Production of synthesis gas and liquid fuels from oligomerization by-products

Abstract

A process for producing liquid fuels from carbon dioxide and hydrogen is disclosed. This process includes reacting a mixture of carbon dioxide and hydrogen to produce methanol. The methanol is contacted with an MTO catalyst to produce an olefin flow. The olefin flow is oligomerized with an oligomerizing catalyst to produce an oligomerized olefin flow, which is then hydrogenated to produce liquid fuel flows: jet fuel, diesel, and naphtha. A waste gas flow is produced at each stage. A synthesis gas flow containing carbon oxides and hydrogen is produced by (1) steam reforming, autothermal reforming, or dry reforming, or (2) partial oxidation.

Description

Technical Field

[0001] The present invention relates to a process for producing liquid fuels from carbon dioxide and hydrogen. This field may particularly relate to processes for converting syngas into liquid fuels such as naphtha, jet fuel, or diesel.

Background Art

[0002] The global need for low-carbon intensity liquid fuels that are compatible with existing infrastructure is driving increased research and investment in fuel refineries that utilize CO2 as a primary feedstock. In such facilities, CO2 reacts with hydrogen (produced from the electrolysis of water) to produce methanol, which can be refined to produce fuels that meet the physical and chemical property requirements of gasoline, jet, or diesel. The fuels are expected to dramatically reduce the life cycle carbon footprint of vehicles that obtain energy from renewable electricity (referred to as "e-fuels") and utilize engines powered by conventional combustion. The e-fuel production process can exist in various forms, one of which includes a water electrolysis unit, a methanol synthesis unit, and a methanol to jet (MTJ) unit for producing jet fuel from methanol. CO2 capture may also be involved.

[0003] The most desirable process for the market is one that produces a low carbon intensity and a high fuel yield. To achieve this, efficient utilization of both light and heavy by-products is essential. A process configuration that recycles waste streams in an efficient manner produces high-yield liquid fuels from carbon dioxide and hydrogen while minimizing incremental CO2 production.

Summary of the Invention

[0004] The inventors have devised a process for producing liquid fuels such as naphtha, jet fuel, or diesel, which involves improved overall selectivity (and reduced carbon intensity) of the CO2-to-jet complex. This improvement includes recirculating the heavier hydrocarbons obtained in this process through a reforming or partial oxidation process. [Brief explanation of the drawing]

[0005] [Figure 1] This is a schematic diagram of a process for producing jet fuel complexes from carbon oxides, showing the main processing blocks, interconnected flows, feeds, and products of the disclosure. [Figure 2] This invention provides different embodiments of a process for producing jet fuel complexes from carbon oxides, illustrating the main processing blocks, interconnected flows, feeds, and products, and including the recirculation of waste flows. [Figure 3] This invention provides different embodiments of a process for producing jet fuel complexes from carbon oxides, illustrating the main processing blocks, interconnected flows, feeds, and products, and including the recirculation of waste flows. [Figure 4] This invention provides different embodiments of a process for producing jet fuel complexes from carbon oxides, illustrating the main processing blocks, interconnected flows, feeds, and products, and including the recirculation of waste flows. [Figure 5] This is a schematic diagram of the methanol synthesis process and apparatus disclosed herein. [Figure 6] This diagram shows a schematic representation of the olefin feed preparation process and apparatus described herein. [Figure 7] This is a schematic diagram of the oligomerization process and apparatus disclosed herein. [Figure 8] This is a schematic diagram of the hydrogenation process and apparatus disclosed herein. [Figure 9] This is a schematic diagram of the partial oxidation section of the process and apparatus of this disclosure.

[0006] definition The term "communication" means that fluid flow is operably permitted between the listed components, and this can be characterized as "fluid communication."

[0007] 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.

[0008] 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.

[0009] The term "direct communication" means that the fluid flow from the upstream component enters the downstream component without passing through any other intervening container.

[0010] The term "indirect communication" refers to the flow of fluid from an upstream component entering a downstream component after passing through an intervening container.

[0011] The term “column” refers to one or more distillation columns 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 feed into 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 to an upper tray and removes the main product from the bottom.

[0012] As used herein, the term “diesel” means hydrocarbons that boil within the “diesel cutpoint” range, including IBP at about 125°C (257°F) to about 175°C (347°F) or T5 at about 150°C (302°F) to about 200°C (392°F), and T95 at about 343°C (650°F) to about 399°C (750°F) using the TBP distillation method, or T90 at 280°C (536°F) to about 340°C (644°F) using ASTM D-86. The term “green diesel” means diesel containing hydrocarbons that are not of fossil fuel origin.

[0013] As used herein, the term "green hydrogen" typically refers to hydrogen produced from non-fossil fuel sources by a water hydrolysis unit.

[0014] As used herein, the terms "T5," "T10," "T90," or "T95" mean, using ASTM D-86 or TBP, the boiling point of 5% by mass or volume, 10% by mass or volume, 90% by mass or volume, and 95% by mass or volume, respectively, of the sample.

[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 about 190°C (374°F) to about 215°C (419°F) and an endpoint of about 290°C (554°F) to about 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] As used herein, the term "concentrated stream of component" or "concentrated stream" means that the concentrated stream exiting the container has a higher concentration of the component than the feed to the container, preferably than all other streams withdrawn from the container.

[0019] As used herein, the term "dilute stream of component" or "dilute stream" means that the dilute stream exiting the container has a lower concentration of the component than the feed to the container, preferably than all other streams withdrawn from the container.

[0020] As used herein, the term "concentrated" means greater than 50%, preferably greater than 75%, more preferably greater than 90%.

[0021] As used herein, the term "separator" means a vessel having an inlet and at least a top vapor outlet and a bottom liquid outlet, and may also have 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 disclosed processes and apparatus include the production of liquid fuels from carbon oxides and hydrogen. The process includes reacting a mixture of carbon oxides and hydrogen to produce methanol, water, and a waste gas containing hydrogen, carbon monoxide, and dimethyl ether (DME). The methanol is contacted with an MTO catalyst and concentrated to produce one or more olefin streams and a waste gas containing hydrogen, carbon monoxide, and methane. The one or more olefin streams are oligomerized with one or more oligomerization catalysts to produce an oligomerized olefin stream containing light to heavy olefins and a waste gas stream containing hydrogen, ethane, propane, and light olefins. The oligomerized olefin stream, and in some embodiments, a portion of the light olefins, are reacted with hydrogen in the presence of a hydrogenation catalyst to produce jet fuel, diesel fuel, naphtha, and a waste gas stream containing hydrogen, propane, butane, and other light hydrocarbons. See FIG. 1.

[0023] The combination of MTO, oligomerization, and hydrogenation processes as described above results in a very high selectivity (greater than 75%) for jet fuel. However, the value of the by-products, diesel, naphtha, and waste gas depends on whether they can be effectively used as fuels. In situations where these streams cannot be used as fuels, it may be more economical to recycle them so that they can be converted to jet fuel. There are many routes for this recycling to occur, each of which ultimately converts unwanted hydrocarbons to a mixture of carbon oxides and hydrogen (syngas), which can be fed to a methanol synthesis unit and converted to methanol. The advantage of these routes is that they create syngas without the need for hydrogen obtained from electrolysis of water, which is the most energy-intensive part of the process of producing jet fuel composites from CO2.

[0024] In one embodiment, oxygen from a water electrolysis process may be introduced into a partial oxidation process to convert undesirable byproducts into synthesis gas, which may then be used as a ready feed for a methanol synthesis unit. In this embodiment, both a thermal partial oxidation process and a catalytic partial oxidation process are suitable, but the catalytic process is preferred due to the fact that the reactants do not contain contaminants to the catalyst. See Figure 2.

[0025] Partial oxidation processes are exothermic, providing their own heat to warm the reactants. In certain embodiments of a partial oxidation process, this process provides additional heat to the hydrogenation process used in the reboiler of a distillation column. See, for example, Figure 9.

[0026] In another embodiment, the methane and hydrogen produced in a specific waste gas stream by the above process can be mixed with CO2 and reacted in a dry reforming process without the use of steam. 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.

[0027] In yet another embodiment, autothermal reforming combines both partial oxidation and steam reforming. Oxygen, steam, and waste streams from a water electrolysis unit are reacted on a catalyst to produce a suitable synthesis gas mixture for conversion to methanol. The heat for this process is generated from partial oxidation by oxygen. See Figure 3.

[0028] In another embodiment, the waste stream reacts with steam in a high-temperature steam reforming process to create a synthesis gas mixture that can be directly supplied to methanol synthesis. This process may include a gas-heated reactor to improve the overall yield. This is an endothermic process and therefore requires a considerable amount of external heat. See Figure 4.

[0029] The following methods and apparatus for methanol synthesis, olefin feed preparation, oligomerization, and hydrogenation are given as examples and are not limited thereto. Other methods and apparatus may be used to carry out these processes.

[0030] Methanol synthesis 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.

[0031] A mixture of carbon oxide and hydrogen is reacted to produce methanol, water, and a waste gas stream containing hydrogen, carbon oxide, and dimethyl ether (DME).

[0032] Referring to Figure 5, 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 5, 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 produced by a water electrolysis unit.

[0033] 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 flow in line 122 and the hydrogen gas flow in line 124 are sent to the first methanol converter 140 of the methanol synthesis section 111. In embodiments, the synthesis gas flow in line 122 and the hydrogen gas flow in line 124 can be mixed to provide a mixed feed logistics 126, which is sent to the first methanol converter 140. However, the synthesis gas flow in line 122 and the hydrogen gas flow in line 124 may be sent to the first methanol converter 140 separately. The mixed feed logistics 126 can be sent to a synthesis gas booster compressor 130 to compress the synthesis gas to a specific pressure and provide a compressed synthesis gas flow into line 132 before being sent to the first methanol converter 140. In an exemplary embodiment, synthesis gas can be compressed to a pressure of approximately 6890 kPa (1000 psia) to approximately 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.

[0034] 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 about 3:2 to about 3:1.

[0035] A suitable methanol synthesis catalyst may be copper on a zinc oxide and alumina support. The synthesis conditions for the first methanol converter 140 in methanol synthesis section 111 may include a temperature of about 200 to about 300°C and a pressure of about 3.5 to about 10 MPa. Reaction equilibrium usually requires methanol separation and recycling of unreacted reagents back into the synthesis reaction to obtain sufficient conversion.

[0036] According to an exemplary embodiment, the first methanol converter 140 can operate at temperatures ranging from about 204°C (400°F) to about 290°C (550°F). According to another exemplary embodiment, the first methanol converter 140 can operate at pressures ranging from about 6890 kPa (1000 psia) to about 8970 kPa (1300 psia).

[0037] The methanol synthesis reaction is highly exothermic. Boiler feed water (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 recirculated BFW in line 149, providing BFW from line 148 for the first methanol converter 140.

[0038] 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.

[0039] 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 approximately 6890 kPa (1000 psia) to approximately 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.

[0040] 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 include a temperature of about 200 to about 300°C and a pressure of about 3.5 to about 10 MPa. Reaction equilibrium typically requires methanol separation and recirculation of unreacted reagents into the synthesis reaction.

[0041] 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.

[0042] In the second methanol converter 160, the first reactor outflow is converted into a methanol composition to provide a second reactor outflow containing methanol in line 162. The methanol flow in the second reactor outflow in line 162 may include methanol, dimethyl ether, ethanol, or a combination thereof. The second reactor outflow in line 162 can be withdrawn from the side of the second methanol converter 160. The second reactor outflow in line 162 is heat-exchanged with the compressed first vapor in line 157 in the heat exchanger 163. The heat-exchanged second reactor outflow in line 164 can be heated in the heater 165 to provide a heated second reactor outflow in line 166a. The heated second reactor outflow in line 166a is separated in the 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.

[0043] According to an exemplary embodiment, the second methanol converter 160 operates at a temperature of approximately 204°C (400°F) to approximately 290°C (550°F). According to another exemplary embodiment, the second methanol converter 160 operates at a pressure of approximately 6890 kPa (1000 psia) to approximately 8970 kPa (1300 psia).

[0044] 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 supply 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.

[0045] The PSA feed logistics in line 184a are 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 feed 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 stream 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 stream 186.

[0046] In the PSA unit 185, hydrogen present in the PSA supply logistics 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.

[0047] 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.

[0048] Crude methanol contains methanol, light fractions, and heavier alcohols. Where used and described herein, the terms “crude methanol” or “crude oxygenated feed” 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 fed to an oxygenated conversion unit for further production of light olefins. According to this disclosure, crude methanol may be fed to an oxygenated conversion unit or an MTO unit.

[0049] According to exemplary embodiments of the present invention, crude methanol may have a composition comprising carbon monoxide at a concentration of about 0 to about 1% by weight, carbon dioxide at a concentration of about 0.05 to about 2% by weight, methane at a concentration of about 0.001 to about 2% by weight, hydrogen at a concentration of about 0.05 to about 2% by weight, oxygen at a concentration of about 0 to about 1% by weight, water at a concentration of about 5 to about 18% by weight, nitrogen at a concentration of about 0 to about 1% by weight, methanol at a concentration of about 75 to about 90% by weight, and alcohol (other than methanol) at a concentration of about 0.05 to about 4% by weight.

[0050] 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 withdrawn from the crude methanol hold-up tank 195.

[0051] Conventionally, the crude methanol stream in line 196 is purified from light gases and heavy oxygenates before being fed into an MTO reactor, such as the oxygenate conversion reactor 16 in Figure 6.

[0052] The crude methanol stream in line 198 may contain methanol, dimethyl ether, ethanol, or a combination thereof.

[0053] 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 5. 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.

[0054] The first distillation column bottom flow containing methanol in line 218 is removed for further separation. The first distillation column bottom flow 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 first distillation bottom. According to an exemplary embodiment, the first distillation column 210 is operated at a pressure of about 689 kPa (100 psia) to about 1379 kPa (200 psia). According to another exemplary embodiment, the first distillation column is operated at a temperature of about 27°C (80°F) to about 177°C (350°F).

[0055] 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 top flow in line 222 is condensed for reflux to the column in line 228. From the heat exchanger 223, the condensed second distillation column top flow in line 224 is sent to the second top receiver 225. In the second top receiver 225, a portion of the condensed liquid from the second distillation column top flow in line 224 is recycled to the second distillation column 220 via line 228. The remaining liquid from the top of the second distillation column is sent via line 199 to the MTO reactor 16 in Figure 3, which is the methanol feed.

[0056] 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 bottom section.

[0057] According to an exemplary embodiment, the second distillation column operates at a pressure of approximately 517 kPa (75 psia) to approximately 862 kPa (125 psia). Furthermore, according to an exemplary embodiment, the second distillation column operates at a temperature of approximately 104°C (220°F) to approximately 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.

[0058] Conversion of methanol to olefins Next, the obtained methanol is brought into contact with the MTO catalyst to generate an olefin flow.

[0059] 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 approximately 2 MPa to approximately 3.8 MPa. The MTO reaction temperature should be approximately 325 to approximately 450°C. The space velocity per second ("WHSV") in the MTO reactor is in the range of approximately 2 to approximately 15 per second. The MTO catalyst is separated from the product olefin stream after the MTO reaction.

[0060] Figure 6 shows the olefin feed preparation process and apparatus 10 of the present disclosure. The methanol feed stream in line 199' in Figure 5 is fed through line 12 in Figure 6 to an oxygenate conversion reactor (MTO reactor) 16, which reacts an oxygenate such as methanol or dimethyl ether (DME) with a flow catalyst. The MTO catalyst may be a silicoaluminophosphate (SAPO) catalyst. SAPO catalysts and their formulations are generally taught in U.S. Patents 4,499,327(A), 10,358,394, and 10,384,986. The MTO reaction conditions include contact with the SAPO catalyst at a pressure of about 2 MPa to about 3.8 MPa. The MTO reaction temperature should be about 325 to about 450°C. The space velocity per second ("WHSV") in the oxygenate conversion reactor 16 is in the range of about 2 to about 15 per second. The MTO catalyst is separated from the product olefin stream after the MTO reaction. The high-temperature vapor reactor effluent in line 14 can be pre-cooled in the 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 may be taken in from the stripped water flow in line 21, and by direct contact with the water flow supplied in line 19, superheating is removed, organic acids are neutralized, and catalyst particles are removed. The quenched reactor effluent in line 22 is discharged from the quenching tower 20 and supplied to the product separator tower 24. The product separator tower 24 may be downstream-communicated with the oxygenate conversion reactor 16.

[0061] 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, resulting in 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 tower 30. A water return flow containing oxygenated by-products from the compression section 80 in the return line 32 can also be passed to the water stripper tower 30. The water stripper tower 330 may be connected downstream to the product separator tower 24.

[0062] The vapor flow from the first section of the product separator column 24 is sent to the second section of the product separator column, i.e., the upper section. The intermediate flow in line 28, containing hydrocarbons, oxygenated by-products, and liquid-phase water, is taken out at the bottom of the upper section. A portion of the intermediate flow in line 28 is cooled and sent as reflux to the top of the second section of the product separator 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 separator column, is sent to the compression section 80.

[0063] 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 at the top line 49, where methanol and at least one other oxygenate are concentrated, and a water-rich stream at the bottom line 46.

[0064] In one embodiment, the temperature of the water stripper tower 30 may be 115°C (239°F) to 180°C (356°F) at the bottom of the water stripper tower, and the pressure may be approximately 75 kPa gauge (11 psig) to approximately 760 kPa (110 psig) at the top of the water stripper tower.

[0065] 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 pressurized first olefin-concentrated stream in the top line 83 at a temperature of about 40°C (104°F) to about 60°C (140°F) and a pressure of about 193 kPa(g)(28 psig) to about 262 kPa(g)(38 psig) and an oxygenated first aqueous stream in the bottom line 84. The olefin-concentrated stream in the top line 83 may be fed to a compressor 85, cooled, and directed to a second knockout drum 86. The aqueous flow in the bottom line 84, along with the product aqueous flow in the bottom line 36 of the separator tower, is pumped via the manifold line 76 to a return line 32 that returns the aqueous flow to the water stripper tower 30.

[0066] The compression section 80 may include a second knockout drum 86, which separates the pressurized first olefin-concentrated flow into a second pressurized olefin-concentrated flow in the top line 87 at a pressure of approximately 330 kPa(g)(48 psig) to approximately 400 kPa(g)(58 psig) and a temperature of approximately 27°C(80°F) to approximately 54°C(130°F) and into an oxygenated second aqueous flow in the bottom line 88. The second olefin-concentrated flow in 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 separator column.

[0067] 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 to the oxygenated absorption tower 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 tower 30 along with the generated water stream in the bottom line 36 of the separator tower.

[0068] 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 may be approximately 1 MPa gauge (145 psig) to approximately 2 MPa gauge (290 psig). The compressor discharge may be cooled to near ambient temperature using conventional heat transfer methods.

[0069] As shown in Figure 6, 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 separator tower 24, under conditions effective for absorbing the oxygenation without prior removal. 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 of the oxygenation absorption tower can be in a bottom temperature range of about 30°C (86°F) to about 60°C (140°F) and a top pressure range of about 700 kPa gauge (101 psig) to about 1 MPa gauge (145 psig).

[0070] The absorption olefin-containing flow in the top line 54 can be supplied 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 to the stripper separator 66 via line 65. The stripper separator separates the aqueous flow containing oxygenates in the boot in line 67, which is supplied 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 tower 70 to remove C3- and lower vapors in the stripper tower top line 71 from the heavy olefin liquid flow in the stripper bottom line 168. Most of the oxygenated material is stripped into the stripper tower top line 71 and separated upon cooling and recirculation to the stripper separator 66. The bottom flow exiting the DME stripper tower 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 of approximately 30°C (86°F) to approximately 60°C (140°F) and pressures of approximately 1.7 MPa (g) (250 psig) to approximately 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.

[0071] 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 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 oligomerized feed preparation section.

[0072] Oligomerization and hydrogenation Olefin oligomerization is a process that allows smaller olefins to be oligomerized into larger olefins. More specifically, light olefins, including oligomerized olefins, can be converted into higher carbon-number olefins, including products in the gasoline, jet, and diesel ranges. Olefins can be saturated for use as transport fuels.

[0073] Jet fuel is one of the few petroleum fuels that cannot be easily replaced by electric motor systems because the fuel for aircraft 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.

[0074] The oligomerization section 310 is shown in Figure 7. 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.

[0075] Referring to the oligomerization section 310 in Figure 7, 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 80 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 about 60°C (140°F) to about 150°C (302°F), preferably about 80°C (176°F) to about 100°C (212°F), and a pressure of about 3.5 MPag (500 psig), preferably about 5.6 MPag (800 psig) to about 8.4 MPag (1200 psig).

[0076] 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.

[0077] 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 7, 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 about 70 to about 90% by volume of the input olefin flow, and the second input olefin flow in the second olefin line 312b may contain about 10 to about 30% by volume of the input olefin flow.

[0078] 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 about 2 to about 8 times, preferably about 3 to about 6 times, the volumetric flow rate of the input olefin stream in the input olefin line 312.

[0079] 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.

[0080] 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 about 10 to about 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 about 10 to about 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 about 10 to about 20% by weight of propylene.

[0081] 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 7, two first-stage oligomerization reactor vessels 321a and 321b are used.

[0082] 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 7 shows only two reaction vessels 321a and 321b.

[0083] 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 and diluted first input olefin stream in line 320a can be introduced at a temperature of about 180°C (356°F) to about 260°C (500°F) and a pressure of about 3.5 MPag (500 psig) to about 8.4 MPag (1200 psig). The input cooler 318a may be equipped with a steam generator.

[0084] The diluted first input olefin stream may be introduced into the first first-stage catalyst bed 322a in line 320a, preferably in a downward flow operation. However, an upward flow operation may be preferable. As oligomerization of ethylene, propylene, and recycled 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 approximately 250°C (482°F).

[0085] The second input olefin stream in line 312b can 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 may be cooled in a second input cooler 318b which can be located outside the first first-stage oligomerization reactor 321a, and the cooled second input olefin stream may be supplied into line 320b and introduced into the second bed 322b of the first-stage oligomerization catalyst of the first first-stage oligomerization reactor 321a. The input cooler 318b may include a steam generator.

[0086] The second input olefin stream in line 320b may be introduced at a temperature of approximately 180°C (356°F) to approximately 230°C (446°F) and a pressure of approximately 3.5 MPag (500 psig) to approximately 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 increased outlet temperature can be limited to a temperature 30°C (54°F) to approximately 50°C (90°F) higher than the inlet temperature to the catalyst bed 322b.

[0087] 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 may be 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 may 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.

[0088] The cooled first recirculating olefin feed stream in line 320c can be fed at a temperature of approximately 180°C (356°F) to approximately 230°C (446°F) and a pressure of approximately 3.5 MPag (500 psig) to approximately 8.4 MPag (1200 psig). The first recirculating olefin feed stream contains a diluent and olefins from the second oligomerized olefin stream and the third recirculating 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 oligomerizing catalyst, as well as the oligomerization of oligomers in the first recirculating olefin feed stream, generates a third oligomerized effluent flow 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 approximately 50°C (90°F) higher than the inlet temperature to the catalyst bed 322c.

[0089] 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 the diluent stream from line 314, which 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 may be 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.

[0090] The cooled second recirculating olefin feed stream in line 320d can be fed at a temperature of approximately 180°C (356°F) to approximately 230°C (446°F) and a pressure of approximately 3.5 MPa(g)(500 psig) to approximately 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 elevated outlet temperature is limited to a temperature 30°C (54°F) to approximately 50°C (90°F) higher than the inlet temperature to the catalyst bed 322d.

[0091] 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.

[0092] 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.

[0093] 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 about 5 to about 95% by weight of the catalyst, for example, about 5 to about 85% by weight. 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.

[0094] 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 create 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.

[0095] One of the components of the catalyst binder used in the present invention is alumina. The alumina source can be any of various hydrated aluminum oxides or alumina gels, such as alpha-alumina monohydrate 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.

[0096] 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 about 5–85% by weight, for example, about 20–82% by weight of MTT zeolite, and the remaining alumina powder provides a well-supported catalyst. Silica supports are also considered.

[0097] 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.

[0098] 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 a temperature of about 260°C (500°F) to about 815°C (1500°F). The MTT catalyst does not have the selectivity to neutralize acidic sites with amines or the like.

[0099] Extruded particles may have any suitable cross-sectional shape, i.e., symmetric or asymmetric, but in most cases 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 typically about 0.635 mm (0.25 inches) to about 12.7 mm (0.5 inches), preferably about 0.79 mm (1 / 32 inch) to about 6.35 mm (0.25 inches), and most preferably about 0.06 mm (1 / 24 inch) to about 4.23 mm (1 / 6 inch).

[0100] In one exemplary embodiment, an MTT-type zeolite catalyst, arranged on a high-purity pseudo-boehmite alumina substrate in a ratio of approximately 90 / 10 to approximately 20 / 80, preferably approximately 20 / 80 to approximately 50 / 50, is provided to the catalyst bed, or more often, to the first-stage oligomerization reactor 322.

[0101] 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 approximately 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.

[0102] 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.

[0103] 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 approximately 80°C (176°F) to approximately 180°C (356°F). The second-stage oligomerization reactor 332 is operated at a pressure of approximately 2.1 MPa (300 psig) to approximately 7.6 MPa (1100 psig), more preferably approximately 3.5 MPa (500 psig) to approximately 6.9 MPa (1000 psig).

[0104] 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.

[0105] 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.

[0106] 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 catalysts of the present invention may have a low-temperature acidity ratio of at least about 0.15, preferably about 0.2, and preferably about 0.25 or greater, as determined by Ammonia Temperature Programmed Desorption (Ammonia TPD) as described below. In addition, preferred catalysts have a surface area of ​​about 50 to about 400 m² / g, as determined by nitrogen BET.

[0107] A preferred second-stage oligomerization catalyst comprises an amorphous silica-alumina support. One of the components of the catalyst support used in the present invention is alumina. The alumina can 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.

[0108] Another component used in the preparation of the second-stage oligomerization catalyst used in the present invention is a surfactant. The surfactant is preferably mixed with the above-mentioned alumina and silica-alumina powders. The resulting mixture of surfactant, alumina, and silica-alumina is then formed, dried, and calcined as described below. Calcination effectively removes the organic components of the surfactant by combustion, but only after the surfactant has faithfully performed its function according to the present invention. Any suitable surfactant can be used according to the present invention. Preferred surfactants are selected from a range of commercially available surfactants marketed by Solvay SA under the trademark "Antarox". "Antarox" surfactants are generally characterized as modified linear aliphatic polyethers and are low-foaming biodegradable detergents and wetting agents.

[0109] 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, amorphous silica-alumina with a silica:alumina ratio of 2.6 at about 75 to about 99% by weight and alumina powder at about 10 to about 20% by weight provides a suitable carrier. In another embodiment, other ratios of amorphous silica-alumina to alumina may be preferred.

[0110] 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 about 0.01 to about 10% by weight, based on the weight of alumina and silica-alumina.

[0111] 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.

[0112] 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 about 260°C (500°F) to about 815°C (1500°F).

[0113] Extruded particles may have any suitable cross-sectional shape, i.e., symmetric or asymmetric, but in most cases 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 typically about 0.635 mm (0.25 inches) to about 12.7 mm (0.5 inches), preferably about 0.79 mm (1 / 32 inch) to about 6.35 mm (0.25 inches), and most preferably about 0.06 mm (1 / 24 inch) to about 4.23 mm (1 / 6 inch).

[0114] 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 about 0.2 to about 2.0 cc / gram, preferably about 0.25 to about 1.0 cc / gram, and most preferably about 0.3 to about 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 about 0.1 cc / gram, preferably less than 0.08 cc / gram, and most preferably less than about 0.05 cc / gram. The surface area, as measured by BET, is typically greater than 50 m² / gram, for example more than about 200 m² / gram, preferably at least 250 m² / gram, and most preferably about 300 m² / gram to about 400 m² / gram.

[0115] 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 about 0.5 to about 15% by weight, and the Group VIB metal, preferably tungsten, must be present at a concentration of about 0 to about 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 approximately 10 percent less and approximately 10 percent more than the volume required to just fill the pores.

[0116] 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.

[0117] A preferred second-stage oligomerization catalyst of the present invention has an amorphous silica-alumina substrate impregnated with about 0.5 to about 15% by weight of nickel in the form of a 3.175 mm (0.125 inch) extruded material, having a density of about 0.45 to about 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.

[0118] The second-stage oligomerization catalyst can be regenerated when deactivated. Suitable regeneration conditions include, for example, exposing the catalyst to hot air at approximately 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.

[0119] The second-stage oligomerization reaction is also inherently exothermic. The final oligomerized olefin stream in line 324d includes a diluent stream from diluent line 314, which is added to the first olefin stream in line 312a and transported 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.

[0120] When the oligomerization reaction is carried out according to the above process conditions, a C4 olefin conversion rate of approximately 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 in the distillate range.

[0121] The oligomerized olefin flow in line 334, which has an increased C8+ olefin concentration compared to the input first-stage oligomerized flow in line 328, is heat-exchanged with the first-stage oligomerized flow in line 324d, its pressure is reduced, and then it is heat-exchanged with the olefin splitter bottom flow in line 30 and supplied to the dealkane column 340. The oligomerized olefin flow in line 334 is at a temperature of approximately 160°C (320°F) to approximately 190°C (374°F) and a pressure of approximately 3.9 MPa (gauge) (550 psig) to approximately 7 MPa (gauge) (1000 psig). The olefin splitter bottom flow in line 330 can be transported to the hydrogenation section.

[0122] 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. The alkane flow may also be converted to synthesis gas via steam reforming, partial oxidation, autothermal reforming, or dry reforming, and the synthesis gas can be recycled to a methanol synthesis unit.

[0123] In the dealkane column 340, light alkanes, preferably C2 hydrocarbons, are separated from the alkane top flow in the top line 342 of Figure 7 and the dealkane bottom flow in the bottom line 344, which probably contains C4+ hydrocarbons, preferably C3+ hydrocarbons. When operating as an ethane dealternator, the dealkane column 340 can operate at a bottom temperature of approximately 177°C (350°F) to approximately 302°C (575°F) and a top pressure of approximately 207 kPa (gauge) (30 psig) to approximately 690 kPa (gauge) (100 psig). When operating as a propane dealternator, the dealkane column 340 can operate at a bottom temperature of approximately 194°C (381°F) to approximately 333°C (630°F) and a top pressure of approximately 207 kPa (gauge) (30 psig) to approximately 1.14 MPa (gauge) (165 psig).

[0124] The alkane top flow in the top line 342 can be cooled and separated in the dealkane receiver 346 to provide dealkalized off-gas in the off-gas line 347, which can then be supplied for further processing, such as to a synthesis gas production reactor, or it can be taken up as fuel gas in line 348 along with the net vapor flow in the receiver top line 368. 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 may 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 bottom line 344 may be divided into a reboiling flow in line 350, which is reboiled by a heat exchanger with a first high-temperature diesel flow in line 352, which is taken in from the jet fractionation bottom heat exchanger flow in 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 feed to the dealkane column 340 is not preheated by the olefin splitter bottom flow in line 330, but the feed to the olefin splitter column 360 in net bottom line 354 is preheated by the olefin splitter bottom flow.

[0125] 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, presumably in the olefin splitter top line 362, and a heavy olefin flow, presumably in the olefin splitter bottom line 364. The olefin splitter top flow is cooled to approximately 66°C (150°F) to approximately 93°C (200°F), and the resulting condensate can be refluxed from the olefin splitter receiver 366 to the olefin splitter column 360. The net vapor flow from the olefin splitter receiver 366 to the receiver top line 368, 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 refluxes back into the tower 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 about 1 to about 15% by weight of the light olefin flow in line 370. The light olefin flow in line 372 may contain about 40 to about 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 the top tower line 377, and the liquid recirculation olefin oligomer flow in line 326 can be recycled to the first-stage oligomerization reactor 322 to oligomerize C4-C8 olefins.

[0126] The heavy olefin flow in the splitter bottom line 364 may be split between a reboiling flow in the splitter reboiling line 365, which is reboiled by heat exchange with a second high-temperature diesel flow in line 373, which is taken from the jet fractionation bottom heat exchange flow in the jet bottom heat exchange line 374 in Figure 8 and returned to the olefin splitter column 360. The cooled second high-temperature diesel flow in line 412 is returned to the hydrogenation section 410 in Figure 8 for reboiling and then returned to the jet fractionation column 400. The heavy olefin flow in the net bottom line 330 is cooled by heat exchange with the second-stage oligomerization flow in line 334 and then transported to a hydrogenation section (not shown). The heavy olefin flow contains C9+ olefins, which, once cooled, may be transported to the hydrogenation section 410.

[0127] Turning to the hydrogenation section 410 in Figure 8, the heavy olefin stream in the net olefin splitter bottom line 330 from Figure 7, containing C9+ oligomerized olefins in the distillate range, can be hydrogenated in the hydrogenation reactor 380 to saturate the olefin bonds and provide fuel. This process is carried out to ensure that the product motor fuel meets or exceeds the thermal oxidation requirements specified in ASTM D7566-10a for hydrogenated synthetic paraffinic kerosene (SPK). Furthermore, hydrogenating the oligomerized heavy olefins yields a paraffin stream that can be used as a diluent stream in line 314. The heavy olefin stream in line 330 can be cooled to generate steam, which can be combined with a light olefin liquid stream containing C2-C8 olefins in line 377 of Figure 7 to produce a composite olefin stream in line 379. Alternatively, the composite olefin flow in line 379 may be combined with the hydrogen flow in line 376 to provide a composite hydrogenation feed stream in line 381, which is cooled and fed into the hydrogenation reactor 380 at 125°C (257°F) to approximately 204°C (400°F) and 3.5 MPa (500 psig) to approximately 6.9 MPa (1000 psig). Excess hydrogen, such as approximately 1.5 to 2.5 stoichiometric hydrogens, can be used to ensure complete saturation.

[0128] Hydrogenation is typically carried out using conventional hydrogenation or hydrogenation catalysts, which may include, for example, metal catalysts containing palladium, rhodium, nickel, ruthenium, platinum, rhenium, cobalt, molybdenum, or combinations thereof, and supported metal catalysts thereof. Catalyst supports may be oxides of silica, alumina, titania, calcium carbonate, barium sulfate, and carbon, but may be any solid inert material. Catalyst supports may be in the form of powders, granules, pellets, etc.

[0129] In exemplary embodiments, hydrogenation is carried out in a hydrogenation reactor 380 containing an alumina-platinum catalyst, for example, about 0.5% to about 0.9% by weight of the alumina-platinum catalyst. In another embodiment, the hydrogenation catalyst contains about 5% to about 30% by weight of the nickel catalyst. The hydrogenation reactor 380 converts the olefin into a paraffin product having the same carbon number distribution as the olefin, thereby forming paraffins with a distillate range suitable for use as jet and diesel fuels.

[0130] The hydrogenated heavy mass flow discharged from the hydrogenation reactor 380 to line 383 can be separated in a high-temperature separator 382 that provides hydrocarbon splits. In the high-temperature separator 382, ​​the hydrogenated heavy mass flow is separated into a high-temperature hydrogenated steam flow in the top line 384 and a high-temperature hydrogenated liquid flow in the bottom line 386 of the high-temperature separator. The hydrogenated heavy mass liquid flow in the bottom line 386 may be heated by heat exchange with the diluent flow in line 314 before the diluent flow is recirculated to the first-stage oligomerization reactor 322 in Figure 7. The heated hydrogenated heavy mass liquid flow in the high-temperature bottom line 386 can be supplied to the stripper column 390. The high-temperature separator can operate at temperatures from about 204°C (400°F) to about 343°C (650°F) and pressures from 3.5 MPa (500 psig) to about 6.9 MPa (1000 psig).

[0131] The high-temperature hydrogenation vapor flow in the high-temperature top line 384 can be cooled and supplied to the cryogenic separator 388. The cryogenic separator separates the cooled high-temperature hydrogenation vapor flow in the high-temperature top line 384 into a cryogenic vapor hydrogenation flow in the cryogenic top line 387 and a cryogenic heavy hydrogenation liquid flow in the cryogenic bottom line 389. The purge flow in the purge line 385 can be taken from the cryogenic vapor hydrogenation flow in the cold top line 387, and the remainder can be compressed and combined with makeup hydrogen in line 388a to provide a hydrogen flow in line 376. The cryogenic heavy hydrogenation liquid flow in the bottom line 389 may be supplied to the stripping tower 390 at a supply position higher than the supply position of the high-temperature heavy hydrogenation liquid flow in the high-temperature separator bottom line 386. The cryogenic separator can operate at temperatures ranging from approximately 32°C (90°F) to approximately 71°C (150°F) and pressures ranging from approximately 3.7 MPa (540 psig) to approximately 4.5 MPa (650 psig).

[0132] The stripping tower 390 may also be a flash stripper for removing light gases from the high-temperature hydrogenated liquid flow in the high-temperature bottom line 386 and the low-temperature hydrogenated liquid flow in the low-temperature bottom line 389. Both of these flows can be fed into the stripping tower 390 together or separately as shown in the figure. The stripping tower 390 removes residual light gases from the liquid hydrogenated flow and provides the stripper tower top flow to the stripper tower top line 392 and the stripped bottom flow to the stripper tower bottom line 394. The stripper tower top flow in the stripper tower top line 392 is cooled and separated in the stripper receiver 396 to provide the stripper off-gas flow in the stripper receiver tower top line 397 and the condensed flow in line 398 which is returned to the tower. The stripping tower 390 can operate at bottom temperatures of approximately 232°C (450°F) to approximately 316°C (600°F) and top pressures of approximately 207 kPa (30 psig) to approximately 689 kPa (100 psig).

[0133] After being stripped in the stripping column 390 to remove volatiles, the stripped fuel flow in the stripper bottom line 394 may be supplied to the jet fractionation column 400 without further heating. Alternatively, the stripping column 390 upstream of the jet fractionation column 400 may be omitted. In the jet fractionation column 400, the stripped fuel flow can be separated into a jet-off gas flow in the top line 402, a green jet flow in the side line 404 from the side of the jet fractionation column 400, and a green diesel flow in the bottom line 406. The jet fractionation column 400 can operate at a bottom temperature of approximately 288°C (550°F) to approximately 371°C (700°F) and a top pressure of approximately 35 kPa (5 psig) to approximately 350 kPa (50 psig).

[0134] The jet fraction top flow in the top line 402 may be cooled, and the resulting condensate is refluxed from the jet fraction receiver 408 and returned to the jet fraction 400 in the jet fraction top liquid line 409, while the net off-gas flow containing C8 hydrocarbons is taken from the jet fraction receiver 408 into the receiver top line 405. The majority of the hydrocarbons in the net off-gas flow in the receiver top line 405 are lighter hydrocarbons and can be used to fuel the reboiling heater 416 for the jet fraction column 400.

[0135] The green jet stream taken into the side line 404 contains C9-C17 hydrocarbons in the kerosene range, is cooled, and can be obtained as a jet fuel product that meets the applicable SPK standards. In an alternative embodiment, the green jet stream can be taken from the condensate flow in line 409 from the jet fractionation receiver 408, instead of refluxing all of the condensate back into the tower. This green jet stream taken from line 409 may need to be further stripped to remove light fractions. In such an embodiment, the side line 404 for recovering the green jet fuel stream is not provided.

[0136] The green diesel bottom flow in bottom line 406 can be split into a reboiling diesel flow in line 407 and a diesel product flow in line 414. The reboiling diesel flow in line 407 can be split into a jet bottom heat exchange flow 374 and a bypass bottom flow in bottom bypass line 411. As shown in Figure 7, the jet bottom heat exchange flow in jet bottom heat exchange line 374 can be split into a first high-temperature diesel flow in line 352 and a second high-temperature diesel flow in line 373 to provide reboiling heat to the dealkane column 340 and the olefin splitter column 360, respectively. The bypass bottom flow in line 411 is taken into the jet reboiling line 413 through a valve above it, or a portion of the cooled second high-temperature diesel flow in line 412, or both, is taken into the jet reboiling line 413, reboiled in the furnace 416, and returned to the jet fractionation column 400.

[0137] The diesel product flow in line 414 is divided into a diesel product flow in diesel product line 418 and a diluent flow in line 314. The diluent flow in line 314 can be cooled by heat exchange with the high-temperature hydrogenated heavy liquid flow in the high-temperature separator bottom line 386, recirculated, and mixed with the olefin flow in line 312 of the oligomerization section 310 in Figure 7, preferably with the first charge olefin flow in line 312a, to provide a first diluent olefin charge flow to line 316a to absorb heat in the oligomerization reactor 322. The green diesel flow in diluent line 314 is paraffinic and therefore inert to the oligomerization and hydrogenation reactions to which it can be supplied. The diesel product flow in diesel product line 418 can be cooled and supplied to the diesel pool. The diesel flow conforms to the ASTM D975 standard for diesel.

[0138] Starting from ethylene and / or propylene, the disclosed process can efficiently produce green jet fuel and green diesel fuel to meet applicable fuel requirements while controlling exothermic heat. The carbon recovery rate in the process can exceed 95%. Both the jet fuel flow in side line 404 and the diesel product flow in line 418 can be cooled and supplied to their respective fuel pools.

[0139] 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. When steam reforming is selected to produce synthesis gas from ethane / propane, a steam feed (which can be supplied entirely by a methanol-jet process) and energy input (which can be supplied partially by steam from the methanol-jet process) are also required. Dry reforming requires energy input and CO2.

[0140] The synthesis gas flow of this disclosure may have an H2 / CO ratio of about 0.5 to about 4, preferably about 2.

[0141] 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 7 using steam in a steam reforming reactor, a self-thermal reforming reactor, or a dry reforming reactor.

[0142] Steam reforming By steam reforming the alkane top flow from the top line 342 in Figure 7 at high temperature using CO2 and steam, a synthesis gas mixture that can be directly supplied to methanol synthesis is created. This is an endothermic process and therefore requires a considerable amount of external heat.

[0143] The alkane top flow 342 is mixed with 1-2 carbon equivalents of carbon dioxide to prevent the formation of elemental carbon. The alkane flow and carbon dioxide mixture are then mixed with steam to produce a mixture containing approximately 1 mole of carbon from hydrocarbons in 1-2 moles of water. The gas / steam mixture enters a steam reforming process unit with a commercially available nickel catalyst. The conversion of hydrocarbons to carbon monoxide and hydrogen takes place at 700°C-900°C and a system pressure that can be 12-500 psia. A typical flow rate for such a reactor is approximately 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.

[0144] 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.

[0145] The alkane top flow from the top line 342 in Figure 7 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.

[0146] 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 7 reacts in the dry reformer to produce synthesis gas. The resulting synthesis gas can be recycled to the methanol synthesis unit.

[0147] 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 7 using oxygen in a partial oxidation reactor.

[0148] Figure 9 shows the partial oxidation of the waste flow to supply heat to the jet fractionator (element 100 in Figure 8) in the oligomerization unit hydrogenation section. The partial oxidation of the waste flow is integrated with the circulating jet fractionator bottom.

[0149] The waste gas / liquid flow is added in line 424 and supplied to a preheater. The preheated flow is then supplied to the partial oxidation unit 420 (which has a partial oxidation catalyst (shaded) at the bottom of the reactor). Oxygen is also added to the partial oxidation unit 420 in line 418 at a substoichiometric ratio to achieve only partial oxidation of hydrocarbons and oxygenates to carbon monoxide. The effluent from the partial oxidation unit 420 is sent to a preheater to heat the feed, then to a jet fractional reboiler, and then to a steam generator to recover any residual heat from the flow. The carbon monoxide-rich partial oxidation synthesis gas flow is recovered into line 422 at approximately 400°F.

[0150] The boiler feedwater 430 is also supplied to the low-temperature side of the steam generator, and steam and water flow through it and are generated in 432.

[0151] The jet fractionation bottom liquid 428 (from either all or part of 113 in Figure 8 (circulating liquid from the jet fractionation reboiler heater (116))) is supplied to the jet fractionation reboiler, where it is partially vaporized. The jet fractionation reboiler vapor and liquid are returned to the jet fractionation column (100 in Figure 8) via line 426.

[0152] The resulting synthesis gas can be recycled to a methanol synthesis unit; for example, the partially oxidized flow from line 422 can be supplied to line 122 in Figure 5 for use in methanol synthesis.

[0153] 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.

[0154] A first embodiment of the present disclosure is a process for producing liquid fuel from carbon dioxide and hydrogen, comprising: (a) reacting a mixture of carbon dioxide and hydrogen to produce a methanol stream, a water stream, and one or more first waste gas streams; (b) contacting methanol with an MTO catalyst to produce one or more olefin streams and a second waste gas stream; (c) oligomerizing one or more of the olefin streams using one or more oligomerization catalysts to produce an oligomerized olefin stream and a third waste gas stream; and (d) hydrogenating the oligomerized olefin streams using a hydrogenation catalyst in the presence of hydrogen to produce naphth A process comprising: (e) a step of generating a naphtha stream, a jet fuel stream, a diesel stream, and a fourth waste gas stream; and (e) a step of generating a synthesis gas stream containing carbon oxides and hydrogen by (1) reforming one or more of the naphtha stream, diesel stream, first waste gas stream, second waste gas stream, third waste gas stream, or fourth waste gas stream with steam in a steam reforming reactor, self-thermal reforming reactor, or dry reforming reactor; or (2) partially oxidizing one or more of the naphtha stream, diesel stream, first waste gas stream, second waste gas stream, third waste gas stream, or fourth waste gas stream. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraph to the first embodiment of this paragraph, wherein the synthesis gas stream is recycled to step (a). Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraph to the first embodiment of this paragraph, wherein the hydrogen in step (a) is green hydrogen. Embodiments of the present disclosure are any or all of the embodiments from the earlier embodiments in this paragraph to the first embodiment in this paragraph, wherein the carbon oxides include CO and CO2. Embodiments of the present disclosure are any or all of the embodiments from the earlier embodiments in this paragraph to the first embodiment in this paragraph, wherein the synthesis gas stream of step (e) has an H2 / CO ratio of about 0.5 to about 4. Embodiments of the present disclosure are any or all of the embodiments from the earlier embodiments in this paragraph to the first embodiment in this paragraph, wherein the alkane stream includes methane, ethane and / or propane.Embodiments of the present disclosure are any or all of the embodiments from the earlier embodiments of this paragraph to the first embodiment of this paragraph, wherein the methanol synthesis process, MTO process, and / or oligomerization process generates vapor, which is used in the reforming step (e). Embodiments of the present disclosure are any or all of the embodiments from the earlier embodiments of this paragraph to the first embodiment of this paragraph, wherein step (a) is carried out in the presence of a catalyst. Embodiments of the present disclosure are any or all of the embodiments from the earlier embodiments of this paragraph to the first embodiment of this paragraph, wherein the liquid fuel flow in step (d) includes a naphtha flow, a jet flow, and a diesel fuel flow.

[0155] A second embodiment of the present disclosure is a process for producing liquid fuel from carbon dioxide and hydrogen, comprising: (a) 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) purifying the crude methanol stream by distillation to remove light contaminants into a first waste gas stream and heavy contaminants and water into a second heavy waste stream to produce a purified methanol stream; and (c) using the purified methanol stream with an MTO catalyst. (d) a step of bringing into contact with 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) purifying the crude olefin stream by distillation to remove light fractions into a third waste gas stream, then absorbing water to remove heavy oxygenated hydrocarbons, extracting water to remove dimethyl ether into a DME (dimethyl ether) recirculation stream, to produce one or more purified olefin streams; (e) using one or more reaction vessels, (f) A step of reacting one or more purified olefin streams with one or more oligomerization catalysts to produce a crude oligomerized olefin stream containing oligomerized olefins having a carbon length of 4 to 28 carbon atoms and contaminants containing one or more of hydrogen, methane, and alkanes lighter than pentane; and (f) fractional distillation of the crude oligomerized olefin stream by distillation to release hydrogen and light alkanes into (1) a fourth waste gas stream, (2) a recirculated stream of light oligomerized olefins, and (3) a purified oligomerized olefin stream containing olefins having a carbon length of 8 to 28 carbon atoms. (g) a step of removing, (h) a step of reacting the purified oligomerized olefin stream with hydrogen using a hydrogenation catalyst to saturate the olefin with paraffin and produce a crude jet fuel stream, (i) a step of fractionally distilling the crude jet fuel stream by distillation to remove excess hydrogen and light hydrocarbons into (1) a fifth waste gas stream, (2) a sixth waste gas stream containing naphtha, (3) a seventh waste gas stream containing diesel, and (4) a liquid fuel containing jet fuel, and (i) the following, (1) in a steam reforming process, an autothermal reforming process, or a dry reforming process,The present invention comprises (i) a step of generating a recirculating synthesis gas stream containing carbon oxides and hydrogen by reforming one or more waste gas streams with steam, or (ii) partially oxidizing one or more waste gas streams, and (j) a step of simultaneously supplying the recirculating synthesis gas stream to the methanol synthesis process of step A. Embodiments of the present invention are any or all of the embodiments described in the preceding paragraph to the second embodiment of this paragraph, further comprising the step of reacting one or more waste gas streams with oxygen in a partial oxidation reactor to generate a second recirculating synthesis gas stream containing CO and hydrogen. Embodiments of the present invention are any or all of the embodiments described in the preceding paragraph to the second embodiment of this paragraph, wherein the reaction of step (i) is carried out by reacting one or more waste gas streams with oxygen and steam in a self-heating reactor to generate a third recirculating synthesis gas stream containing CO and hydrogen. Embodiments of this disclosure are any or all of the earlier embodiments of this paragraph to the second embodiment of this paragraph, wherein the reaction of step (i) is carried out by reacting a sixth waste gas stream with steam in a steam reforming process. Embodiments of this disclosure are any or all of the earlier embodiments of this paragraph to the second embodiment of this paragraph, wherein the reaction of step (i) is carried out by reacting a seventh waste gas stream with steam in a steam reforming process. Embodiments of this disclosure are any or all of the earlier embodiments of this paragraph to the second embodiment of this paragraph, wherein the reaction of step (i) is carried out 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. Embodiments of this disclosure are any or all of the earlier embodiments of this paragraph to the second embodiment of this paragraph, wherein the hydrogen for steps (a) and / or (g) is produced by a water electrolysis unit. Embodiments of the present disclosure are any or all of the embodiments described in the preceding paragraph to the second embodiment described in this paragraph, wherein the reaction in step (i) is carried out by partial oxidation, and the oxygen for partial oxidation is produced by a water electrolysis unit. Embodiments of the present disclosure are wherein the oxygen for self-thermal reforming is produced by a water electrolysis unit.The embodiments of this disclosure are one, any, or all of the embodiments described in the preceding paragraph to the second embodiment of this paragraph. The embodiments of this disclosure are one, any, or all of the embodiments described in the preceding paragraph to the second embodiment of this paragraph, wherein the reaction of step (i) is carried out using steam in a steam reforming process, and one or more waste gas streams supplied to the steam reforming reactor contain about 10% to about 50% by weight of propane. The embodiments of this disclosure are one, any, or all of the embodiments described in the preceding paragraph to the second embodiment of this paragraph, wherein the reaction of step (i) is carried out using a dry reforming process.

[0156] 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.

[0157] 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 producing liquid fuel from carbon oxides and hydrogen, (a) A step of reacting a mixture of carbon oxide and hydrogen to produce a methanol stream, a water stream, and one or more first gaseous waste streams, (b) A step of contacting the methanol with an MTO catalyst to generate one or more olefin streams and a second gaseous waste stream, (c) A step of oligomerizing one or more olefin streams with one or more oligomerization catalysts to produce an oligomerized olefin stream and a third gaseous waste stream, (d) A step of hydrogenating the oligomerized olefin flow using a hydrogenation catalyst in the presence of hydrogen to produce a naphtha flow, a jet fuel flow, a diesel flow, and a fourth gaseous waste flow, (e) Below, (1) Partially oxidizing one or more of the naphtha flow, the diesel flow, the first gaseous waste flow, the second gaseous waste flow, the third gaseous waste flow, or the fourth gaseous waste flow, (2) A process comprising the step of generating a synthesis gas flow containing carbon oxides and hydrogen by reforming one or more of the naphtha flow, the diesel flow, the first gaseous waste flow, the second gaseous waste flow, the third gaseous waste flow, or the fourth gaseous waste flow using steam in a steam reforming reactor, a self-thermal reforming reactor, or a dry reforming reactor.

2. The process according to claim 1, wherein the synthesis gas flow is recycled to step (a).

3. The process according to claim 1, wherein the hydrogen in step (a) is green hydrogen.

4. The carbon oxides mentioned above are CO and CO ２ The process according to claim 1, including the process described in claim 1.

5. The synthesis gas flow in step (e) is approximately 0.5 to approximately 4 H ２ The process according to claim 4, having a CO ratio.

6. The process according to claim 1, wherein the methanol synthesis process, the MTO process, and / or the oligomerization process generate vapor, which is used in the reforming step (e).

7. The process according to claim 1, wherein step (a) is carried out in the presence of a catalyst.

8. The above step (a) is, f) Reacting a mixture of carbon oxide and hydrogen to produce methanol, water, and hydrogen, CO, CO ２ A step of generating a crude methanol stream containing methane, ethanol, and one or more other oxygenated hydrocarbons, g) A step of purifying the crude methanol stream by distillation to remove light contaminants into the first gaseous waste stream, and to remove heavy contaminants and water into the second heavy waste stream to produce a purified methanol stream, h) The process according to claim 1, further comprising the step of contacting the purified methanol stream with the MTO catalyst.

9. In step (b), the one or more olefin streams contain ethylene, propylene, and hydrogen, CO, CO ２ The process according to claim 1, comprising a crude olefin stream containing methane, dimethyl ether, ethanol, and one or more other oxygenated hydrocarbons.

10. The crude olefin stream is purified by distillation to remove light fractions into the third gaseous waste stream, followed by water absorption to remove heavy oxygenated hydrocarbons and water extraction to remove dimethyl ether, thereby generating one or more purified olefin streams. The process according to claim 9, further comprising the step of reacting one or more of the purified olefin streams with one or more oligomerization catalysts.

Images