Use of recycled by-products to create hydrogen for use in the conversion of olefins into jet fuel.

Abstract

A process for supplying hydrogen in a process for producing jet fuel from methanol, comprising generating supply hydrogen from a hydrogen generation unit to be sent to one or more vessels in the process, wherein additional supply hydrogen is recovered from one or more recirculation flows from one or more reactors or vessels and sent to replenish the first supply hydrogen. The additional supply of hydrogen is recovered from a hydrogenation reactor downstream of an oligomerization reactor, a methanol synthesis reactor, or a hydrogenation reactor downstream of a DME water tower and sent to a methanol synthesis reactor, a hydrogenation reactor, or an acetylene converter.

Description

Technical Field

[0001] (Priority Claim) This application claims priority to U.S. Provisional Patent Application No. 63 / 502,333, filed May 15, 2023, the entire disclosure of which is incorporated herein by reference.

[0002] (Field of the Invention) The present invention relates to a process for producing liquid fuels from carbon dioxide and hydrogen. This field may particularly relate to a process for using recycled hydrogen to supplement hydrogen produced by a hydrogen production unit.

Background Art

[0003] The viability of methanol-jet (MTJ) composites in the market highly depends on the overall carbon intensity of the entire CO2-jet process of which MTJ is a part. There are additional upstream units including an electrolyzer and methanol synthesis. CO2 capture may also be present.

[0004] Renewable fuels produced from methanol have at least two process steps that each require additional hydrogen, namely the hydrogenation of CO2 to methanol and a hydrogenation unit to produce fuel-grade hydrocarbons from upstream dimerization and / or oligomerization units. The additional H2 can come from hydrocarbons with or without carbon capture (gray and blue H2 respectively), or most likely from renewable energy sources (green H2). Those hydrogen are costly to generate. It is desirable to reduce the cost of hydrogen required for renewable fuels.

Summary of the Invention

[0005] A process for providing hydrogen in a process for producing jet fuel from methanol is provided, comprising generating feed hydrogen from a hydrogen generation unit to be delivered to one or more vessels in the process, wherein additional feed hydrogen is recovered from one or more recirculation flows from one or more reactors or vessels and delivered to replenish the first feed hydrogen. The additional feed hydrogen is taken from a reaction in which hydrogen is not completely consumed. The additional feed hydrogen may be taken from a purification or hydrogen concentration unit. [Brief explanation of the drawing]

[0006] [Figure 1] This is a schematic diagram of a methanol synthesis process according to an exemplary embodiment of the process for producing the light olefins of the present disclosure. [Figure 2] This is a schematic diagram of an MTO process according to an exemplary embodiment of the process for producing the light olefins of this disclosure. [Figure 3] This diagram shows a schematic representation of the olefin supply raw material preparation process and apparatus in this disclosure. [Figure 4] This is a schematic diagram illustrating the overall process of this disclosure. [Figure 5] This is a schematic diagram of the light olefin fractionation portion of the process described herein. [Figure 6] This is a schematic diagram of the hydrogenation portion of the process disclosed herein.

[0007] definition The term "communication" means that fluid flow is operably permitted between the listed components, and this can be characterized as "fluid communication."

[0008] The term "downstream communication" means that at least a portion of the fluid flowing to the downstream-communicating object can be operably flowed from the fluid-communicating object.

[0009] The term "upstream communication" means that at least a portion of the fluid flowing from an upstream communication object can flow operably into a fluid-communicated object.

[0010] The term "direct communication" means that the fluid flow from the upstream component enters the downstream component without passing through any other intervening container.

[0011] The term "indirect communication" refers to the flow of fluid from an upstream component entering a downstream component after passing through an intervening container.

[0012] The term “column” refers to a distillation column (singular or plural) for separating one or more different volatile components. Unless otherwise specified, each column includes a condenser at the top to condense and reflux a portion of the top flow returning to the top of the column, and a reboiler at the bottom to vaporize a portion of the bottom flow and return it to the bottom of the column. The feedstock to the column may be preheated. The top pressure is the pressure of the top vapor at the column's vapor outlet. The bottom temperature is the liquid bottom outlet temperature. The top line and bottom line refer to the net lines from column to column downstream of any reflux or reboil. A stripping column may omit the reboiler at the bottom of the column, but instead may provide heating requirements and separation propulsion from fluidizing inert media such as vapor. A stripping column typically feeds the feedstock to an upper tray and removes the main product from the bottom of the column.

[0013] As used herein, the term “diesel” means hydrocarbons that boil within the “diesel cutpoint” range, including IBP at 125°C (257°F) to 175°C (347°F) or T5 at 150°C (302°F) to 200°C (392°F), and T95 at 343°C (650°F) to 399°C (750°F) using the TBP distillation method, or T90 at 280°C (536°F) to 340°C (644°F) using ASTM D-86. The term “green diesel” means diesel containing hydrocarbons that are not of fossil fuel origin.

[0014] As used herein, the terms "T5," "T90," or "T95" mean the boiling point of 5 mass percent, 90 mass percent, or possibly 95 mass percent of the sample, respectively, as may be the case when using ASTM D-86 or TBP.

[0015] As used herein, the term “end point” (EP) means, in some cases, the temperature at which the sample has completely evaporated, as may be the case when using ASTM D-7169, ASTM D-86, or TBP.

[0016] As used herein, the term “jet fuel” means hydrocarbons that boil at a T10 of 190°C (374°F) to 215°C (419°F) and an endpoint of 290°C (554°F) to 310°C (590°F). The term “green jet fuel” means jet fuel containing hydrocarbons that are not of fossil fuel origin.

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

[0018] In this specification, the terms “concentrated stream” or “concentrated stream” mean that the concentrated stream coming out of the container has a higher concentration of components than the raw materials supplied to the container, and preferably all other streams being taken out of the container.

[0019] In this specification, the terms “dilute stream” or “dilute stream” mean that the dilute stream leaving the container has a lower concentration of the component than the supply to the container, and preferably all other streams withdrawn from the container.

[0020] As used herein, the term “concentrated” means more than 50%, preferably more than 75%, and more preferably more than 90%.

[0021] As used herein, the term "separator" means a vessel having an inlet, at least a top vapor outlet and a bottom liquid outlet, and optionally also having an aqueous outlet from the boot. A flash drum is a type of separator that can be in downstream communication with a separator that can operate at a higher pressure.

DETAILED DESCRIPTION OF THE INVENTION

[0022] The processes and apparatus described herein are one embodiment of possible configurations. Simpler configurations, as well as the multi-stage oligomerization processes described below, are contemplated.

[0023] The disclosed processes and apparatus include the production of liquid fuels from carbon dioxide and hydrogen. The process includes reacting a mixture of carbon dioxide and hydrogen to produce methanol, carbon monoxide, and water. The methanol is contacted with an MTO catalyst to produce an olefin stream. The olefin stream is oligomerized with an oligomerization catalyst to produce an oligomerized olefin stream comprising jet fuel, diesel fuel, and alkanes. The oligomerized olefin stream is separated into (1) a liquid fuel stream and (2) an alkane stream. A syngas stream comprising carbon oxides and hydrogen is produced by (1) reforming the alkane stream with steam in a steam reformer, an autothermal reformer, or a dry reformer, or (2) partially oxidizing the alkane stream.

[0024] The conversion of methanol to a liquid fuel stream, such as sustainable aviation fuel (SAF), is highly selective, but the light and heavy by-products of the reaction must be disposed of. One way to improve the overall selectivity of CO2 to the liquid fuel stream (and reduce the carbon intensity), for example, to a jet complex, is to recycle heavier hydrocarbons through a route from CO2 to syngas, such as reforming (steam reforming, autothermal reforming or dry reforming) or partial oxidation. The resulting syngas can be recycled to the methanol synthesis unit, thereby increasing the jet yield of CO2 to the jet facility.

[0025] Excess oxygen from the hydrolyzer can be introduced into a partial oxidation reactor to convert the by-product into synthesis gas, which can then be used as a ready-to-use feedstock for the methanol synthesis unit.

[0026] The methane and hydrogen produced by this process can be mixed with CO2 and reacted without the use of vapor. This requires a non-precious metal catalyst developed by Linde and BASF. The resulting synthesis gas can be used to produce methanol. Small amounts of other light hydrocarbons may also be treatable in this manner. The advantage of this route is that it produces synthesis gas without requiring hydrogen from the electrolytic cell, which is the most energy-intensive part of the CO2 jet complex.

[0027] The liquid byproduct from the MTJ complex is mixed with oxygen from the hydrolysis unit and reacted on a catalyst to convert hydrocarbons to CO. To improve the overall CO yield, CO2 can be co-supplied as a raw material. This process is exothermic when it reaches a sufficiently high temperature.

[0028] There are two main alternatives to this approach. (1) High-temperature steam reforming using CO2 and steam produces a synthesis gas mixture that can be directly supplied as raw materials for methanol synthesis. This is an endothermic process and therefore requires a considerable amount of external heat. (2) Autothermal reforming combines both partial oxidation and steam reforming. Oxygen, steam, and CO2 from the hydrolysis unit react on a catalyst to produce a synthesis gas mixture suitable for conversion to methanol. Methane and hydrogen (and possibly ethane) from the MTJ off-gas stream are mixed with CO2 to form a synthesis gas mixture of H2 and CO. This is an endothermic process but can act as a heat sink from other exothermic processes.

[0029] A third alternative method, dry reforming, can also be used. Methane and carbon dioxide react on a catalyst in an endothermic process to form a synthesis gas mixture of H2 and CO.

[0030] In the oligomerization process, the flow leaving the depropane unit is sent to a reactor, where it is converted to synthesis gas via steam reforming, partial oxidation, autothermal reforming, or dry reforming. The synthesis gas is then sent to a methanol synthesis unit. When partial oxidation or autothermal reforming is selected to produce synthesis gas, air or oxygen must also be supplied, although additional steam can be extracted from the reactor for use elsewhere in the facility. If steam reforming is selected to produce synthesis gas from ethane / propane, steam feedstock (which can be supplied entirely by a methanol-jet process) and energy input (which can be supplied partially by steam from a methanol-jet process) are also required. Liquid fuels can be produced from carbon dioxide and hydrogen through the following process: (a) A step of reacting a mixture of carbon dioxide and hydrogen to produce methanol, carbon monoxide, and water. (b) A step of generating an olefin flow by contacting methanol with an MTO catalyst, (c) A step of oligomerizing the olefin stream with an oligomerizing catalyst to produce an oligomerized olefin stream containing jet fuel, diesel fuel, and alkanes. (d) A step of separating the oligomerized olefin flow into (1) a liquid fuel flow and (2) an alkane flow, (e) The following, (1) Reforming the alkane stream with steam in a steam reforming reactor, a self-thermal reforming reactor, or a dry reforming reactor, or (2) A step of generating a synthesis gas stream containing carbon dioxide and hydrogen by partially oxidizing an alkane stream.

[0031] The process of producing liquid fuel from carbon dioxide and hydrogen may include the following steps: (a) A step of reacting a mixture of carbon dioxide and hydrogen to produce a crude methanol stream containing methanol, water, and other contaminants including one or more of hydrogen, CO, CO2, methane, ethanol, and other oxygenated hydrocarbons. (b) A step of purifying the crude methanol stream by distillation to remove light contaminants into the first exhaust gas stream, and removing heavy contaminants and water into the second heavy waste stream to produce a purified methanol stream. (c) A step of contacting a purified methanol stream with an MTO catalyst to produce a crude olefin stream containing ethylene, propylene, butylene, and other contaminants including one or more of hydrogen, CO, CO2, methane, dimethyl ether, ethanol, and other oxygenated hydrocarbons. (d) A step of purifying the crude olefin stream by distillation to remove the light fraction into a third exhaust gas stream, then absorbing water to remove heavy oxygenated hydrocarbons, extracting water to remove dimethyl ether into a DME (dimethyl ether) recirculation stream, and generating one or more purified olefin streams. (e) A step of reacting one or more purified olefin streams with one or more oligomerization catalysts using one or more reaction vessels to produce a crude oligomerized olefin stream containing oligomerized olefins having a carbon length of 4 to 24 carbon atoms and contaminants containing one or more of hydrogen, methane, and alkanes lighter than pentane. (f) A step of removing hydrogen and light alkanes from the crude oligomerized olefin stream by fractional distillation, into (1) a fourth exhaust gas stream, (2) a recirculated stream of light oligomerized olefins, and (3) a purified oligomerized olefin stream containing olefins having a carbon length of 12 to 24 carbon atoms. (g) A step of preparing a crude jet fuel flow by reacting a purified oligomerized olefin flow with hydrogen using a hydrogenation catalyst to saturate the olefin with paraffin. (h) A step of fractionally distilling the crude jet fuel stream by distillation to remove excess hydrogen and light hydrocarbons into (1) a fifth exhaust stream, (2) a sixth exhaust stream containing naphtha, (3) a seventh exhaust stream containing diesel, and (4) liquid fuel containing jet fuel. (i) Below: (1) In a steam reforming process, a self-thermal reforming process, or a dry reforming process, one or more exhaust gas flows are reformed using steam, or (2) A step of generating a recirculated synthesis gas flow containing carbon dioxide and hydrogen by partially oxidizing one or more of the exhaust gas flow, (j) A step in which the recirculated synthesis gas flow is supplied as a raw material to the methanol synthesis process of step A.

[0032] One or more exhaust gas streams may also be reacted with oxygen in a partial oxidation reactor to produce a second recirculated synthesis gas stream containing CO and hydrogen.

[0033] The reaction in step (i) can also be carried out by reacting one or more flue gas streams with oxygen and steam in a self-thermal reactor to produce a third recirculated synthesis gas stream containing CO and hydrogen. The reaction in step (i) can also be carried out by reacting a sixth flue gas stream with steam in a steam reforming process, by reacting a seventh flue gas stream with steam in a steam reforming process, or by partial oxidation, and the resulting recirculated synthesis gas stream is thermally integrated with the fractional distillation process of step (j) to provide some or all of the energy required for distillation.

[0034] Hydrogen for steps (a) and / or (g) may be produced by a water electrolysis unit, and / or oxygen for self-thermal reforming may be produced by a water electrolysis unit. If the reaction in step (i) is carried out by partial oxidation, the oxygen for partial oxidation may be produced by a water electrolysis unit. If the reaction in step (i) is carried out using steam in a steam reforming process, one or more exhaust gas streams supplied to the steam reforming reactor may contain 10% to 50% by weight of propane. The reaction in step (i) can also be carried out using a dry reforming process. A water electrolysis unit (electrolytic cell) produces hydrogen and oxygen from water.

[0035] Methanol synthesis using MTO reactor Carbon dioxide is a so-called greenhouse gas, and its concentration in the atmosphere is desired by many. Carbon dioxide can be converted into 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 facilitate the conversion of oxygenated substances into hydrocarbon mixtures, especially hydrocarbon mixtures consisting mostly of light olefins. A highly efficient methanol-to-olefin (MTO) process can convert oxygenated substances into light olefins, which have typically been considered for plastic production. Light olefins produced from the MTO process contain high concentrations of ethylene.

[0036] Methanol is converted to light olefin products in the methanol-to-olefin (MTO) process. Molecular sieves such as microporous crystalline zeolites and non-zeolite catalysts, particularly silicoaluminophosphates (SAPOs), are known to facilitate the conversion of oxygenated materials such as methanol into hydrocarbon mixtures, especially those composed primarily of light olefins. SAPO catalysts and their formulations are generally taught in U.S. Patents 4,499,327(A), 10,358,394, and 10,384,986. The light olefins produced from the MTO process are concentrated with ethylene and propylene, but also contain C4-C6 olefins.

[0037] A mixture of carbon dioxide and hydrogen is reacted to produce methanol, carbon monoxide, and water. The resulting methanol is then contacted with an MTO catalyst to generate an olefin flow.

[0038] The methanol stream is introduced into the MTO reactor and, under MTO reaction conditions, contacts the MTO catalyst to convert methanol into olefins and water. The methanol stream may contain methanol, dimethyl ether, ethanol, or a combination thereof. The MTO reaction conditions involve contact with the SAPO catalyst at a pressure of 2 MPa to 3.8 MPa. The MTO reaction temperature should be 325 to 450°C. The weight hourly space velocity (WHSV) in the MTO reactor is in the range of 2 to 15 per hour. The MTO catalyst is separated from the product olefin stream after the MTO reaction.

[0039] Referring to Figure 1, the integrated process and apparatus 101 for producing light olefins includes a methanol synthesis section 111 and a methanol purification section 208. As shown in Figure 1, the synthesis gas stream in line 122 and the hydrogen gas stream in line 124 are fed to the methanol synthesis section 111. Synthesis gas is defined as a gas primarily consisting of carbon monoxide (CO), carbon dioxide (CO2), and hydrogen (H2). Optionally, the synthesis gas may also include methane (CH4), as well as small amounts of ethane and propane. Conventional processes for converting carbon components into synthesis gas include steam reforming, partial oxidation, autothermal reforming, and combinations of these processes. According to embodiments of this disclosure, the synthesis gas stream in line 122 can be taken from any suitable source. According to another embodiment of this disclosure, the hydrogen gas stream in line 124 can be taken from any suitable source. In an exemplary embodiment, the hydrogen gas stream in line 124 is taken from a pressure swing adsorption (PSA) unit.

[0040] According to exemplary embodiments of the present disclosure, the methanol synthesis section 111 comprises a first methanol converter 140 and a second methanol converter 160. The synthesis gas stream in line 122 and the hydrogen gas stream in line 124 are fed to the first methanol converter 140 of the methanol synthesis section 111. In embodiments, the synthesis gas stream in line 122 and the hydrogen gas stream in line 124 can be mixed to provide a mixed feed stream 126, which is then fed to the first methanol converter 140. However, the synthesis gas stream in line 122 and the hydrogen gas stream in line 124 may be fed to the first methanol converter 140 separately. The mixed feed stream 126 can be fed to a synthesis gas booster compressor 130 to compress the synthesis gas to a specific pressure and provide a compressed synthesis gas stream in line 132 before being fed to the first methanol converter 140. In an exemplary embodiment, the synthesis gas can be compressed to a pressure of 6890 kPa (1000 psia) to 8970 kPa (1300 psia) in a synthesis gas booster compressor 130. The synthesis gas flow can be heated before being sent to a first methanol converter 140. The compressed synthesis gas flow in line 132 can be heat-exchanged with the first reactor outflow flow in line 144 in a heat exchanger 133, thereby providing a heated synthesis gas flow to line 134. The heated synthesis gas flow in line 134 is sent to the first methanol converter 140.

[0041] In the first methanol converter 140 of the methanol synthesis section 111, the synthesis gas is converted into a methanol composition. The methanol synthesis process is achieved in the presence of a methanol synthesis catalyst. In an exemplary embodiment, the synthesis gas stream in line 122 to the methanol synthesis section 111 has a carbon dioxide to carbon monoxide molar ratio of 1:2 to 1:4 and a hydrogen to carbon oxide (CO+CO2) molar ratio in the range of 3:2 to 3:1.

[0042] Suitable methanol synthesis catalysts may include zinc oxide and copper on an alumina support. The synthesis conditions for the first methanol converter 140 in methanol synthesis section 111 may be a temperature of 200-300°C and a pressure of 3.5-10 MPa. Reaction equilibrium usually requires methanol separation and recycling of unreacted reagents back into the synthesis reaction to obtain sufficient conversion.

[0043] According to an exemplary embodiment, the first methanol converter 140 can operate at temperatures ranging from 204°C (400°F) to 290°C (550°F). According to another exemplary embodiment, the first methanol converter 140 can operate at pressures ranging from 6890 kPa (1000 psia) to 8970 kPa (1300 psia).

[0044] The methanol synthesis reaction is highly exothermic. Boiler feedwater (BFW) in line 148 is sent to the first methanol converter 140, generating a steam flow in line 142, which is taken out of the first methanol converter 140. The generation of steam absorbs the heat generated in the methanol synthesis reaction. The steam flow in line 142 is sent to the top separator 145, which separates the steam in line 146 from the water flow in line 147. The water flow in line 147 is replenished by the recirculated BFW in line 149, providing BFW from line 148 for the first methanol converter 140.

[0045] In the first methanol converter 140, the synthesis gas is converted into a methanol composition in the first reactor effluent containing methanol in line 144. The methanol stream in the first reactor effluent in line 144 may contain methanol, dimethyl ether, ethanol, or a combination thereof. The first reactor effluent in line 144 is heat-exchanged with the compressed synthesis gas stream in line 132 in the heat exchanger 133. The heat-exchanged first reactor effluent in line 135 can be heated in the heater 131 to provide the heated first reactor effluent in line 136. The heated first reactor effluent in line 136 can be further heated in the heater 137 to provide the further heated first reactor effluent in line 138. The further heated first reactor effluent in line 138 is separated in the first gas-liquid separator 150 to provide the first vapor stream in line 152 and the first liquid stream in line 154. The first vapor stream in line 152 and the first liquid stream in line 154 can be further processed to recover methanol.

[0046] The first vapor flow in line 152 contains carbon dioxide that has not yet been converted to methanol. The first vapor flow in line 152 can be compressed in the first compressor 155. In an embodiment, the first vapor flow in line 152 can be combined with a feed hydrogen flow in line 153 to provide a combined first vapor flow in line 156. The combined first vapor flow in line 156 is compressed in the first compressor 155 to provide a compressed first vapor flow in line 157 at a pressure of 6890 kPa (1000 psia) to 8970 kPa (1300 psia). In an embodiment, the feed hydrogen flow in line 153 can be taken from any suitable source. According to this disclosure, the feed hydrogen flow in line 153 may be taken from one or more units of process 101.

[0047] The compressed first vapor flow in line 157 is heat-exchanged with the second reactor outflow flow in heat exchanger 163 to provide a heat-exchanged first vapor flow in line 158, which is sent to the second methanol converter 160. In the second methanol converter 160 of methanol synthesis section 111, unconverted carbon dioxide in the synthesis gas is converted into methanol composition. The methanol synthesis process is achieved in the presence of a methanol synthesis catalyst. A suitable methanol synthesis catalyst may be zinc oxide and copper on an alumina support. The synthesis conditions in the second methanol converter 160 of methanol synthesis section 111 may be a temperature of 200-300°C and a pressure of 3.5-10 MPa. Reaction equilibrium typically requires methanol separation and recirculation of unreacted reagents into the synthesis reaction.

[0048] The boiler feedwater (BFW) in line 176 is sent to the second methanol converter 160, which generates a steam flow in line 166, taken out of the second methanol converter 160, and manages the heat generation. The steam flow in line 166 is sent to the top separator 172, which separates the steam in line 171 from the water flow in line 173. The water flow in line 173 is replenished by the recirculated BFW in line 174, which provides BFW in line 176 for the second methanol converter 160.

[0049] In the second methanol converter 160, the first reactor effluent is converted into a methanol composition to provide a second reactor effluent containing methanol in line 162. The methanol flow in the second reactor effluent in line 162 may include methanol, dimethyl ether, ethanol, or a combination thereof. The second reactor effluent in line 162 can be withdrawn from the side of the second methanol converter 160. The second reactor effluent in line 162 is heat-exchanged with the compressed first vapor in line 157 in a heat exchanger 163. The heat-exchanged second reactor effluent in line 164 can be heated in a heater 165 to provide a heated second reactor effluent in line 166a. The heated second reactor effluent in line 166a is separated in a second gas-liquid separator 180 to provide a second vapor flow in line 182 and a second liquid flow in line 184. The second vapor stream in line 182 and the second liquid stream in line 184 can be further processed to recover methanol.

[0050] According to an exemplary embodiment, the second methanol converter 160 operates at a temperature of 204°C (400°F) to 290°C (550°F). According to another exemplary embodiment, the second methanol converter 160 operates at a pressure of 6890 kPa (1000 psia) to 8970 kPa (1300 psia).

[0051] According to this disclosure, the second vapor flow in line 182 is sent to the PSA unit 185, where hydrogen is separated from the second vapor flow in line 182. In an exemplary embodiment, the second vapor flow in line 182 may be separated into a recirculation flow in line 153 and a PSA feed material flow in line 184a. In another exemplary embodiment, the recirculation flow in line 153 may be sent to the first compressor 155 as a replenishment hydrogen flow. In an embodiment, the replenishment hydrogen flow in line 153 to the first compressor 155 includes the recirculation flow in line 153.

[0052] The PSA feedstock flow in line 184a is processed in PSA unit 185. Typically, a PSA unit includes a series of adsorption beds, each containing one or a combination of adsorbents suitable for adsorbing specific components to be adsorbed. These adsorbents include, but are not limited to, activated alumina, silica gel, activated carbon, zeolite molecular sieve-type materials, or any combination thereof. The adsorbents are organized in any order required by the adsorption process to adsorb impurities or components. In PSA unit 185, the PSA feedstock gas flows over the adsorbents, and components that are more readily adsorbable are adsorbed during the adsorption process. The remaining gas leaves the adsorption beds in the PSA top gas flow 186, which is concentrated with components and impurities. When the adsorbents reach their adsorption capacity, they are regenerated to prevent hydrogen leakage into the PSA top gas flow 186.

[0053] In the PSA unit 185, hydrogen present in the PSA feed material flow in line 184a is separated into a hydrogen-concentrated flow in line 124. As shown in the figure, the purge flow in line 186 from the PSA unit is separated from the hydrogen-concentrated flow in line 124. The purge flow in line 186 can be used as fuel. In an exemplary embodiment, the hydrogen-concentrated flow in line 187 can be sent as a hydrogen flow to the synthesis gas booster compressor 130. In an embodiment, the hydrogen flow in line 124 to the synthesis gas booster compressor 130 includes the hydrogen-concentrated flow in line 187.

[0054] Returning to the second gas-liquid separator 180, the second liquid flow in line 184 is taken out from the bottom of the second gas-liquid separator 180 and sent to the third gas-liquid separator 190. The first liquid flow in line 154 may also be sent to the second gas-liquid separator 180. In an exemplary embodiment, the second liquid flow in line 184 may be combined with the first liquid flow in line 154 to provide a combined liquid flow in line 188, which is sent to the third gas-liquid separator 190. In the third gas-liquid separator 190, the first liquid flow in line 154 and the second liquid flow in line 184 are separated into a third vapor flow in line 192 and a third liquid flow in line 194. The third liquid flow in line 194 contains crude methanol. Alternatively, the third liquid flow in line 194 may be a crude methanol flow. The crude methanol stream may contain at least 100 ppmw of carbon oxides and / or at least 100 ppmw of C2+ oxygenates.

[0055] Crude methanol contains methanol, light fractions, and heavier alcohols. Where used and described herein, the terms “crude methanol” or “crude oxygenated feedstock” may include methanol, ethanol, water, light fractions, and fuel off. Light fractions may include ethers, ketones, aldehydes, and dissolved gases such as hydrogen, methane, carbon oxides, and nitrogen. Crude methanol contains fusel oil. Fusel oil in crude methanol typically contains higher alcohols and is commonly burned as fuel in methanol plants. Crude methanol containing fusel oil can be sent to an oxygenated conversion unit for further production of light olefins. According to this disclosure, crude methanol may be sent as feedstock to an oxygenated conversion unit or an MTO unit.

[0056] According to embodiments of the present invention, crude methanol may have a composition with concentrations of 0-1% by weight of carbon monoxide, 0.05-2% by weight of carbon dioxide, 0.001-2% by weight of methane, 0.05-2% by weight of hydrogen, 0-1% by weight of oxygen, 5-18% by weight of water, 0-1% by weight of nitrogen, 75-90% by weight of methanol, and 0.05-4% by weight of alcohol (excluding methanol).

[0057] The third liquid flow in line 194 may be sent to the crude methanol hold-up tank 195. The crude methanol flow in line 196 is taken out of the crude methanol hold-up tank 195. According to the present invention, the crude methanol flow in line 196 may be sent to the oxygenate conversion unit 200, as shown in Figure 1.

[0058] Conventionally, the crude methanol stream in line 196 is purified of light gases and heavy oxygenates before being fed into the MTO reactor 202 of the oxygenate conversion unit 200. According to the embodiment of Figure 1, at least 100 ppmw of carbon oxides and / or at least 100 ppmw of C 2+ The crude methanol stream containing oxygenated material is vaporized in the heating furnace and then sent directly to the MTO reactor 202.

[0059] The superheated crude methanol stream in line 199' is introduced into MTO reactor 202 and contacted with the MTO catalyst under MTO reaction conditions to convert methanol and other oxygenates into olefins and water. The crude methanol stream in line 198 may contain methanol, dimethyl ether, ethanol, or a combination thereof. MTO reactor 202 can fluidize the catalyst under high-speed flow conditions. 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. MTO reaction conditions include contact with the SAPO catalyst at a pressure of 2 MPa to 3.8 MPa. The MTO reaction temperature should be 325 to 450°C. The gravitational space velocity (WHSV) in MTO reactor 202 is in the range of 2 to 15 per hour. The MTO catalyst is separated from the olefin product stream after the MTO reaction.

[0060] In the MTO process, catalyst particles are repeatedly circulated between the MTO reactor 202 and the MTO regenerator unit 204. During regeneration, coke deposited on the catalyst particles during the reaction in the reaction zone is removed at high temperatures by oxidation in the regenerator unit 204. Removal of the coke deposits restores the activity of the catalyst particles to a point where they can be reused in the MTO reactor 202. The regenerated catalyst is discharged from the regenerator unit 204 in line 206 and recirculated back to the MTO reactor 202. The MTO effluent, including ethylene, propylene, and other olefins, along with water and oxygenated materials, is discharged from the MTO reactor 202 in effluent line 207.

[0061] According to an exemplary embodiment, the crude methanol stream in line 196 may be sent to a methanol purification section 208, which includes at least two distillation columns, a first distillation column 210 and a second distillation column 220. The crude methanol stream in line 196 is heat-exchanged with the product stream in a heat exchanger 197 to provide a heat-exchanged crude methanol stream in line 194 or a heat-exchanged third liquid stream in line 198. The heat-exchanged crude methanol stream in line 198 may be sent to the first distillation column 210. In the first distillation column 210, light gases are separated from the crude methanol in the top stream of the first distillation column in line 212. The light gases separated from the crude methanol stream include carbon monoxide, carbon dioxide, methane, hydrogen, and dimethyl ether. The top stream of the first distillation column in line 212 is sent to a first top receiver 215, where the light gases are separated into a first top receiver vapor stream in line 214. The first top receiver vapor flow in line 214 can be sent to the fuel section or possibly used as fuel in combustor 254 in line 256 in Figure 1. From the first top receiver 215, the top receiver liquid flow is taken out in line 216 and sent to the top of the first distillation column 210.

[0062] The bottom flow of the first distillation column containing methanol in line 218 is removed for further separation. The bottom flow of the first distillation column in line 218 is separated into a first reboiling flow in line 218b and a first distillation column outflow flow in line 218a. The first reboiling flow in line 218b is reboiled in reboiler 219 before being sent to the bottom of the first distillation column. According to an exemplary embodiment, the first distillation column 210 is operated at a pressure of 689 kPa (100 psia) to 1379 kPa (200 psia). According to another exemplary embodiment, the first distillation column is operated at a temperature of 27°C (80°F) to 177°C (350°F).

[0063] The first distillation column outflow in line 218a contains heavy oxygenated substances such as C2+ alcohols, ketones, and aldehydes that should be removed from the crude methanol stream. Therefore, the first distillation column outflow in line 218a is further separated in the second distillation column 220. In the second distillation column 220, the first distillation column outflow in line 218a is separated into the second distillation column top flow in line 222, which contains methanol, and the second distillation column bottom flow in line 226. The second distillation column top is in the gas phase. Instead of condensing it, the methanol feed stream is taken in from the second distillation column top line 222 and supplied to the MTO reactor 202 in line 199'. A portion of the top flow in line 222 may be condensed in line 228 for reflux into the column. From the heat exchanger 223, the partially condensed top flow of the second distillation column in line 224 is sent to the second top receiver 225. In the second top receiver 225, the condensed portion of the top flow of the second distillation column in line 224 is recirculated to the second distillation column 220.

[0064] The second distillation column bottom flow in line 226 is removed from the column. The second distillation column bottom flow in line 226 is separated into the second reboiling flow in line 226b and the second distillation column outflow flow in line 226a. The second reboiling flow in line 226b is reboiled in reboiler 230 before being sent to the second distillation column bottom section.

[0065] According to an exemplary embodiment, the second distillation column operates at a pressure of 517 kPa (75 psia) to 862 kPa (125 psia). Furthermore, according to an exemplary embodiment, the second distillation column operates at a temperature of 104°C (220°F) to 149°C (300°F). The outflow logistics of the second distillation column in line 226a include heavy oxygenated material and water, and aqueous oxygenated logistics.

[0066] In the embodiment shown in Figure 1, heavy oxygenates in the crude methanol stream are processed together with the MTO effluent stream in line 207 and separated into an oxygenate-concentrated stream. According to embodiments of the present disclosure, the methanol purification section 208 may also include a third distillation column (not shown) for further removal of heavy oxygenates from the crude methanol stream. According to exemplary embodiments, the third distillation column may operate at a pressure of 106 kPa (15 psia) to 345 kPa (50 psia). According to other embodiments of the present invention, the third distillation column may be an atmospheric distillation column operating at atmospheric pressure. According to further exemplary embodiments, the second distillation is performed at a temperature of 48°C (120°F) to 122°C (250°F).

[0067] If a third column is also used, the outflow from the second distillation column in line 226a is separated in the third distillation column to provide a top flow and a bottom flow containing methanol. The top flow, along with the methanol product flow from line 222 in MTO feed line 199', can be sent to the MTO reactor 202.

[0068] Preparation of olefin feedstock Figure 2 shows the olefin feedstock preparation process and apparatus of the present disclosure. In process and apparatus 10, the superheated MTO effluent of light olefins from the effluent line 207 of Figure 1 is supplied as feedstock via line 12 to an oxygenate conversion reactor 16 that reacts an oxygenate such as methanol or dimethyl ether (DME) with a fluidized catalyst. The high-temperature vapor reactor effluent in line 14 can be pre-cooled in a reactor effluent heat exchanger 15 to recover heat before being sent to the quenching tower 20. In the quenching tower 20, the vapor reactor effluent is deheated, organic acids are neutralized, and catalyst particles are removed by direct contact with a water stream supplied in line 19, which may be taken in from a stripped water stream in line 21. The quenched reactor effluent in line 22 is discharged from the quenching tower 20 and supplied as feedstock to a product separator tower 24. The product separator tower 24 may be downstream in communication with the MTO reactor 202 of Figure 1.

[0069] The product separator 24 comprises two sections for separating the reactor effluent into a product olefin flow in the top line 40, an intermediate liquid flow in the intermediate line 28, and a water flow in the bottom line 26. The first, i.e., lower section, receives the quenched reactor effluent in line 22. In the lower section, most of the heat is removed from the quenched reactor effluent while partially condensing the water in the quenched reactor effluent, producing a product water flow in the bottom line 26 containing a portion of the oxygenated by-products in the quenched reactor effluent in line 22. A portion of the product water flow is cooled and pumped to the top of the first section of the product separator 24 to cool the quenched reactor effluent in line 22. A second portion of the second bottom flow 26 is sent to the water stripper column 30. A water return flow containing oxygenated by-products from the compression section 80 in the return line 32 may also be passed to the water stripper column 30. The water stripper tower 330 may be connected downstream to the product separator tower 24.

[0070] The vapor flow from the first section of the product separation column 24 is sent to the second section of the product separation column, i.e., the upper section. The intermediate flow in line 28, containing hydrocarbons, oxygenated by-products, and liquid-phase water, is removed at the bottom of the upper section. A portion of the intermediate flow in line 28 is cooled and sent as reflux to the upper part of the second section of the product separation column 24. The remainder of the intermediate flow in line 28 is passed to the coalescer 29, which separates the hydrocarbon top flow from the aqueous flow in line 34, which is then fed back into the product aqueous flow and sent to the water stripper column 30 in line 36. The top product flow in line 40, containing olefins discharged from the product separation column, is sent to the compression section 80.

[0071] The product water stream in line 36 contains dilute hydrocarbon oxygenates such as DME, methanol, acetaldehyde, acetone, and MEK (methyl ethyl ketone). The water stripper column 30 separates or strips the oxygenates, dividing the mixture into a methanol and oxygenate-rich stream in the top line 49, which is concentrated with methanol and at least one other oxygenate, and a water-rich stream in the bottom line 46.

[0072] In one embodiment, the temperature of the water stripper tower 30 may be the temperature at the bottom of the water stripper tower, and the pressure may be 75 kPa gauge (11 psig) to 345 kPa (50 psig) at the top of the water stripper tower.

[0073] The product olefin stream in the product top line 40 carries the valuable olefin product that needs to be recovered. The compression section 80 increases the pressure of the product olefin stream required for downstream processing, as used in conventional light olefin recovery units. The compression section 80 may include a first knockout drum 82 that separates the product olefin stream into a first concentrated olefin stream pressurized at a temperature of 40°C (104°F) to 60°C (140°F) and a pressure of 193 kPa(g) (28 psig) to 262 kPa(g) (38 psig) at the top line 83 and an oxygenated first aqueous stream at the bottom line 84. The concentrated olefin stream in the top line 83 may be fed to a compressor 85, cooled, and directed to a second knockout drum 86. The aqueous stream in the bottom line 84 is pumped via a manifold line 76 to a return line 32 that returns the aqueous stream, along with the product aqueous stream in the bottom line 36 of the product separation column, back to the water stripper column 30.

[0074] The compression section 80 may include a second knockout drum 86 that separates a pressurized first olefin-rich flow into a second pressurized olefin-rich flow at a top line 87 at a pressure of 330 kPa(g)(48 psig) to 400 kPa(g)(58 psig) and a temperature of 27°C(80°F) to 54°C(130°F) and into an oxygenated second aqueous flow at a bottom line 88. The second olefin-rich flow at the top line 87 may be supplied to a compressor 89, cooled, and directed to a third knockout drum 90. The aqueous flow in the bottom line 88 is pumped to a return line 32 via a manifold line 76 that returns the aqueous flow to the water stripper column 30, along with the product aqueous flow in the bottom line 36 of the product separation column.

[0075] The compression section 80 may include a third knockout drum 90 that separates a pressurized second olefin-enriched stream into a third pressurized olefin-enriched stream in the top line 91 and an oxygenated third aqueous stream in the bottom line 92. The third olefin-enriched stream in the top line 91 may be supplied as feed to the oxygenated absorption column 50. The aqueous stream in the bottom line 92 is pumped through the manifold line 76 to the return line 32, from where the water stream is returned to the water stripper column 30 along with the product water stream in the bottom line 36 of the separation column.

[0076] Suitable compressor types may include centrifugal, positive displacement, piston, diaphragm, and screw types. In one embodiment, compressors 85 and 89 in compressor section 80 are centrifugal compressors. The final discharge pressure can range from 1 MPa gauge pressure (145 psig) to 2 MPa gauge pressure (290 psig). The compressor discharge can be cooled to ambient temperature using conventional heat transfer methods.

[0077] As shown in Figure 2, in a preferred embodiment, at least a portion of the compression product flow passing through the top line 91 is brought into contact in the oxygenation absorption tower 50 with a cooled, dilute, water-free flow directly taken in from the product separation tower 24, under conditions effective for absorbing the oxygenation without prior removal of it. The contact in the oxygenation absorption tower 50 generates a concentrated absorbed olefin flow in the top line 54 and a concentrated absorbed water flow containing a certain amount of effluent oxygenation in the bottom line 52. The operating conditions for the oxygenation absorption tower may be a bottom temperature range of 30°C (86°F) to 60°C (140°F) and a top pressure range of 700 kPa (101 psig) to 1 MPa (145 psig).

[0078] The absorption olefin-containing flow in the top line 54 can be supplied as raw material to the absorption column separator 60, where the gaseous olefin flow is taken in from the top line 61 to the third compressor 62, while water and oxygenates are taken in from the bottom line 59 to the manifold line 76. The gaseous olefin flow in line 61 is compressed in the third compressor and merges with the flow in line 71 via line 63, where it is cooled and partially condensed in the heat exchanger 64 and supplied as raw material to the stripper separator 66 via line 65. The stripper separator separates the aqueous flow containing oxygenates in the boot in line 67 that supplies raw material to the manifold line 76, the light olefin vapor flow in the top line 68 containing C3-olefins, and the heavy olefin liquid flow in line 69 containing C4+ olefins. The heavy olefin liquid flow in line 69 is stripped in the DME stripper column 70 to remove C3- and lower vapors in the stripper column top line 71 from the heavy olefin liquid flow in the stripper column bottom line 168. Most of the oxygenated material is stripped into the stripper column top line 71 and separated upon cooling and recirculation to the stripper separator 66. The bottom flow exiting the DME stripper column 70 can be sent to hydrogenation through line 168. This flow mainly contains C4+ olefins but also contains diolefins that interfere with oligomerization catalysts requiring selective hydrogenation. The stripper separator may operate at temperatures from 30°C (86°F) to 60°C (140°F) and pressures from 1.7 MPa (g) (250 psig) to 2.1 MPa (g) (300 psig). The light olefin vapor flow in the top line 68 is washed in the caustic alkali scrubber column 73 by countercurrent contact with the caustic alkali solution in line 342, and acidic gases such as carbon dioxide are absorbed from the light olefin vapor. The vapor flows out of the caustic alkali scrubber 73 to the top line 74. The concentrated caustic alkali solution containing the acidic gases flows out of the scrubber 73 to line 44.

[0079] The washed light olefin vapor in the top line 74 is cooled with a propylene refrigerant in the cryogenic cooler 75 to liquefy a portion of the light olefin flow, which is separated in the dry separator 46a to provide an aqueous flow from the boot that is taken into the manifold line 76, a vaporized light olefin flow containing C2- hydrocarbons and gas in the top line 77, and a liquid light olefin flow in the bottom line 78 containing C3+ hydrocarbons. The vaporized light olefin flow in the top line 77 is dried in the dryer 79a to provide a vaporized product olefin flow in line 112. The liquid light olefin flow in the bottom line 78 is dried in the dryer 79b to provide a liquid product olefin flow in line 114. The product olefin flows in lines 112 and 114 are processed in the oligomerization feedstock preparation section.

[0080] Oligomerization Olefin oligomerization is a process that allows smaller olefins to be oligomerized into larger olefins. More specifically, olefins, including oligomerized olefins, can be converted into distillates containing products in the jet fuel and diesel range. The oligomerized distillates can be saturated for use as transport fuels.

[0081] Jet fuel is one of the few petroleum fuels that cannot be easily replaced by electric motor systems because the fuel for airplanes requires a high energy output that electric motors cannot provide. Jet fuel has an endpoint boiling point specification of less than 300°C when using ASTM D86. Significant incentives are currently available for green jet fuel in certain regions.

[0082] The oligomerization section 310 is shown in Figure 3. The input olefin flow and the recirculated olefin flow are oligomerized by a first-stage oligomerization catalyst to produce a first-stage oligomerized olefin flow. The first-stage oligomerized olefin flow is oligomerized by a second-stage oligomerization catalyst to provide a second-stage oligomerized flow. A dealkane column can be used to remove light alkanes that are inert and accumulate in the recirculation loop. The second-stage oligomerized product flow is hydrogenated to produce fuel.

[0083] Referring to the oligomerization section 310 in Figure 3, the input olefin stream in line 312 is supplied to the oligomerization section 310. The input olefin stream may substantially contain ethylene and propylene. The input olefin stream may mainly contain ethylene and / or propylene. In one embodiment, the input olefin stream may contain at least 95 mol% ethylene and / or propylene. The input olefin stream in line 312 may be referred to as a light olefin stream. Additional olefin species having carbon numbers in the range of C4 to C6 can be expected to be present in the input stream. The light olefin stream may be supplied by dehydration of ethanol or from an MTO unit. The input olefin stream may have a temperature of 60°C (140°F) to 150°C (302°F), preferably 80°C (176°F) to 100°C (212°F), and a pressure of 3.5 MPag (500 psig), preferably 5.6 MPag (800 psig) to 8.4 MPag (1200 psig).

[0084] The input olefin stream may first be brought into contact with the first-stage oligomerization catalyst to oligomerize ethylene and propylene, and then into contact with the second oligomerization catalyst to oligomerize the unconverted ethylene and propylene from the first-stage oligomerization. Alternatively, the olefin stream may first be brought into contact with the second-stage oligomerization catalyst to oligomerize ethylene and propylene, and then into contact with the first-stage oligomerization catalyst to oligomerize the oligomerized ethylene and propylene.

[0085] Oligomerization reactions generate significant heat. For example, the dimerization of ethylene can generate 612 kcal / kg (1100 BTU / lb) of heat. Consequently, this large heat generation must be managed. Therefore, the input olefin flow in line 312 can be divided into multiple olefin flows. In Figure 3, the input olefin flow is divided into two separate flows: the first input olefin flow goes into the first input olefin line 312a, and the second input olefin flow goes into the second input olefin line 312b. More or fewer separate olefin flows can also be used. Up to six input olefin flows are easily conceivable. The input olefin flow in line 312 can be divided into multiple olefin flows with equal distributions. Alternatively, the input olefin flow in line 312 may be divided into unequal flows. For example, the input olefin flow may be divided into flows with decreasing flow rates, where the input olefin flow to a preceding reactor has a higher flow rate than the input olefin flow to a subsequent reactor. In one embodiment, the input olefin flow is divided into two flows of equal flow rates, each containing 50% by volume of the input olefin flow. In another embodiment, the first input olefin flow in the first input olefin line 312a may contain 70-90% by volume of the input olefin flow, and the second input olefin flow in the second olefin line 312b may contain 10-30% by volume of the input olefin flow.

[0086] To manage the heat generation, the input olefin stream may be diluted with a diluent stream to provide a diluted olefin stream for heat absorption. The diluent stream may include a paraffin stream in the diluent line 314. The diluent stream in the diluent line 314 may be added to the first-stage input olefin stream in the first-stage input olefin line 312a before being introduced into the first-stage oligomerization reactor 322. Preferably, the diluent stream is added to the first-stage input olefin stream in line 312a after the input olefin stream in line 312 has been divided into multiple olefin streams to create a first diluted olefin input stream in line 316a, so that the diluent stream passes through the entire first-stage oligomerization reaction. Alternatively, the diluent stream may be divided into multiple streams, and each diluent stream may be added to the corresponding input olefin stream. The volumetric flow rate of the diluent stream may be 2 to 8 times, preferably 3 to 6 times, the volumetric flow rate of the input olefin stream in the input olefin line 312.

[0087] The recirculating olefin stream in the recirculation line 326, which contains C4-C8 olefins, can be mixed with the input olefin stream and oligomerized in the first-stage oligomerization reactor 322. In one embodiment, the recirculating olefin stream in line 326 is divided into several recirculating olefin streams 326a-326d. The recirculating olefin stream in the first recirculating olefin line 326a can be mixed with the first input olefin stream in line 312a and fed into the first-stage oligomerization reactor 322. In a further embodiment, the first recirculating olefin stream in the first recirculating olefin line 326a is mixed with the first input olefin stream in line 312a and the diluent stream in line 314 to provide a diluted first input olefin stream in line 316a.

[0088] The first diluted olefin stream may contain 35% by weight or less of olefin, preferably 30% by weight or less of olefin, and more preferably 20% by weight or less of olefin. In embodiments, the first diluted olefin stream contains 10 to 30% by weight of C2 to C8 olefins. The first diluted olefin stream may contain 30% by weight or less of ethylene, preferably 25% by weight or less of ethylene, and more preferably 20% by weight or less of ethylene. In embodiments, the first diluted olefin stream contains 10 to 20% by weight of propylene. The first diluted olefin stream may contain 30% by weight or less of propylene, preferably 25% by weight or less of propylene, and more preferably 20% by weight or less of propylene. In embodiments, the first diluted olefin stream contains 10 to 20% by weight of propylene.

[0089] The first-stage oligomerization reactor 322 may comprise a series of first-stage oligomerization catalyst beds 322a, 322b, 322c, and 322d for feeding olefin feed streams 312a, 312b, 312c, and 312d, respectively. The first-stage oligomerization reactor 322 preferably includes four fixed first-stage oligomerization catalyst beds 322a, 322b, 322c, and 322d. Each first-stage oligomerization catalyst bed 322a, 322b, 322c, and 322d may be located in a dedicated first-stage oligomerization reactor, or multiple first-stage oligomerization catalyst beds may be located in two or more separate first-stage oligomerization reaction vessels. Up to six first-stage oligomerization catalyst beds are readily conceivable. In Figure 3, two first-stage oligomerization reaction vessels 321a and 321b are used.

[0090] A parallel first-stage oligomerization reactor may be used when the first-stage oligomerization reactor 322 is deactivated, during which time the first-stage oligomerization reactor 322 is regenerated in situ by the combustion of coke from the catalyst. In another embodiment, each first-stage oligomerization reactor may include a reed reactor, a lag reactor, and a pre-reactor to facilitate regeneration. Figure 3 shows only two reaction vessels 321a and 321b.

[0091] The diluted first input olefin stream in line 316a is cooled in the first input cooler 318a and supplied to line 320a, where it can be introduced into the first bed 322a of the first-stage oligomerization catalyst in the first-stage oligomerization reaction vessel 321a of the first-stage oligomerization reactor 322. The cooled diluted first input olefin stream in line 320a can be introduced at a temperature of 180°C (356°F) to 260°C (500°F) and a pressure of 3.5 MPag (500 psig) to 8.4 MPag (1200 psig). The input cooler 318a may be equipped with a steam generator.

[0092] The diluted first input olefin stream can preferably be introduced into the first first-stage catalyst bed 322a in line 320a in a downward flow operation. However, an upward flow operation may be preferable. As oligomerization of ethylene, propylene, and recirculated olefins occurs in the first first-stage oligomerization catalyst bed 322a, heat is generated due to the high exothermic nature of the olefin oligomerization reaction. Once the first input olefin stream is oligomerized, a first oligomerized effluent flow is generated in the first oligomerized effluent line 324a at an elevated outlet temperature despite cooling and dilution. The elevated outlet temperature is limited to 150°C (302°F) to 250°C (482°F).

[0093] The second input olefin stream in line 312b may be mixed with the second recirculated olefin stream in the second recirculated olefin line 326b and the first oligomerized effluent stream in the first oligomerized effluent line 324a, which has been removed from the first first-stage oligomerization catalyst bed 322a of the first first-stage reactor 321a, to provide a mixed second input olefin stream in line 316b. The first oligomerized effluent stream in line 324a includes a diluent stream from diluent line 314 that has been added to the first olefin input stream in line 312a. The second diluent olefin stream may contain 35% by weight or less of C2-C8 olefins, preferably 25% by weight or less of C2-C8 olefins, and preferably 20% by weight or less of ethylene. The second input olefin stream may contain 30% by weight or less of ethylene, preferably 25% by weight or less of ethylene, and preferably 20% by weight or less of ethylene. The second input olefin stream may contain 30% by weight or less of propylene, preferably 25% by weight or less of propylene, and more preferably 20% by weight or less of propylene. The second input olefin stream in line 316b is cooled in a second input cooler 318b, which may be located outside the first first-stage oligomerization reactor 321a, and the cooled second input olefin stream is supplied into line 320b and fed into the second bed 322b of the first-stage oligomerization catalyst in the first first-stage oligomerization reactor 321a. The input cooler 318b may include a steam generator.

[0094] The second input olefin stream in line 320b can be introduced at a temperature of 180°C (356°F) to 230°C (446°F) and a pressure of 3.5 MPag (500 psig) to 8.4 MPag (1200 psig). The second input olefin stream contains a diluent and olefins from the first oligomerized stream. The olefins from the first oligomerized olefin stream are oligomerized in the second catalyst bed 322b. The oligomerization of ethylene, propylene, recycled olefins, and oligomers in the second olefin stream in the second bed 322b of the first-stage oligomerized catalyst generates a second oligomerized olefin effluent stream in the second oligomerized effluent line 324b at a high outlet temperature. The elevated outlet temperature can be limited to a temperature 30°C (54°F) to 50°C (90°F) higher than the inlet temperature to the catalyst bed 322b.

[0095] The second oligomerization effluent in line 324b, removed from the second first-stage oligomerization catalyst bed 322b of the first first-stage reaction vessel 321a, can be mixed with the third recirculating olefin flow in the third recirculating olefin line 326c to provide the first recirculating olefin input flow in line 316c. The input olefin flow in line 312 is not directly added to the first recirculating olefin input flow in line 316c. Alternatively, the second oligomerization effluent may be added to a portion of the input olefin flow in line 312, together with the first recirculating olefin input flow in line 316c. The second oligomerization effluent in line 324b includes the diluent flow from the added diluent line 314 in the first input olefin flow 312a. The first recirculating olefin input flow may contain 30% by weight or less of ethylene, preferably 25% by weight or less of ethylene, and more preferably 20% by weight or less of ethylene. The first recirculated olefin feed stream may contain 30% by weight or less of propylene, preferably 25% by weight or less of propylene, and more preferably 20% by weight or less of propylene. The first recirculated olefin feed stream may contain 30% by weight or less of C2-C8 olefins, preferably 25% by weight or less of C2-C8 olefins, and more preferably 20% by weight or less of C2-C8 olefins. The first recirculated olefin feed stream in line 316c is cooled in a third feed cooler 318c, which may be located outside the oligomerization reactor 322, to provide a cooled first recirculated olefin feed stream in line 320c, which can be fed into the third bed 322c of the first-stage oligomerization catalyst in the first-stage oligomerization reactor 322. In embodiments, the third bed 322c of the first-stage oligomerization catalyst is provided in the second first-stage oligomerization reaction vessel 321b. The feed cooler 318c may include a steam generator.

[0096] The first cooled and diluted olefin stream in line 320c can be introduced at a temperature of 180°C (356°F) to 230°C (446°F) and a pressure of 3.5 MPag (500 psig) to 8.4 MPag (1200 psig). The first recirculated olefin feed stream contains a diluent and olefins from the second oligomerized olefin stream and the third recirculated olefin stream. The olefins are oligomerized in the third catalyst bed 322c. The oligomerization of ethylene and propylene in the third bed 322c of the first-stage oligomerized catalyst, as well as the oligomerization of oligomers in the first recirculated olefin feed stream, generates a third oligomerized effluent stream in the third oligomerized effluent line 324c at an elevated outlet temperature. In this embodiment, the third oligomerized effluent logistics is the second-to-last oligomerized effluent logistics, and the third oligomerized effluent line 324c is the second-to-last oligomerized effluent line 324c. The elevated outlet temperature is limited to a temperature 30°C (54°F) to 50°C (90°F) higher than the inlet temperature to the catalyst bed 322c.

[0097] The third oligomerization effluent in line 324c, removed from the second first-stage oligomerization reaction vessel 321b of the first-stage oligomerization reactor 322, can be mixed with the fourth recirculated olefin stream in line 326d to provide the second recirculated olefin feed stream in line 316d. The third oligomerization effluent in line 324c includes a diluent stream from line 314 that is added to the first olefin stream in diluent line 312a. The feed olefin stream in line 312 is not directly added to the second recirculated olefin feed stream in line 316d. In embodiments, the third oligomerization effluent in line 324c can also be mixed with the olefin feed stream from olefin feed line 322 and oligomerized. The second recirculated olefin feed stream may contain 35% by weight or less of C2-C8 olefins, preferably 30% by weight or less of C2-C8 olefins, and preferably 25% by weight or less of C2-C8 olefins. The second recirculated olefin feed stream may contain 30% by weight or less of ethylene, preferably 25% by weight or less of ethylene, and more preferably 20% by weight or less of ethylene. The second recirculated olefin feed stream may contain 30% by weight or less of propylene, preferably 25% by weight or less of propylene, and more preferably 20% by weight or less of propylene. The second recirculated olefin feed stream in line 316d is cooled in a fourth feed cooler 318d, which may be located outside the second vessel 321b of the first-stage oligomerization reactor 322, to provide a cooled second recirculated olefin feed stream in line 320d, which can be fed into the fourth bed 322d of the first-stage oligomerization catalyst in the second vessel of the first-stage oligomerization reactor 322. The feed cooler 318d may include a steam generator.

[0098] The cooled second recirculating olefin feed stream in line 320d can be fed at a temperature of 180°C (356°F) to 230°C (446°F) and a pressure of 3.5 MPa (g) (500 psig) to 8.4 MPa (g) (1200 psig). The cooled second recirculating olefin feed stream in line 320d contains a diluent, and olefins from the third or second-to-last oligomerized effluent stream, and C4-C8 olefins from the fourth recirculating olefin stream. The olefins are oligomerized on the fourth catalyst bed 322d. The oligomerization of ethylene and propylene in the second recirculating olefin feed stream in the fourth bed 322d of the first-stage oligomerization catalyst generates the fourth oligomerization stream in the fourth oligomerized effluent line 324d at an elevated outlet temperature. The increased outlet temperature is limited to a temperature 30°C (54°F) to 50°C (90°F) higher than the inlet temperature to the catalyst bed 322d.

[0099] The fourth oligomerized effluent in line 324d exits the second reaction vessel 321b of the first-stage oligomerization reactor 322. In this embodiment, the fourth oligomerized effluent in line 324d is the final oligomerized effluent, and the fourth oligomerized effluent line 324d is the final oligomerized effluent line 324d.

[0100] The first-stage oligomerization reaction occurs mainly in the liquid phase or gas-liquid mixed phase at an LHSV of 0.5 to 10 per hour based on olefins. The inventors have found that across the first-stage oligomerization catalyst bed, typically 30 to 50% by weight of ethylene in the olefin stream is converted to higher olefins. The ethylene is first dimerized on the catalyst to butene. The majority of propylene and butene in the olefin stream introduced into the first-stage oligomerization catalyst bed are oligomerized. In embodiments, at least 99 mol% of propylene and butene in the olefin stream are oligomerized.

[0101] The first-stage oligomerization catalyst may include a zeolite catalyst. The first-stage oligomerization catalyst can be considered as a solid acid catalyst. The zeolite may constitute 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, MOR, OFF, MTW, TON, MTT, AFO, ATO, and AEL. The three-letter codes indicating the zeotype are as defined by the Structure Commission of the International Zeolite Association and are managed at http: / / www.iza-structure.org / databases. UZM-8 is as described in U.S. Patent No. 6,756,030. In a preferred embodiment, the first-stage oligomerization catalyst may include a zeolite having a skeleton with a 10-membered ring pore structure. Examples of suitable zeolites having a 10-membered ring pore structure include TON, MTT, MFI, MEL, AFO, AEL, EUO, and FER. In a further preferred embodiment, the first-stage oligomerization catalyst, which includes a zeolite having a 10-membered ring pore structure, may include a one-dimensional pore structure. The one-dimensional pore structure represents a zeolite containing non-crossing pores substantially parallel to one of the crystal axes. The pores preferably extend through the zeolite crystal. A suitable example of a zeolite having a 10-membered ring one-dimensional pore structure is MTT. In a further embodiment, the first-stage oligomerization catalyst includes an MTT zeolite.

[0102] The first-stage oligomerization catalyst can be formed by combining a zeolite with a binder, and then forming the catalyst into a pellet. The pellet may be optionally treated with a phosphorus reagent to produce a zeolite having 0.5 to 15% by weight of phosphorus in 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.

[0103] 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.

[0104] A suitable first-stage oligomerization catalyst is prepared by mixing proportional volume amounts of zeolite and alumina to achieve a desired zeolite-alumina ratio. In embodiments, the MTT content may be 5-85, for example, 20-82% by weight of MTT zeolite, and the remaining alumina powder provides a suitably supported catalyst. Silica supports are also considered.

[0105] A monobasic acid such as nitric acid or formic acid may be added to the mixture in an aqueous solution to dissolve the alumina in the binder. Additional water may be added to the mixture to provide sufficient wettability to form a dough with sufficient viscosity for extrusion or spray drying. Extrusion aids such as cellulose ether powder may also be added. Preferred extrusion aids are available from The Dow Chemical Company under the trademark Methocel.

[0106] The paste or dough may be prepared in the form of molded particles, a preferred method of which involves extruding the dough through a die having an opening of a desired size and shape, then dividing the extruded material into extruders of a desired length and drying 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 with amines or the like.

[0107] 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 can be as small as about 40 μm. However, it is usually 0.635 mm (0.25 inches) to 12.7 mm (0.5 inches), preferably 0.79 mm (1 / 32 inch) to 6.35 mm (0.25 inches), and most preferably 0.06 mm (1 / 24 inch) to 4.23 mm (1 / 6 inch).

[0108] In one exemplary embodiment, an MTT-type zeolite catalyst, arranged on a high-purity pseudo-boehmite alumina substrate in a ratio of 90 / 10 to 20 / 80, preferably 20 / 80 to 50 / 50, is provided to the catalyst bed, or more often, to the first-stage oligomerization reactor 322.

[0109] The first-stage oligomerization catalyst can be regenerated when deactivated. Preferred regeneration conditions include, for example, exposing the first-stage oligomerization catalyst to hot air at 400°C to 500°C for 3 hours in situ. To promote regeneration without downtime, a rocking bed apparatus may be used in conjunction with an alternative first-stage oligomerization reactor. The regeneration gas stream may be introduced into the first-stage oligomerization reactor 322 that requires regeneration. The regeneration gas may contain air with increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to that of a fresh catalyst.

[0110] Zeolite catalysts are advantageous as first-stage oligomerization catalysts. Zeolite catalysts have relatively low susceptibility to oxygenate contamination. As a result, when produced from an ethanol dehydration process, the degree of oxygenate removal required for the olefin feedstock in line 312 is reduced.

[0111] The final first-stage oligomerization flow in the last first-stage oligomerization effluent line 324d has increased concentrations of ethylene oligomer and propylene oligomer compared to the input olefin flow in line 312. The final first-stage oligomerization flow in the last first-stage oligomerization effluent line 324d is cooled by steam generation in the steam generator 318e or by other heat exchange, further cooled by heat exchange with the second-stage oligomerization flow in line 334, and possibly further cooled by an air cooler to become the input first-stage oligomerization flow, which is then fed into the second-stage oligomerization reactor 332 in the second-stage oligomerization input line 328. To achieve the most desirable olefin product, the second-stage oligomerization reactor 332 operates at a temperature of 80°C (176°F) to 180°C (356°F). The second-stage oligomerization reactor 332 is operated at a pressure of 2.1 MPa (300 psig) to 7.6 MPa (1100 psig), more preferably 3.5 MPa (500 psig) to 6.9 MPa (1000 psig).

[0112] The second-stage oligomerization reactor 332 may be downstream-communicated with the first-stage oligomerization reactor 322. The second-stage oligomerization reactor 332 preferably operates in a downward flow manner. However, an upward flow manner may be preferable. The second-stage oligomerization feedstream is brought into contact with the second-stage oligomerization catalyst, dimerizing and trimerizing the unconverted ethylene from the first-stage oligomerization reactor 322, while also dimerizing, trimerizing, and tetramerizing the higher olefins to provide distillate-range olefins. With respect to the second-stage oligomerization reactor 332, process conditions are selected to increase the proportion of jet-range olefins that yield desirable jet-range hydrocarbon products when hydrogenated in subsequent steps. The majority of the unconverted ethylene from the first-stage oligomerization reactor 322 is dimerized, trimerized, and tetramerized. In embodiments, at least 99% by weight of the ethylene in the second-stage oligomerization feedstream is converted, usually to butene.

[0113] The second-stage oligomerization reactor 332 may include a first reaction vessel 331a containing a first bed 332a of the second-stage oligomerization catalyst and a second reaction vessel 331b containing a second bed 332b of the second-stage oligomerization catalyst. The first second-stage oligomerization flow is discharged from the first second-stage reaction vessel 331a, cooled, and introduced into the second second-stage reaction vessel 331b. The second-stage oligomerization flow, having a higher average number of carbon atoms than the first-stage oligomerization flow introduced in line 328, exits the second-stage oligomerization reactor 332 in line 334.

[0114] The second-stage oligomerization catalyst is preferably an amorphous silica-alumina base having a metal from either Group VIII and / or Group VIB of the periodic table using Chemical Abstracts Service notation. In one embodiment, the catalyst has a Group VIII metal promoted by a Group VIB metal. Typically, since silica and alumina are present only in the base, the ratio of silica to alumina is the same for the catalyst as for the base. The metal can be impregnated onto the silica-alumina base or ion-exchanged in the silica-alumina base. Co-mulling is also considered. The catalyst of the present invention may have a low-temperature acidity ratio of at least 0.15, preferably 0.2, and more preferably greater than 0.25, as measured by Ammonia Temperature Programmed Desorption (ammonia TPD) as described below. In addition, a suitable catalyst has a surface area of ​​50-400 m² / g as measured by nitrogen BET.

[0115] A preferred second-stage oligomerization catalyst comprises an amorphous silica-alumina support. One component of the catalyst support used in this disclosure is alumina. The alumina may be any of various hydrated aluminum oxides or alumina gels, such as alpha-alumina monohydrate with a boehmite or pseudoboehmite structure, alpha-alumina trihydrate with a gibbsite structure, or beta-alumina trihydrate with a bayerite structure. A particularly preferred alumina is available from Sasol North America Alumina Product Group under the trademark CATAPAL. This material is an extremely high-purity alpha-alumina monohydrate (pseudoboehmite) and has been shown to yield high-purity gamma-alumina after calcination at high temperatures. Another component of the catalyst support is amorphous silica-alumina. A preferred silica-alumina with a silica-to-alumina ratio of 2.6 is available, for example, from CCIC, a subsidiary of JGC in Japan.

[0116] Another component used in the preparation of the second-stage oligomerization catalyst used in this disclosure is a surfactant. The surfactant is preferably mixed with the alumina and silica-alumina powder. 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 fully performed its function according to this disclosure. Any suitable surfactant can be used according to this disclosure. 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.

[0117] 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, 75-99% by weight of amorphous silica-alumina with a silica:alumina ratio of 2.6 and 10-20% by weight of alumina powder provide a suitable carrier. In another embodiment, other ratios of amorphous silica-alumina to alumina may be preferred.

[0118] 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.

[0119] 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.

[0120] The paste or dough can be prepared in the form of molded particles. A preferred method involves extruding a dough mixture of alumina, silica-alumina, a surfactant, and water through a die having an opening of the desired size and shape, then dividing the extruded material into extruders of the desired length and drying them. A further calcination step may be used to provide additional strength to the extruders. Generally, calcination is carried out in a stream of dry air at a temperature of 260°C (500°F) to 815°C (1500°F).

[0121] 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 can be as small as about 40 μm. However, it is usually 0.635 mm (0.25 inches) to 12.7 mm (0.5 inches), preferably 0.79 mm (1 / 32 inch) to 6.35 mm (0.25 inches), and most preferably 0.06 mm (1 / 24 inch) to 4.23 mm (1 / 6 inch).

[0122] Typical properties of the amorphous silica-alumina supports used herein are total pore volume, average pore diameter, and surface area large enough to provide substantial space and area for depositing active metal components. The total pore volume of the support, as measured by conventional mercury porosimeters, is typically 0.2–2.0 cc / gram, preferably 0.25–1.0 cc / gram, and most preferably 0.3–0.9 cc / gram. Typically, the amount of pore volume of the support in pores with a diameter greater than 100 angstroms is less than 0.1 cc / gram, preferably less than 0.08 cc / gram, and most preferably less than 0.05 cc / gram. The surface area, as measured by BET, is typically 50 m² / gram or more, for example 200 m² / gram or more, preferably at least 250 m² / gram, and most preferably 300 m² / gram to 400 m² / gram.

[0123] To prepare the second-stage oligomerization catalyst, the carrier material is compounded with one or more precursors of at least one metal component from Group VIII or Group VIB of the periodic table, such as by single or multiple impregnation of calcined amorphous refractory oxide carrier particles. The Group VIII metal, preferably nickel, must be present at a concentration of 0.5 to 15% by weight, and the Group VIB metal, preferably tungsten, must be present at a concentration of 0 to 12% by weight. Impregnation can be achieved by any method known in the art, for example, by spray impregnation, in which a solution containing the metal precursor in a dissolved form is sprayed onto the carrier particles. Another method is a multi-dip procedure, in which the carrier material is repeatedly brought into contact with the impregnation solution with or without intermittent drying. Yet another method includes immersing the carrier in a large volume of impregnation solution or circulating the carrier therein, and yet another method is a pore volume or pore saturation technique in which the carrier particles are introduced into an impregnation solution of just enough volume to fill the pores of the carrier. In some cases, the pore saturation technique may be modified to utilize an impregnation solution having a volume between 10 percent less and 10 percent more than the volume to which the pores are just filled.

[0124] When the active metal precursor is incorporated by impregnation, the metal is converted to its respective oxide form by subsequent or second calcination at a high temperature, for example, 399°C (750°F) to 760°C (1400°F). In some cases, calcination may be performed after each impregnation of the individual active metals. Subsequent calcination yields a catalyst containing the active metals in their respective oxide forms.

[0125] A preferred second-stage oligomerization catalyst of this disclosure has an amorphous silica-alumina base impregnated with 0.5–15 wt% nickel, in the form of a 3.175 mm (0.125 inch) extruded material with a density of 0.45–0.65 g / mL. It is also intended that the metal may be incorporated onto the support by other methods such as ion exchange and co-miscibility.

[0126] The second-stage oligomerization catalyst can be regenerated when deactivated. Suitable regeneration conditions include, for example, exposing the catalyst to hot air at 400-500°C for 3 hours in situ. To facilitate regeneration without downtime, a rocking bed apparatus may be used in conjunction with an alternative second-stage oligomerization reactor. The regeneration gas may contain air with increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to those of a fresh catalyst.

[0127] The second-stage oligomerization reaction is also inherently exothermic. The final oligomerized olefin stream in line 324d is added to the first olefin stream in line 312a and includes a diluent stream from diluent line 314, which is carried through the first-stage oligomerization catalyst beds 322a-322d. The diluent stream is then transported in line 328 to the second-stage oligomerization reactor 332 to absorb the heat generated in the second-stage oligomerization reactor. A dedicated diluent line to the second-stage oligomerization reactor 332 may be considered to rapidly control the rise in heat generation, or if it is desired to cool only the second stage by cooling the second-stage oligomerization reactor 332.

[0128] When the oligomerization reaction is carried out according to the process conditions, a C4 olefin conversion rate of 95% or more, or 97% or more, is achieved. The second-step oligomerization stream obtained in line 334 contains multiple olefin products, which are hydrocarbons within the distillate range.

[0129] The oligomerized olefin flow in line 334, which has an increased C8+ olefin concentration compared to the input first-stage oligomerization flow in line 328, is heat-exchanged with the first-stage oligomerization flow in line 324d, its pressure is reduced, and then it is heat-exchanged with the olefin splitter bottom flow in line 30 before being supplied as raw material to the dealkane column 340. The oligomerized olefin flow in line 334 is at a temperature of 160°C (320°F) to 190°C (374°F) and a pressure of 3.9 MPa (gauge) (550 psig) to 7 MPa (gauge) (1000 psig). The olefin splitter bottom flow in line 330 can be transported to the hydrogenation section.

[0130] The inventors have found that light alkanes, such as ethane and / or propane, are generated in the first-stage oligomerization reactor 322 and / or the second-stage oligomerization reactor 332, and that these must be removed from the second-stage oligomerization reactor, particularly for fuel production to facilitate the recirculation of light olefins to the first-stage oligomerization reactor 322. The light alkanes are inert and accumulate in the recirculation loop. Therefore, the second-stage oligomerization stream in line 334 is dealkaneted by fractional distillation in the dealkane column 340 to provide a light alkane stream and a dealkane stream. In one embodiment, the light alkane stream is an ethane stream, in which case the dealkane column 340 is a deethane column. In another embodiment, the light alkane stream is a propane stream, in which case the dealkane column 340 is a depropane column. The alkane stream may be a mixture of ethane and propane. The alkane stream can be used as a heating fuel in process 310.

[0131] In the dealkane column 340, light alkanes such as C3 hydrocarbons, preferably C2 hydrocarbons, are separated from the alkane top flow in the top line 342 of Figure 3, or from the dealkanization bottom flow in the bottom line 344, which contains C4+ hydrocarbons, preferably C3+ hydrocarbons. When the dealkane column 340 is operated as an ethane column, it can operate at a bottom temperature of 177°C (350°F) to 302°C (575°F) and a top pressure of 207 kPa (gauge pressure) (30 psig) to 690 kPa (gauge pressure) (100 psig). When the dealkane column 340 is operated as a propane column, it can operate at a bottom temperature of 194°C (381°F) to 333°C (630°F) and a top pressure of 207 kPa (gauge pressure) (30 psig) to 1.14 MPa (gauge pressure) (165 psig).

[0132] The alkane top flow in the top line 342 can be cooled and separated in the dealkane receiver 346, thereby providing dealkaneated off-gas in the off-gas line 347. In the off-gas line 347, this off-gas is cooled and, together with the net vapor flow in the receiver top line 368, can be supplied as a feedstock for further processing, such as being taken up as fuel gas in line 348. The condensate from the dealkane receiver 344 may be refluxed back to the dealkane column 340 in the dealkane top liquid line 349. In some embodiments, a portion of the condensate from the dealkane receiver 344 in line 349 can be taken up as recirculation in line 351 and sent to the first-stage oligomerization reactor in lines 372 and 326. The dealkane flow in the bottom line 344 may be divided into a reboiling flow in line 350, which is reboiled by a heat exchanger with the first high-temperature diesel flow in line 352, which is taken in from the jet fractionation bottom heat exchanger flow in the jet bottom heat exchanger line 374, and a net bottom flow in line 354, which is supplied directly to the olefin splitter column 360 without heating. The reboiled bottom flow in line 350 may be returned to the dealkane column 340 boiling to provide heating requirements. In another embodiment, the feedstock to the dealkane column 340 is not preheated by the olefin splitter bottom flow in line 330, but the feedstock to the olefin splitter column 360 in the net bottom line 354 is preheated by the olefin splitter bottom flow.

[0133] The dealkane flow in the net bottom line 354 of the dealkane process is divided by fractional distillation in the olefin splitter column 360 into a light olefin flow, possibly in the top line 362 of the olefin splitter column, and a heavy olefin flow, possibly in the bottom line 364 of the olefin splitter column. The top line of the olefin splitter column may be cooled to 66°C (150°F) to 93°C (200°F), and the resulting condensate is refluxed back into the olefin splitter column 360 from the olefin splitter receiver 366. The net vapor flow in the top line 368 of the receiver column from the olefin splitter receiver 366, along with the off-gas flow in the off-gas line 347, can be cooled and further processed, similar to the fuel gas in line 348. Light olefin condensates from the bottom of the olefin splitter receiver in line 370 can be divided into a reflux flow that is refluxed back into the column in line 371 and a light olefin recirculation flow in recirculation line 372 that can be recycled to the first-stage oligomerization reactor 322 or the second-stage oligomerization reactor 332. The light olefin flow in line 372 may contain 1 to 15% by weight of the light olefin flow in line 370. The light olefin flow in line 372 may contain 40 to 80% by weight of C4-C8 olefins. In embodiments, the light olefin flow in line 372 is flushed in a knockout drum 375 to remove vapor from the light olefin vapor flow, which is then transported to a hydrogenation section in top-of-column line 377, and the liquid recirculation olefin oligomer flow in line 326 may be recycled to the first-stage oligomerization reactor 322 to oligomerize C4-C8 olefins.

[0134] In the oligomerization process, the flow leaving the depropane unit is sent to a reactor, where it is converted to synthesis gas via steam reforming, partial oxidation, autothermal reforming, or dry reforming. The synthesis gas is then sent to a methanol synthesis unit. When partial oxidation or autothermal reforming is selected to produce synthesis gas, air or oxygen must also be supplied, although additional steam can be extracted from the reactor for use elsewhere in the facility. If steam reforming is selected to produce synthesis gas from ethane / propane, steam feedstock (which can be supplied entirely by a methanol-jet process) and energy input (which can be supplied partially by steam from a methanol-jet process) are also required.

[0135] Modification A synthesis gas flow containing carbon oxides and hydrogen can be produced by reforming the alkane top flow in the top line 342 of Figure 3 using steam in a steam reforming reactor, a self-thermal reforming reactor, or a dry reforming reactor.

[0136] Steam reforming By steam reforming the alkane top flow from the top line 342 in Figure 3 at high temperature using CO2 and steam, a synthesis gas mixture is created that can be directly supplied as a raw material for methanol synthesis. This is an endothermic process and therefore requires a considerable amount of external heat.

[0137] The alkane top flow 342 is mixed with 1-2 carbon equivalents of carbon dioxide to prevent the formation of elemental carbon. The mixture of alkane flow and carbon dioxide is then mixed with steam to produce a mixture containing 1 mole of hydrocarbon-derived carbon per 1-2 moles of water. The gas / steam mixture enters a steam reforming process unit with a commercially available Nicel catalyst. The conversion of hydrocarbons to carbon monoxide and hydrogen takes place at 700-900°C and a system pressure that can be 12-500 psia. A typical flow rate for such a reactor is 300 lbs / hour / cubic feet of catalyst. After water condensation, the product gas from steam reforming may be fed directly into a synthesis gas mixer, depending on the intended use of the synthesis gas, or carbon dioxide and other trace components may be removed first.

[0138] Self-thermal reforming Autothermal reforming combines both partial oxidation and steam reforming. Oxygen, steam, and CO2 from the hydrolysis unit react on a catalyst to produce a synthesis gas mixture suitable for conversion to methanol.

[0139] The alkane top flow from the top line 342 in Figure 3 reacts in the self-thermal reactor to produce a recirculated synthesis gas flow containing mostly CO and hydrogen. The heat from the exothermic partial oxidation helps drive the endothermic steam reforming process. The resulting synthesis gas can be recycled to the methanol synthesis unit.

[0140] Dry modification Dry reforming is an endothermic reaction between methane and carbon dioxide, producing carbon monoxide and hydrogen. The alkane top flow from the top line 342 in Figure 3 reacts in the dry reformer to produce synthesis gas. The resulting synthesis gas can be recycled to the methanol synthesis unit.

[0141] Partial oxidation Alternatively, a synthesis gas stream containing carbon oxides and hydrogen can be produced by partially oxidizing the alkane top stream from the top line 342 in Figure 3 using oxygen in a partial oxidation reactor.

[0142] Hydrogen recycling It has been found that hydrogen from a hydrogen generator, such as a reformer or electrolytic cell, can first be sent to a hydrocarbon hydrogenation reactor to produce renewable fuel. Excess hydrogen and other off-gases flashed off in the separation drum are then sent to a CO2 hydrogenation reactor to produce methanol. Excess hydrogen from the CO2 hydrogenation reactor is then compressed, purified (if necessary), and recycled to the hydrocarbon hydrogenation reactor, where it is mixed with fresh hydrogen from the hydrogen generator. Purge of hydrogen during recycling may be necessary. Hydrogen from a hydrogen generator can first be sent to a CO2 hydrogenation reactor to produce methanol. Excess hydrogen and any other off-gases flashed off in the separation drum are then sent to a hydrocarbon hydrogenation reactor to produce renewable fuel. Excess hydrogen from the hydrocarbon hydrogenation reactor is then compressed, purified (if necessary), and recycled to the CO2 hydrogenation reactor, where it is mixed with fresh H2 from the H2 generator. Purge of H2 during recycling may be necessary.

[0143] Figure 4 shows a typical flow scheme with hydrogen recirculation. The carbon dioxide stream 500 is sent to the methanol synthesis unit 502. The crude methanol stream 503 is then sent to the methanol purification unit 504, and the green feed methanol 530 is then sent to the methanol-to-olefin vaporization section 528. The dry methanol and recovered dimethyl ether stream 529 is then sent to the methanol-olefin reactor and regeneration unit 525, along with the product stream 524, which is then sent to the light olefin recovery process 522. The off-gas stream 523 is removed, and the dry liquid and vaporized olefin product stream 540 and the water recirculation or water removal stream 526 are shown. The stream 540 is then sent to the light olefin recovery process section 521 to produce the purified liquid and vaporized olefin stream 520, which is then sent to the oligomerization section 516, along with the off-gas 517, recirculated oil 513, and C9+ olefin stream 515. The C9+ olefin stream 515 is sent to the hydrogenation section 512, from which the off-gas stream 511, sustainable aviation fuel 510, and recoverable diesel and naphtha streams 509 emerge. Also illustrated is a recirculated hydrogen stream 514 from the hydrogenation section 512, which is sent to a light olefin recovery process for use in selective hydrogenation or other processes requiring hydrogen. In the lower left is a hydrogen production unit 508, which may be an electrolytic cell or reformer as described above. Hydrogen 507 from the methanol synthesis unit 502 can be recirculated and used in the hydrogenation unit 512.

[0144] Figure 5 shows the light olefin recovery portion of the process after fractional distillation and concentration. The dry olefin liquid 602 is pumped and combined with the dry olefin vapor stream 604 and enters the demethane feed chiller 606, which is cooled by a cooler 608. The cooled olefin is sent to the demethane column 610 with the off-gas 612, and then proceeds to the deethane column 614. A portion of the recycled hydrogen stream 616 is added after being separated from the hydrogenation unit and then proceeds to the acetylene converter 618. The olefin then passes through the water and carbon dioxide guard bed 620, and the purified olefin stream 622 is sent to the oligomerization section. The bottom stream from the deethane column 614 is sent to the DME washing column 624 and the DME washing water stripper column 626, so that the recovered DME can be sent to the methanol-olefin reactor through line 628. The top flow from the DME scrubbing tower, along with some C4+ olefins from the DME stripper, is sent to the selective hydrogenation unit 634. The selective hydrogenation unit 634 is in fluid downstream communication with the DME scrubbing tower 624. It is illustrated that some of the hydrogen 636 recirculated from the hydrogenation unit is added, and then the hydrogenation products are sent to the oxygenation removal unit 642 from which oxygenates 644 have been removed, and the purified olefin flow 640 is sent for oligomerization.

[0145] Figure 6 shows the hydrogenation section of the process, including where hydrogen can be recycled into and out of the process. The C9+ olefin flow 700 is sent to the hydrogenation reactor 704 along with the drag flow 702 from the olefin splitter. The bottom flow from the hydrogenation reactor is sent to the high-temperature separator 706 and then to the low-temperature separator 710, and the resulting hydrogen flow 708 can be sent as needed in the process shown in Figure 5 or to the methanol synthesis process. The remainder of the separated flow is sent to the flash stripper 712 and the jet fuel fractionation column 732 to produce a renewable naphtha flow 736, a sustainable aviation fuel jet fuel flow 734, and a renewable diesel flow 724. Reboiler sections 730 and 728 are also shown. The recycled oil is sent to the oligomerization section in line 722. [Examples]

[0146] The following examples cascade hydrogen from a recirculating flow in a dimerization and / or oligomerization unit to two different units within the plant's Light Olefin Recovery Process (LORP) section, referred to as LORP Unit 1 and LORP Unit 2 in the examples. To address the requirements of the units being simulated, a basic case simulation was developed using appropriate component and fluid packages.

[0147] Example 1 The cascade hydrogen from the recirculated flow was compared to a baseline case utilizing fresh, pure hydrogen. In the baseline case, fresh, pure hydrogen was sent to the dimerization and / or oligomerization unit, as well as to two separate units in the Light Olefin Recovery Process (LORP), LORP Unit 1 and LORP Unit 2. Fresh, pure hydrogen is typically 99.9 mol% or higher.

[0148] The dimerization and / or oligomerization unit also has a hydrogen recirculation flow that recirculates slightly impure hydrogen (about 97% hydrogen) back into the unit, and also has a purge associated with it of the same composition to prevent the accumulation of impurities in the oligomerization / dimerization hydrogen loop. One possible location for this purge flow was the fuel gas header in the plant. The cascade hydrogen network delivered fresh, pure hydrogen to the dimerization and / or oligomerization unit. However, it replaced the fresh, pure hydrogen to LORP unit 1 and LORP unit 2 with recirculated hydrogen (about 97% H2) from the dimerization and / or oligomerization unit. In this embodiment, the recirculation hydrogen purge rate was kept constant. Table 1 below shows the mass balance for both the basic cascade and the hydrogen cascade in this embodiment.

[0149] [Table 1]

[0150] The hydrogen pressure required for LORP Unit 1 and LORP Unit 2 was lower than the hydrogen pressure required for the oligomerization / dimerization unit. Table 2 below shows the hydrogen pressure required for the units.

[0151] [Table 2]

[0152] In the basic case, hydrogen is compressed to 630 psig for all units and then reduced to the required pressure for LORP Unit 1 and LORP Unit 2. In the cascade case, all hydrogen is pressurized to 630 psig and sent to the oligomerization / dimerization units, so it is utilized at 630 psig. This minimizes wasted losses from compression by reducing the load on the recirculating hydrogen compressor by 3% while maintaining the same load on the fresh, pure hydrogen compressor.

[0153] Example 2 Example 2 has the exact same basis as Example 1, but instead of maintaining a constant purge gas rate from the oligomerization / dimerization recirculation loop, a portion of the purge gas from the recirculation loop is used to supply hydrogen to LORP units 1 and 2. Similar to Example 1, in this example the fresh pure hydrogen was typically 99.9 mol% or higher, and the purity of the recirculated hydrogen was 97 mol%.

[0154] The purge gas pressure was 598 psig at extraction, meaning no additional compression was required to cascade the purge gas. This modification reduces the amount of fresh, pure hydrogen required for the entire system, helping to reduce the cost of producing and compressing fresh, pure H2 by 2%. This also has the additional benefit of reducing compression costs, as cited in Example 1. Table 3 shows the mass balance for both the basic cascade and hydrogen cascade cases for this example.

[0155] [Table 3]

[0156] Specific Embodiments The following will be explained in conjunction with specific embodiments, but it should be understood that this explanation is intended to illustrate the scope of the preceding explanation and the attached claims, and is not intended to limit them.

[0157] A first embodiment of the present disclosure is a process for providing hydrogen in a process for producing jet fuel from methanol, comprising generating a supply of hydrogen from a hydrogen production unit that is sent to one or more vessels in the process, wherein the additional supply of hydrogen is recovered from a reaction in which hydrogen has not been completely consumed and is incorporated into one or more recirculation flows from one or more reactors or vessels and sent to replenish the first supply hydrogen. Embodiments of the present disclosure are any or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the hydrogen production unit is an electrolytic device, a steam reformer, or a self-thermal reformer. Embodiments of the present disclosure are any or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the first supply hydrogen is sent to a hydrogenation reactor and a methanol synthesis reactor. Embodiments of the present disclosure are any or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the additional supply hydrogen is a second supply hydrogen recovered from a hydrogenation reactor downstream of an oligomerization reactor. Embodiments of the present disclosure are one or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the additional supply of hydrogen is a third supply hydrogen recovered from a hydrogenation reactor downstream of the DME scrubbing tower in the light olefin recovery process section. Embodiments of the present disclosure are one or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the additional supply of hydrogen is a fourth supply hydrogen recovered from a methanol synthesis unit. Embodiments of the present disclosure are one or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the second supply of hydrogen is sent from the hydrogenation reactor to a selective hydrogenation reactor in the light olefin recovery process section of the process. Embodiments of the present disclosure are one or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the second supply of hydrogen is sent from the hydrogenation reactor to an acetylene conversion reactor in the light olefin recovery process section of the process.Embodiments of this disclosure are one or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein a portion of the second supply hydrogen is sent to a methanol synthesis unit. Embodiments of this disclosure are one or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the third supply hydrogen is sent from the hydrogenation reactor to an acetylene conversion reactor in the light olefin recovery process (LORP) section of the process. Embodiments of this disclosure are one or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein a portion of the third supply hydrogen is sent to a methanol synthesis unit. Embodiments of this disclosure are one or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the fourth supply hydrogen is sent to a hydrogenation reactor downstream of the oligomerization reactor. Embodiments of the present disclosure are any or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the fourth supply hydrogen is sent from the hydrogenation reactor to a selective hydrogenation reactor in the light olefin recovery process (LORP) section of the process. Embodiments of the present disclosure are any or all of the embodiments from the first embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the fourth supply hydrogen is sent from the hydrogenation reactor to an acetylene conversion reactor in the light olefin recovery process (LORP) section of the process.

[0158] A second embodiment of the present disclosure is a process for providing hydrogen in a process for producing jet fuel from methanol, comprising generating a supply of hydrogen from a hydrogen production unit that is sent to one or more vessels in the process, wherein the additional supply of hydrogen is recovered from a reaction in which hydrogen has not been completely consumed, recovered from a purification unit in one or more recirculation flows from one or more reactors or vessels, and sent to replenish the first supply hydrogen. Embodiments of the present disclosure are any or all of the embodiments from the second embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the hydrogen production unit is an electrolytic device, a steam reformer, or a self-thermal reformer. Embodiments of the present disclosure are any or all of the embodiments from the second embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the first supply hydrogen is sent to a hydrogenation reactor and a methanol synthesis reactor. Embodiments of the present disclosure are any or all of the embodiments from the second embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the additional supply hydrogen is a second supply hydrogen recovered from a hydrogenation reactor downstream of an oligomerization reactor.

[0159] A third embodiment of the present disclosure is a process for providing hydrogen in a process for producing jet fuel from methanol, comprising generating a supply of hydrogen from a hydrogen production unit that is sent to one or more vessels in the process, the hydrogen production unit being an electrolytic cell, a steam reformer, or a self-thermal reformer, and the additional supply of hydrogen is recovered from a reaction in which hydrogen has not been completely consumed, recovered from a purification unit in one or more recirculation flows from one or more reactors or vessels, and sent to replenish the first supply hydrogen. Embodiments of the present disclosure are any or all of the embodiments from the third embodiment of this paragraph to the embodiment immediately preceding this paragraph, wherein the first supply hydrogen is sent to a hydrogenation reactor and a methanol synthesis reactor.

[0160] Without further detail, it is expected that those skilled in the art will be able to utilize the Disclosure to the fullest extent without departing from the spirit and scope of the Disclosure, readily identify its essential characteristics, and make various changes and modifications to adapt it to various uses and conditions. Accordingly, the prior preferred specific embodiments should be interpreted as merely illustrative examples and not in any way limiting the remainder of the Disclosure, but are intended to cover various modifications and equivalent configurations that fall within the scope of the appended claims.

[0161] In the above, all temperatures are given in degrees Celsius, and all parts and percentages are based on weight unless otherwise indicated.

Claims

1. A process for providing hydrogen in a process for producing jet fuel from methanol, comprising generating supply hydrogen from a hydrogen generation unit to be sent to one or more vessels in the process, recovering additional supply hydrogen from a reaction in which hydrogen has not been completely consumed, incorporating it into one or more recirculation flows from one or more reactors or vessels, and sending it to replenish the first supply hydrogen.

2. The process according to claim 1, wherein the hydrogen generation unit is an electrolytic cell, a steam reformer, or a self-heating reformer.

3. The process according to claim 2, wherein the first supply hydrogen is sent to a hydrogenation reactor and a methanol synthesis reactor.

4. The process according to claim 1, wherein the additional supply hydrogen is a second supply hydrogen recovered from a hydrogenation reactor downstream of the oligomerization reactor.

5. The process according to claim 1, wherein the additional supply hydrogen is a third supply hydrogen recovered from a hydrogenation reactor downstream of the DME scrubbing tower in the light olefin recovery process section.

6. The process according to claim 1, wherein the additional supply hydrogen is a fourth supply hydrogen recovered from a methanol synthesis unit.

7. The process according to claim 4, wherein the second supply hydrogen is sent from the hydrogenation reactor to a selective hydrogenation reactor in the light olefin recovery process section of the process, to an acetylene conversion reactor in the light olefin recovery process section of the process, or to a methanol synthesis unit.

8. The process according to claim 5, wherein the third supply hydrogen is sent from the hydrogenation reactor to an acetylene conversion reactor in the light olefin recovery process (LORP) section of the process, or to a methanol synthesis unit.

9. The process according to claim 6, wherein the fourth supply hydrogen is sent to a hydrogenation reactor located downstream of the oligomerization reactor.

10. The process according to claim 6, wherein the fourth supply hydrogen is sent from the hydrogenation reactor to a selective hydrogenation reactor in the light olefin recovery process (LORP) section of the process.

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