components. The total pore volume of the support measured by the conventional mercury porosimeter method is usually 0.2 to 2.0 cc / g, preferably 0.25 to 1.0 cc / g, and most preferably 0.3 to 0.9 cc / g. The surface area measured by the B.E.T. method is typically greater than 50 m 2 / g, for example greater than 200 m 2 / g, preferably at least 250 m 2 / g, and most preferably 300 m 2 / g to 550 m 2 / g.

[0044] The most preferred first-stage oligomerization catalyst is calcined amorphous refractory oxide support particles in the form of oil-droplet spheres, containing a metal from Group VIII at a molar concentration of 0.04 to 0.70 M8 / Al2 (M8 = Group VIII). The Group VIII element is preferably nickel. The catalyst preferably contains a Group IA element, and optionally a Group IIA element, and M IA / Al(M IA The concentration of the group IA metals 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% by weight.

[0045] The incorporation of metals can be achieved by any method known in the art, for example, by ion exchange, evaporation impregnation, or pore filling / spray impregnation. Ion exchange of 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 first carried out using an alkaline nitrate, followed by ion exchange of nickel nitrate. Ion exchange can be carried out using a solution containing both nickel and an alkaline nitrate. Impregnation can be carried out sequentially or simultaneously, preferably using a solution-to-support ratio greater than 1.0 by volume to ensure that the ion exchange process takes place during the impregnation operation.

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

[0047] The first-stage oligomerization catalyst can be regenerated when deactivated. Preferred regeneration conditions include, for example, exposing the catalyst to hot air at 450-550°C for 3 hours in situ. To facilitate regeneration without downtime, a swing-floor configuration is used with an alternative first-stage oligomerization reactor 60'. If the first-stage oligomerization catalyst in the first-stage oligomerization reactor 60 is deactivated, valves on the primary charge line 57 and the inter-floor charge line 58 to the first-stage oligomerization reactor 60 are closed, and valves on the alternative primary charge line 57' and the alternative inter-floor charge line 58' are opened to charge the alternative first-stage oligomerization reactor 60' with an olefin flow. The regeneration gas flow from line 62 is then introduced into the first-stage oligomerization reactor 60 that requires regeneration. The regeneration gas may contain air with increased or decreased oxygen concentration. An alternative first-stage oligomerization reactor 60' can be regenerated in the reverse manner by closing valves on lines 57' and 58' and introducing regenerative gas from line 62'. Each first-stage oligomerization reactor 60, 60' may include vent lines 63, 63' for discharging regenerative combustion exhaust gas, respectively. The activity and selectivity of the regenerated catalyst are comparable to that of the fresh catalyst.

[0048] The first-stage oligomerized olefin stream in line 66 is collected from lines 64 and 64' with an increased butene concentration compared to the charging olefin stream in line 54, and is split into a first-stage oligomerized recirculation stream in line 52 and a charging second-stage oligomerized olefin stream in line 68. The intermediate olefin stream in net light olefin splitter bottom line 108 containing C3-C8 olefins and the oligomerized recirculation stream in line 72 containing oligomerized olefins can be added to the charging second-stage oligomerized olefin stream in line 68 to provide a charging second-stage oligomerized olefin stream in line 74. The charging second-stage oligomerized olefin stream in line 74 may be cooled and charged into the second-stage oligomerization reactor 70 in line 75. The second-stage oligomerization reactor 70 may communicate downstream with the first-stage oligomerization reactor 60. The second-stage oligomerization reactor 70 preferably operates in a downward flow manner. However, an upward flow manner may be preferable. The charged second-stage oligomerized olefin stream is brought into contact with the second-stage oligomerization catalyst to oligomerize C2-C8 olefins and provide olefins in the distillate range. The second-stage oligomerization stream, having an increased average carbon number beyond the charged second-stage oligomerized olefin stream in line 74, exits the second-stage oligomerization reactor 70 in line 76.

[0049] The oligomerization catalyst in the second stage may include a zeolite catalyst. The zeolite may constitute 5 to 95% by weight, for example, 5 to 85% by weight, of the catalyst. Suitable zeolites include those having structures from one of the following classes: MFI, MEL, ITH, IMF, TUN, FER, BEA, FAU, BPH, MEI, MSE, MWW, UZM-8, UZM-8HS, UZM-37, MOR, OFF, MTW, MRE, MFS, TON, MTT, AFO, ATO, and AEL. The three-letter codes indicating the zeotype are as defined by the Structure Commission of the International Zeolite Association and are managed at http: / / www.iza-structure.org / databases. UZM-8 is as described in U.S. Patent No. 6,756,030. In a preferred embodiment, the oligomerization catalyst may include a zeolite having a skeleton with a 10-membered ring pore structure. Examples of suitable zeolites having a 10-membered ring pore structure include TON, MTT, MFS, MRE, MFI, MEL, AFO, AEL, EUO, and FER. In a further preferred embodiment, the second-step oligomerization catalyst, which includes a zeolite having a 10-membered ring pore structure, may include a one-dimensional pore structure. The one-dimensional pore structure represents a zeolite containing non-crossing pores substantially parallel to one of the crystal axes. The pores preferably extend through the zeolite crystal. A preferred example of a zeolite having a 10-membered ring one-dimensional pore structure is MTT. In a further embodiment, the oligomerization catalyst includes an MTT zeolite. A preferred silica-to-alumina ratio for the MTT zeolite is 30-100.

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

[0051] One of the components of the catalyst binder used in this disclosure is alumina. The alumina source may be any of various hydrated aluminum oxides or alumina gels, such as alpha-alumina monohydrate with a boehmite or pseudo-boehmite structure, alpha-alumina trihydrate with a gibbsite structure, or beta-alumina trihydrate with a bayerite structure. Preferred alumina is available from UOP LLC under the trademark "VERSAL". Preferred alumina is available from Sasol North America Alumina Product Group under the trademark "Catapal". This material is extremely high-purity alpha-alumina monohydrate (pseudo-boehmite) and has been shown to yield high-purity gamma-alumina after calcination at high temperatures.

[0052] A suitable second-stage oligomerization catalyst is prepared by mixing proportionally volume amounts of zeolite and alumina to achieve a desired zeolite-alumina ratio. In embodiments, the zeolite content may be 5 to 90, for example, 10 to 85% by weight, preferably 25 to 75% by weight, with the remaining alumina powder providing a suitably supported catalyst. Silica supports are also considered. In an exemplary embodiment, an MTT-type zeolite catalyst, arranged on a high-purity pseudoboehmite alumina substrate in a ratio of 10 / 90 to 90 / 10, preferably 25 / 75 to 75 / 25, is provided in an oligomerization reactor 70.

[0053] A monobasic acid such as nitric acid or formic acid may be added to the mixture in an aqueous solution to dissolve the alumina in the binder. Additional water may be added to the mixture to provide sufficient wettability to form a dough with sufficient viscosity for extrusion or spray drying. Extrusion aids such as cellulose ether powder may also be added. 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 molded particles, a preferred method of which is to extrude the dough through a die having an opening of the desired size and shape, and then divide the extruded material into extruders of the desired length and dry them. Further calcination steps may be used to provide additional strength to the extruders. Generally, calcination is carried out in an airflow at temperatures of 260°C (500°F) to 815°C (1500°F). The MTT catalyst does not have the selectivity 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 they have a symmetrical cross-sectional shape, preferably spherical, cylindrical, or multi-lobed. The cross-sectional diameter of the particles may be as small as about 40 μm. However, it is preferably 0.79 mm (1 / 32 inch) to 6.35 mm (0.25 inch), and most preferably 1.06 mm (1 / 24 inch) to 4.23 mm (1 / 6 inch).

[0056] With respect to the oligomerization reactor 70, process conditions are selected to produce a higher proportion of jet-range olefins, which, when hydrogenated in subsequent steps as described below, yield 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 supplied to the oligomerization reactor 70. The first-stage oligomerized olefin stream charged in line 74 is cooled and supplied to the oligomerization reactor 70. To obtain the most desirable olefin product, the 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 of 21 barg (300 psig) to 69 barg (1000 psig), more preferably 49 barg (710 psig) to 63 barg (900 psig).

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

[0058] The second-stage oligomerization catalyst can be regenerated when deactivated. Preferred regeneration conditions include, for example, exposing the oligomerization catalyst to hot air at 500°C for 3 hours in situ. To facilitate regeneration without downtime, a swing-bed configuration is used with an alternative second-stage oligomerization reactor 70'. When the second-stage oligomerization reactor 70 is 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 into the alternative second-stage oligomerization reactor 70'. The regeneration gas stream from line 78 is then introduced into the second-stage oligomerization reactor 70 that requires regeneration. The regeneration gas may contain air with increased or decreased oxygen concentration. An alternative second-stage oligomerization reactor 70' can be regenerated in the reverse manner by closing a valve on line 75' and introducing regenerative gas from line 78'. Each second-stage oligomerization reactor 70, 70' may include vent lines 73, 73' for discharging regenerative combustion exhaust gas, respectively. The activity and selectivity of the regenerated catalyst are comparable to that of the fresh catalyst.

[0059] Compared to the second-stage oligomerized olefin stream charged into line 74, the second-stage oligomerized olefin stream from line 80, collected from lines 76 and 76' with increased C9+ olefin concentration, is split into the second-stage oligomerization recirculation stream from line 72 and the second-stage oligomerization product stream from line 84.

[0060] The second-stage oligomerization product stream from the oligomerization unit 50 is carried to the olefin recovery section 88, where it is fed into the heavy olefin splitter column 90. In the heavy olefin splitter column 90, hydrocarbons in the jet range, oligomers boiling lower than C8 hydrocarbons with atmospheric pressure boiling points below 150°C, are separated from the net heavy olefin splitter bottom stream in line 94, which contains C9+ hydrocarbons in the distillate range, typically C9-C20 olefins, into the net heavy olefin splitter overhead stream in line 92. The heavy olefin splitter column 90 can operate at bottom temperatures of 300°C to 500°C and overhead pressures of 9 barg to 15 barg. The heavy olefin splitter overhead stream is condensed and can be taken out as a net vapor overhead stream from the heavy olefin splitter receiver 96 in line 92.

[0061] The net vapor heavy olefin splitter overhead flow of C8- in line 92 can be fed into the light olefin splitter column 100. The light olefin splitter overhead flow can be cooled and separated in the light olefin splitter overhead receiver 102 into a net vapor overhead flow in line 104 containing methane and light off-gas, and a net liquid light olefin splitter overhead flow in line 106 containing ethylene. The net liquid light olefin splitter overhead flow in line 106, along with the diluent flow in line 148 and the first-stage oligomerization recirculation flow in line 52, can recirculate unconverted ethylene into the feed olefin flow in line 48, providing the charged olefin flow in line 54 to the first-stage oligomerization reactor 60. The net liquid light olefin splitter overhead stream may primarily contain ethylene, which can be recycled to the first-stage oligomerization reactor 60. In an alternative embodiment, the net liquid olefin splitter overhead stream may primarily contain propylene, which can be recycled to the first-stage oligomerization reactor 60. In this case, the ethylene is still incorporated into the net liquid light olefin splitter overhead stream in line 106. The intermediate olefin stream in the net light olefin splitter bottom line 108, containing C3-C8 or C4-C8 olefins, can be recycled and combined with the charged second-stage oligomerized olefin stream in line 68, along with the oligomerization recirculation stream in line 72, to provide the charged second-stage oligomerized olefin stream in line 74, which is oligomerized in the second-stage oligomerization reactor 70. Drag streams can be taken out from lines 106 and 108. The light olefin splitter column 100 can operate at a bottom temperature of 100°C to 300°C and an overhead pressure of 5 barg to 11 barg.

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

[0063] In exemplary embodiments, hydrogenation is carried out in a hydrogenation reactor 120 containing an alumina-supported platinum catalyst, for example, 0.5% to 0.9% by weight of alumina-supported platinum catalyst. Using this catalyst, hydrogenation preferably occurs at a temperature of 125 to 175°C and a pressure of 35 barg (500 psig) to 105 barg (1500 psig). According to these process conditions, the hydrogenation reactor 120 converts the olefin into a paraffin product having the same carbon number distribution as the olefin, thereby forming paraffins with a distillate range suitable for use as jet and diesel fuels.

[0064] The hydrogenation distillate flow discharged from the hydrogenation reactor 120 in line 124 can be cooled and supplied to the hydrogenation separator 130. In the hydrogenation separator 130, the hydrogenation distillate flow is separated into a hydrogenation separator vapor flow in overhead line 132 and a hydrogenation separator liquid flow in bottom line 134. The hydrogenation separator vapor flow in line 132 can be compressed and combined with supplemental hydrogen in line 136 to supply the hydrogen flow in line 122 and / or the hydrogen flow in line 36 to the selective hydrogenation reactor 40. The hydrogenation separator liquid flow in bottom line 134 can be heated by heat exchange with the hydrogenation distillate flow in line 124 and supplied to the stripper column 140.

[0065] The stripper column 140 strips the light gas from the hydrogenation separator liquid flow to provide a stripper off-gas flow in the off-gas line 142 from the stripper overhead receiver 144. The net stripped hydrogenation flow in the stripper bottom line 146 is split into a product fuel flow in line 147 and a diluent flow in line 148. The diluent flow in line 148 can be recirculated into the feed olefin flow in line 48 along with the light olefin splitter overhead flow 106 and the first-stage oligomerization recirculation flow in line 52 to provide the charge olefin flow in line 54. The diluent flow is inert in the oligomerization reactors and helps absorb the heat generated in the first-stage oligomerization reactors 60, 60' and the second-stage oligomerization reactors 70, 70'. Stripper column 140 can operate at bottom temperatures of 250°C to 500°C and overhead pressures of 2 barg to 8 barg.

[0066] The product fuel flow in line 147 is fed into the jet fractionation column 150, where it can be separated into an off-gas flow in the overhead line 152 from the jet receptor overhead 154, a green jet fuel flow from the jet receptor bottom line 156, and a green diesel flow in the net diesel bottom line 158. Both the jet fuel flow in line 156 and the diesel flow in line 158 can be supplied to their respective fuel pools. The jet fractionation column 150 can operate at a bottom temperature of 350°C to 600°C and an overhead pressure of 1 barg to 5 barg.

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

[0068] In Figure 2, the olefin splitter column 90# separates dimers and oligomers of C8 hydrocarbons, typically with atmospheric pressure boiling points below 150°C, which boil lower than jet-range hydrocarbons, from the net olefin splitter bottom flow in line 94#, which contains distillate-range C9+ hydrocarbons, typically C9-C20 olefins, in the net olefin splitter overhead flow in line 92#. The olefin splitter column 90# can operate at bottom temperatures of 100°C to 500°C and overhead pressures of 5 barg to 15 barg. The olefin splitter overhead flow is condensed and can be extracted as a net liquid overhead flow from the olefin splitter receptor 96 in line 106#. The net liquid olefin splitter overhead flow in line 106#, along with the diluent flow in line 148 and the first-stage oligomerization recirculation flow in line 52, can recirculate unconverted ethylene and oligomers into the feed olefin flow in line 48, providing the charge olefin flow in line 54 to the first-stage oligomerization reactor 60. The net liquid light olefin splitter overhead flow in line 106# may contain at least 10% by weight, preferably at least 20% by weight, and more preferably at least 25% by weight of ethylene, which can be recirculated to the first-stage oligomerization reactor 60. In the embodiment of Figure 2, the intermediate olefin flow in net olefin splitter bottom line 108, containing C3-C8 olefins, is not returned to the oligomerization reactor 70 in the oligomerized olefin flow in line 74. In the embodiment shown in Figure 2, the light olefins and intermediate olefins in the net liquid overhead flow can be recycled to the oligomerization unit 50 in line 106#. Alternatively, the olefin splitter column 90# may operate to recycle the light olefins, C2-C3 olefins, or C2-C4 olefins back to the oligomerization unit 50 in line 106#, while sending the heavier C4+ or C5+ oligomers to hydrogenation in the bottom line 94#.

[0069] In the alternative embodiment shown in Figure 3, even when only a single olefin splitter column 90# is used in the olefin recovery section 88#, the intermediate olefin stream is returned to the second-stage oligomerization reactor 70. Elements in Figure 3 having the same configuration as in Figure 2 will have the same reference numerals as in Figure 2. Elements in Figure 3 having a different configuration from the corresponding elements in Figure 2 will have the same reference numerals, but will be marked with an asterisk ( * ) is specified. The configuration and operation of the embodiment in Figure 3 are essentially the same as those in Figure 2, with a notable exception.

[0070] Olefin Recovery Section 88 * Olefin splitter column 90 * This is line 108, which is taken from the side of the column. * Provides internal side-cut flow. Olefin splitter side line 108 containing C3-C8 olefins. * The intermediate olefin flow can be recirculated to merge the first-stage oligomerized olefin flow charged in line 68 with the oligomerized recirculation flow in line 72, providing the second-stage oligomerized olefin flow charged in line 74, which is oligomerized in the second-stage oligomerization reactor 70. In addition, the olefin splitter overhead flow is condensed into line 106 * The net liquid overhead flow can be extracted from the internal olefin splitter receptor 96. Line 106 * The net liquid olefin splitter overhead flow within, along with the diluent flow in line 148 and the first-stage oligomerization recirculation flow in line 52, can recirculate unconverted ethylene into the feed olefin flow in line 48, providing the charge olefin flow in line 54 to the first-stage oligomerization reactor 60. Line 106 * The net liquid light olefin splitter overhead flow inside may contain ethylene as the main component, which can be recycled to the first-stage oligomerization reactor 60.

[0071] Figure 4 shows an embodiment in which the oligomerization reactor 60+ includes beds for both the first-stage oligomerization catalyst and the second-stage oligomerization catalyst. Elements in Figure 4 having the same configuration as in Figure 2 will have the same reference numerals as in Figure 2. Elements in Figure 4 having a different configuration from the corresponding elements in Figure 2 will have the same reference numerals but will be designated with a plus sign (+). The configuration and operation of the embodiment in Figure 4 are essentially the same as those in Figure 2.

[0072] The oligomerization reactor 60+ includes a catalyst bed in which both the first-stage oligomerization catalyst and the second-stage oligomerization catalyst of the above-described embodiment are either stacked or mixed within the same catalyst bed, resulting in both ethylene conversion and higher-grade olefin conversion occurring within the same catalyst bed. In one embodiment, the first-stage oligomerization catalyst may be stacked on top of the oligomerization catalyst such that the charged olefin encounters first-stage oligomerization, and then second-stage oligomerization. The first-stage oligomerization stream, containing the first-stage oligomerized olefin and the unconverted feed olefin, is collected from reactors 60+ and 60'+ into line 66+. The oligomerization recirculation stream in line 52+, containing the dimer, oligomer, and unconverted feed olefin, is recirculated to the oligomerization reactor 60+ via the charged olefin stream in line 54. The oligomerization product stream in line 84+ may be fed to olefin splitter column 90# for separation. The net liquid light olefin splitter overhead stream in line 106# and / or the oligomerization recirculation stream in line 52+ may contain at least 10% by weight, preferably at least 20% by weight, and more preferably at least 25% by weight of ethylene, which may be recirculated to the first-stage oligomerization reactor 60.

[0073] Figure 5 shows an embodiment in which the oligomerized olefin flow from the first stage in line 66+ is supplied to the interstage separator 160 for separation. Elements in Figure 5 having the same configuration as in Figure 4 will have the same reference numerals as in Figure 4. Elements in Figure 5 having a different configuration from the corresponding elements in Figure 4 will have the same reference numerals but will be designated with a dash (-). The configuration and operation of the embodiment in Figure 5 are essentially the same as those in Figure 4, with some notable exceptions.

[0074] Reactor 60+ may contain the first-stage oligomerization catalyst, similar to the embodiment in Figure 1. The first-stage oligomerization stream in line 66- may be fed to the interstage separator 160 to separate the overhead olefin stream containing unreacted ethylene in overhead line 162 from the bottom first-stage oligomerized olefin stream in bottom 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 olefin stream in line 162 may be mixed with the primary charge olefin stream in line 56-, which may then be charged into the first-stage oligomerization reactor 60+ or ​​60'+ in line 57 or 57', respectively. The overhead olefin stream in line 162 contains ethylene as its main component. The bottom first oligomerized olefin stream in bottom line 164 can be mixed with the unoligomerized recirculated olefin in olefin splitter overhead line 106- and the oligomerized recirculated stream in line 72- containing oligomerized olefin 72 to provide a charged second-stage oligomerized olefin stream in line 74-, which can be charged into the second-stage oligomerized olefin reactor 70 or 70'-. The second-stage oligomerized olefin reactors 70-, 70'- may contain a second-stage oligomerization catalyst. The second-stage oligomerization product stream in line 84- can be fed into olefin splitter column 90# for separation.

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

[0076] The charging olefin stream in line 54^ is delivered by the inter-bed olefin charging stream in line 58^ to the downstream catalyst beds 65, 67, 65', 67' in the inter-catalyst reactors 60+, 60'+. The downstream catalyst beds 65, 67, 65', 67' may contain the second-stage oligomerization catalyst, or may contain the first-stage oligomerization catalyst and the second-stage oligomerization catalyst stacked or mixed together. The oligomerized olefin stream in line 64 or 64', containing the dimers, oligomers, and unconverted feed olefins, is collected in line 66^. The oligomerized recirculation stream in line 52^, containing the dimers, oligomers, and unconverted feed olefins removed from the oligomerized olefin stream in line 66^, is recirculated to the oligomerization reactors 60+, 60'+ via the charging olefin stream in line 54^. The oligomerized olefin stream in line 66, having the same composition, can be taken out and fed into olefin splitter column 90 for separation.

[0077] In the olefin separation column 90#, fractionation is performed to separate the light olefins and intermediate olefins from the net liquid overhead stream and recirculate them to the oligomerization unit 50^ in line 106^. Preferably, the olefin splitter column 90# may operate to recirculate the unconverted ethylene back to the first-stage oligomerization unit 50^ in line 106^, while sending the heavier oligomers to hydrogenation in bottom line 94#. The net liquid light olefin splitter overhead stream in line 106^ may mainly consist of ethylene, which can be mixed with a paraffin-containing diluent stream, which may be the net stripped hydrogenation stream in line 148^, and then recirculated to the oligomerization reactors 60+ and 60'+ in primary charging lines 57^ and 57'^, respectively.

[0078] The net liquid overhead flow in line 106^ may be heated as needed and sent by primary charging lines 57^, 57'^ to upstream catalyst beds 61, 61' which may contain only the first-stage oligomerization catalyst or both the first-stage and second-stage oligomerization catalysts stacked or mixed together. Heat generation can be more easily controlled by recirculating the ethylene-dominant flow to the upstream end of reactors 60+, 60+'.

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

[0080] Example 1 The inventors prepared catalyst supports using the aforementioned oil-drop method, which combines synthesis and shaping in a single continuous step, in order to achieve compositional homogeneity in which aluminum oxide is completely dispersed in the silica matrix. One of the catalysts in Example 2 was prepared using CCIC powder containing a crystalline alumina phase dispersed throughout the support. An example of a catalyst support prepared by the predetermined method is shown in Table 1 below.

[0081] [Table 1]

[0082] Example 2 The catalysts shown in Table 2 were provided by incorporating metals onto the carriers shown in Table 1 using either an evaporation impregnation method or an ion exchange method.

[0083] [Table 2]

[0084] Example 3 To demonstrate the efficacy of the specified catalysts, the selective catalysts in Table 2 were tested for the conversion of ethylene to dimers and oligomers. In one apparatus, the test was performed at 6.2 MPa (g) (900 psig) using a chemical-grade ethylene feed. In this test, n-butane with a nominal ratio of 4 on a weight basis to n-butane to ethylene was used as a diluent to reduce the heat of reaction. The test was performed with catalyst loadings ranging from 15 to 40 grams and at a rate of 5.0 hours relative to the total feed rate. -1 The test was carried out at a combined feed rate to achieve WHSV. The catalyst was pre-dried in liquid nitrogen at a temperature of 300-400°C until the effluent dew point stabilized at approximately -40--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 flow was analyzed by online GC gas chromatography using PONA and alumina columns. The overhead gas was analyzed by alumina column for the paraffin / olefin / naphthene / aromatic compound ratio 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 at the effective space velocity using n-butane co-feed as a heat-mediated diluent, example catalysts 5.1 and 6.1 exhibited stable ethylene conversion rates exceeding 90%, C8-C16 distillate yields of 30-50% by weight and 10-15% by weight, respectively, approaching 400 hours and 200 hours, respectively.

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

[0089] [Table 5]

[0090] Comparative Example Catalyst 1.1 was prepared on the oil-droplet spherical support of Comparative Example 1. When Comparative Example Catalyst 1.1 was tested according to a predetermined test method, the catalytic activity, measured by the ethylene conversion rate and C8-C16 yield, was very low, particularly under conditions of appropriate diluent-to-ethylene ratios for heat transfer, as shown in Table 6 below. The catalyst contains a crystalline alumina phase as determined 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, the inventors evaluated various catalysts using C2-C4 olefin feed blends containing 30-40 wt% ethylene, 40-55 wt% propylene, and 15 wt% butylene, respectively, simulating 100% ethylene and olefin mixtures emerging from MTO units. In these tests, n-butane was also used as a diluent to mitigate reaction heat, 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 an MTT catalyst, Example 4.1; in another test, the ethylene feed was tested with nickel on an amorphous silica alumina (ASA) catalyst, Example 4.2. In a third test, Example 4.3, nickel on amorphous silica alumina and the MTT catalyst were tested in a layered configuration using a given C2-C4 olefin blend. Nickel was incorporated into Example 3's ASA oil-dropped spherical carrier of 75 SiO2 and 25% Al2O3 with a diameter of 1 / 16 inch using evaporation impregnation technique to obtain Example Catalyst 3.1. MTT catalysts were prepared by extruding 45 Si / Al2 zeolite MTT and 25 / 75 zeolite / alumina-formed boehmite alumina and calcining them in fluid air at 550°C for 2 hours. In the former, nickel was incorporated into amorphous silica-alumina spherical carriers of 75 mol% SiO2 and 25 mol% Al2O3, prepared using a predetermined oil-dropped process. In all three tests, n-butane was used as a diluent. Ten grams of the entire catalyst were loaded in Test Examples 4.1 and 4.2 using a predetermined C2-C4 olefin mixture and 100% ethylene, respectively. In Test Example 4.3, a laminated packing consisting of 5 grams of Example Catalyst 3.1, made from amorphous silica-alumina, and 10 grams of zeolite catalyst made from 25% MTT and 75% alumina was loaded. The results are shown in Table 7 below.

[0093] [Table 7]

[0094] As shown in Table 7, the 25% MTT and 75% alumina zeolite catalysts readily converted propylene and mixed butene to over 90% levels and ethylene to a moderate level of 40-70% when operated at an average bed temperature of 190-230°C in Test Example 4.1. Meanwhile, the nickel catalyst on amorphous silica-alumina oil droplets, Catalyst 3.1, converted nearly 100% of ethylene in Test Example 4.2. Test Examples 4.1 and 4.2 appear to demonstrate that nickel on amorphous silica-alumina oil droplets can achieve ethylene conversion rates 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, conversion rates of over 90% for ethylene, propylene, and mixed butene can be achieved 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-directed template, such as diisopropylethylamine (DIPEA), to a pre-prepared solution of phosphoric acid in cold / ice water. While stirring the reaction mixture, the desired amounts of silica and alumina precursors (Ludox colloidal silica and Versal 251 alumina, respectively) were added in a sequential manner at 5-minute intervals. The reaction slurry was then transferred to a reactor, which was sealed and heated from room temperature to 175°C at a rate of 15°C / hour and a stirring speed of 300 RPM. Once the set temperature was reached, the reactor was stirred for 48 hours, at which point the reactor was cooled, the product powder was washed, and isolated by centrifugation. The synthesized MTT was subjected to NH4+ ion exchange, dried, bonded with an alumina binder, and calcined. Alternatively, the synthesized MTT could be calcined before ion exchange and the addition of the binder. A typical yield, equal to the ratio of the mass of the calcined powder to the total mass of the synthesized gel, was 17%.

[0096] Specific Embodiments The following will be explained in conjunction with specific embodiments, but it should be understood that this explanation is intended to illustrate the scope of the above-mentioned description and the attached claims, and is not intended to limit them.

[0097] A first embodiment of the present disclosure is a process for oligomerizing an olefin flow, comprising: oligomerizing the olefin flow with a first-step oligomerization catalyst to provide a first-step oligomerized olefin flow; oligomerizing the first-step oligomerized olefin flow with an oligomerization catalyst to provide a second-step oligomerized olefin flow; and recirculating the ethylene flow to the first-step oligomerization step. Embodiments of the present disclosure are any or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, wherein the first-step oligomerization catalyst comprises a metal on amorphous silica alumina. Embodiments of the present invention are any or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, wherein the first-step or second-step oligomerization catalyst comprises one or more of MRE, TON, MTT, MFS, MFI, MEL, AFO, AEL, EUO, and FER. Embodiments of the present disclosure are one or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, wherein the oligomerization catalyst of the first step or the oligomerization catalyst of the second step is regenerated by contact with air. Embodiments of the present disclosure are one or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, further comprising recirculating the oligomerized olefin from the second step oligomerization flow to the second step oligomerization process. Embodiments of the present disclosure are one or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, further comprising recirculating the oligomerized olefin from the first step oligomerization process to the first step oligomerization process. Embodiments of the present disclosure are any or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, further comprising separating an ethylene stream from a first-stage oligomerized olefin stream and recirculating the ethylene stream to the first-stage oligomerization step. Embodiments of the present disclosure are any or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, wherein a second-stage oligomerization step is downstream of the first-stage oligomerization step.Embodiments of the present disclosure are any or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, further comprising separating an ethylene stream from the oligomerized olefin stream of the second stage and recirculating the ethylene stream to the first oligomerization step. Embodiments of the present disclosure are any or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, further comprising separating an intermediate olefin stream from the oligomerized olefin stream of the second stage and recirculating the intermediate olefin stream to the second oligomerization step. Embodiments of the present disclosure are any or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, further comprising separating a distillate olefin stream from the oligomerized olefin stream of the second stage and hydrogenating the distillate olefin stream. Embodiments of the present disclosure are any or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, and further include separating the hydrogenation distillate flow into a jet flow and a diesel flow. Embodiments of the present disclosure are any or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, and further include extracting a diluent flow from the hydrogenation distillate flow and charging the diluent flow into a first-stage oligomerization step. Embodiments of the present disclosure are any or all of the prior embodiments of this paragraph to the first embodiment of this paragraph, and the olefin flow comprises at least 5% by weight of ethylene.

[0098] A second embodiment of the present disclosure is a process for converting methanol into a distillate fuel, comprising contacting an oxygenated flow with an MTO catalyst to generate an olefin flow, and oligomerizing the olefin flow with an oligomerizing catalyst to generate an oligomerized olefin flow. Embodiments of the present disclosure are any or all of the prior embodiments to the second embodiment of the present paragraph, wherein the oligomerization step further comprises first oligomerizing the olefin flow with a first-stage oligomerizing catalyst to generate a first-stage oligomerized olefin flow, and oligomerizing the first-stage oligomerized olefin flow. Embodiments of the present disclosure are any or all of the prior embodiments described in this paragraph to the second embodiment described in this paragraph, and further include separating an ethylene flow and an intermediate olefin flow from an oligomerized olefin flow, recirculating the ethylene flow to a first-stage oligomerization process, and recirculating the intermediate olefin flow to a second-stage oligomerization process.

[0099] 1. A third embodiment of the present disclosure is a process for converting ethylene into a distillate fuel, 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. Embodiments of the present disclosure are any or all of the prior embodiments to the third embodiment of this paragraph, further comprising: separating an ethylene stream and an intermediate olefin stream from the second-stage oligomerized olefin stream; recirculating the ethylene stream to the first-stage oligomerization step; and recirculating the intermediate olefin stream to the second-stage oligomerization step. Embodiments of the present disclosure are one or all of the prior embodiments to the third embodiment of this paragraph, wherein the ethylene flow is obtained by contacting an oxygenated flow with an MTO catalyst to generate an olefin flow. Embodiments of the present disclosure are one or all of the prior embodiments to the third embodiment of this paragraph, wherein the olefin flow contains at least 5% by weight of ethylene.

[0100] A fourth embodiment of the present disclosure is a process for converting ethylene, comprising oligomerizing an ethylene stream with an oligomerization catalyst containing a group VIII metal on a silica aluminum oxide support, wherein the support contains 70-99.5% by weight of SiO2. Embodiments of the present disclosure are any or all of the prior embodiments to the fourth embodiment of this paragraph, wherein the support contains 70-99.5% by weight of SiO2. Embodiments of the present disclosure are any or all of the prior embodiments to the fourth embodiment of this paragraph, wherein the aluminum oxide is completely dispersed in a silica matrix. Embodiments of the present disclosure are any or all of the prior embodiments to the fourth embodiment of this paragraph, wherein the aluminum oxide is amorphous. Embodiments of the present disclosure are any or all of the prior embodiments to the fourth embodiment of this paragraph, wherein the support is synthesized and shaped in a single step. Embodiments of this disclosure are one or all of the prior embodiments to the fourth embodiment of this paragraph, and the carrier is an oil droplet sphere. Embodiments of this disclosure are one or all of the prior embodiments to the fourth embodiment of this paragraph, and the carrier has a surface area of ​​300 to 600 m² / g as determined by the nitrogen BET method. Embodiments of this disclosure are one or all of the prior embodiments to the fourth embodiment of this paragraph, and the porosity of the carrier, as measured by the total mercury intrusion volume, is 0.40 to 1.4 mL / g. Embodiments of this disclosure are one or all of the prior embodiments to the fourth embodiment of this paragraph, and the group VIII metal is nickel. Embodiments of this disclosure are one or all of the prior embodiments to the fourth embodiment of this paragraph, and the molar ratio of the group VIII metal to Al² in the catalyst is 0.04 to 0.70. Embodiments of the present disclosure are one or all of the prior embodiments described in this paragraph to the fourth embodiment described in this paragraph, wherein the catalyst also includes a group IA metal, and the molar ratio of the group IA metal to aluminum is 0.04 to 0.4.

[0101] Without further detail, it is expected that a person skilled in the art can use the foregoing description to the fullest extent without departing from the spirit and scope of the disclosure, and can easily identify the essential characteristics of the disclosure, and can make various changes and modifications to the disclosure to suit various uses and conditions. Accordingly, the prior preferred specific embodiments should be interpreted as merely illustrative examples and not in any way limiting the remainder of the disclosure, but are intended to cover various modifications and equivalent configurations that fall within the scope of the appended claims.

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

Claims

1. This is a process for oligomerizing olefin flows. As a first-stage oligomerization step, the olefin flow is oligomerized with a first-stage oligomerization catalyst to provide a first-stage oligomerized olefin flow. The oligomerized olefin flow from the first step is oligomerized with the oligomerizing catalyst from the second step to provide the oligomerized olefin flow from the second step. The second step involves separating the oligomerized olefin flow into a heavy olefin splitter overhead flow and a heavy olefin splitter bottom flow, The heavy olefin splitter bottom flow is subjected to hydrogenation treatment to provide a hydrogenated distillate flow, and the hydrogenated distillate flow is separated into a hydrogenation separator vapor flow and a hydrogenation separator liquid flow. The aforementioned hydrogenation separator liquid flow is separated into a stripper-off gas flow and a C9 + paraffin diluent flow, The C9 + paraffin diluent flow is introduced into the oligomerization step of the first stage, The heavy olefin splitter overhead flow is separated into a vapor overhead flow and a liquid light olefin splitter overhead flow containing ethylene. A process comprising recirculating the liquid light olefin splitter overhead flow to the first oligomerization step.

2. A process for converting methanol into distillate fuel, The process involves contacting a methanol stream with an MTO catalyst to generate an olefin stream, To recover the deoxygenated olefin stream from the aforementioned olefin stream, The deoxygenated olefin flow is selectively hydrogenated to produce a selectively hydrogenated olefin flow, The process comprises oligomerizing the selectively hydrogenated olefin stream in the process described in claim 1, wherein the selectively hydrogenated olefin stream is the olefin stream of the first oligomerization step. process.

3. The olefin stream is an ethylene stream, and the oligomerization catalyst in the first step is a catalyst containing a group VIII metal on an amorphous silica aluminum oxide support. Here, the support contains 87 to 98% by weight of SiO2, the contact time of the oligomerization catalyst in the first step is shorter than the contact time of the oligomerization catalyst in the second step, and the contact time of the oligomerization catalyst in the first step is less than 45% of the total contact time. The process according to claim 1.

Images