Integrated process for converting carbon oxides to olefins

Abstract

An integrated process for producing light olefins by combining methanol purification and olefin production from methanol. The methanol supplied to the MTO process may be crude methanol or purified methanol from which heavy oxygenates in the MTO effluent have been removed and treated with heavy oxygenates. By eliminating or reducing the need for methanol purification by methanol suppliers, the cost of converting methanol to olefins or fuels is reduced.

Description

Technical Field

[0001] This field relates to an integrated process for producing light olefins from carbon oxides. This field may relate particularly to integrating a methanol synthesis process with an oxide conversion process.

Background Art

[0002] Olefins have traditionally been produced from petroleum feedstocks by catalytic or steam cracking processes. These cracking processes, particularly steam cracking, produce light olefins such as ethylene and propylene from various hydrocarbon feedstocks. Ethylene and propylene are important general-purpose petrochemicals useful in various processes for manufacturing plastics and other compounds.

[0003] In the petrochemical industry, it has long been known that oxides, particularly alcohols, can be converted into light olefins. For example, methanol, which is a preferred alcohol for light olefin production, can be converted mainly into ethylene and propylene in the presence of a molecular sieve catalyst. This process is called the methanol-to-olefin (MTO) reaction process and occurs in an MTO reaction system. A very efficient MTO process can convert oxides into light olefins that have typically been used in plastic production. The light olefins produced from the MTO process are concentrated in ethylene and propylene but also include C4 - C6 olefins.

[0004] Methanol is typically synthesized from the catalytic reaction of syngas in a methanol reactor in the presence of a catalyst. Syngas is defined as a gas mainly containing carbon monoxide (CO), hydrogen (H2), and preferably carbon dioxide (CO2). Other components may be present. Syngas production processes are well-known and include conventional steam reforming, autothermal reforming, dry reforming, or combinations thereof.

[0005] Methanol synthesis is known to produce numerous by-product impurities, including methane, dimethyl ether, methyl formate, higher alcohols, and ketones (e.g., acetone, methyl ethyl ketone). The amount and type of these impurities depend on the raw materials and methods used in methanol synthesis. Effective use of the product methanol generally requires capital and energy-intensive separation processes to remove both impurities more volatile than methanol and those less volatile than methanol. These separations negatively impact the overall economics of methanol production. Downstream use of methanol for the production of olefins, gasoline, jet fuel, and distillates produces the same by-product impurities that occur in methanol synthesis. Furthermore, other by-products such as acetaldehyde are also present. As with methanol synthesis, removing these impurities requires considerable capital and energy costs.

[0006] An effective means is needed to integrate the purification process for methanol synthesis with the product purification process for downstream methanol conversion. [Overview of the project]

[0007] The inventors have devised an integrated process for producing light olefins by integrating methanol purification with olefin production from methanol. The methanol supplied to the MTO process may be crude methanol or purified methanol from which heavy oxygenates have been removed and which has been treated with heavy oxygenates in the MTO effluent. By eliminating or reducing the need for methanol purification by methanol suppliers, the cost of converting methanol into olefins or fuels is reduced. [Brief explanation of the drawing]

[0008] [Figure 1] This is a schematic diagram of the methanol synthesis and oxygenate conversion process according to the present disclosure. [Figure 2] This is a schematic diagram of an MTO recovery process according to an exemplary embodiment of the integrated process for producing the light olefins of this disclosure. [Figure 3] This is a schematic diagram of the alternative methanol synthesis and oxygenate conversion process described herein.

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

[0010] 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.

[0011] 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.

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

[0013] The term "indirect communication" refers to the flow of fluid from an upstream component entering a downstream component after passing through an intervening container.

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

[0015] As used herein, the terms “main” or “major” mean more than 50%, preferably more than 75%, and more preferably more than 90%.

[0016] 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. Stripping columns may omit the reboiler at the bottom of the column, but instead may provide heating requirements and the propulsion for separation from fluidizing inert media such as steam. Stripping columns typically feed to an upper tray and remove the main product from the bottom.

[0017] 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 which may also have an aqueous outlet from the boot. A flash drum is a type of separator that may be downstream-communicated with a separator that may operate at higher pressures. As used herein, the term “boiling temperature” means the atmospheric equivalent boiling point (AEBP) calculated from the observed boiling point and distillation pressure using the formulas provided in ASTM D1160 Appendix A7, entitled “Practice for Converting Observed Vapor Temperatures to Atmospheric Equivalent Temperatures.”

[0018] As used herein, the term “True Boiling Point” (TBP) means a test method for determining the boiling point of a substance, which corresponds to ASTM D-2892 for producing standardized quality liquefied gas, distilled fractions, and residues from which analytical data can be obtained, and for determining the yield of the above fractions by both mass and volume, with a temperature versus mass % graph produced using 15 theoretical stages in a column with a reflux ratio of 5:1.

[0019] As used herein, the terms "T5", "T10", "T90", or "T95" mean the temperature at which 5 weight percent, 10 weight percent, 90 weight percent, or 95 weight percent, respectively, of a sample boils, depending on the case, using ASTM D-86 or TBP.

[0020] As used herein, the term "initial boiling point" (IBP) means the temperature at which a sample begins to boil, depending on the case, using ASTM D-7169, ASTM D-86, or TBP.

[0021] As used herein, the term "end point" (EP) means the temperature at which a sample has completely evaporated, depending on the case, as may be the case when using ASTM D-7169, ASTM D-86, or TBP.

[0022] As used herein, the term "diesel" means hydrocarbons that boil within the range of an IBP of about 125 °C (257 °F) to about 175 °C (347 °F) or a T5 of about 150 °C (302 °F) to about 200 °C (392 °F), and "diesel cut point" includes a T95 between about 343 °C (650 °F) and about 399 °C (750 °F) using the TBP distillation method, or a T90 between 280 °C (536 °F) and about 340 °C (644 °F) using ASTM D-86. The term "green diesel" means diesel containing hydrocarbons not of fossil fuel origin.

[0023] As used herein, the term "jet fuel" means hydrocarbons that boil within the range of a T10 of about 190 °C (374 °F) to about 215 °C (419 °F) and an end point of about 290 °C (554 °F) to about 310 °C (590 °F). The term "green jet fuel" means jet fuel containing hydrocarbons not of fossil fuel origin.

[0024] As used herein, the term "component-rich stream" means that the rich stream exiting the vessel has a higher concentration of the component than the feed to the vessel, preferably than all other streams withdrawn from the vessel.

[0025] As used herein, the term "component-lean stream" means that the lean stream exiting the vessel has a lower concentration of the component than the feed to the vessel, preferably than all other streams withdrawn from the vessel.

Embodiments for Carrying out the Invention

[0026] An integrated process and apparatus for producing methanol from carbon oxides, converting the methanol to light olefins, and possibly to fuel. In one embodiment, the crude methanol is fed directly to the MTO unit. This embodiment utilizes reacting DME and other oxygenates in the crude methanol to produce additional olefins. Only a portion of the heavy oxygenates in the methanol feed react in the MTO reactor. However, the product recovery section for the MTO reactor effluent can be used to recover the unreacted oxygenates. The major light impurity in the crude methanol is carbon dioxide. The additional carbon dioxide in the MTO reactor effluent increases the duty for downstream carbon dioxide removal in the caustic scrubber. To address the additional carbon dioxide, bulk carbon dioxide removal is used upstream of the caustic scrubber to remove the carbon dioxide and reduce the caustic consumption. An additional advantage is that the removal of bulk carbon dioxide provides additional carbon dioxide that can be utilized by other means such as the production of hydrogen by reforming.

[0027] In the second embodiment, the crude methanol fractionation column is used in the MTO unit. The energy requirements for operating these columns can be supplied by waste energy from the MTO unit. Furthermore, the purified methanol product is taken out as vapor at the top of the column, but instead of condensing it at the end of the purification process as in the conventional method and then evaporating it for the MTO reactor, it can be sent directly to the MTO reactor as feed. In the conventional MTO process, liquid methanol is vaporized before being supplied to the MTO unit, thereby consuming energy to vaporize the methanol. Additionally, the specifications for methanol distillation can be relaxed, thereby reducing the associated capital and energy consumption. Water containing heavy oxygenates can be treated together with the water condensed from the MTO reactor effluent.

[0028] Referring to Figure 1, the integrated process and apparatus 101 for producing light olefins includes a methanol synthesis section 111 and a methanol purification section 201. As shown in Figure 1, the synthesis gas stream in line 122 and the hydrogen gas stream in line 124 are fed to the methanol synthesis section 111. Synthesis gas is defined as a gas primarily consisting of carbon monoxide (CO), carbon dioxide (CO2), and hydrogen (H2). Optionally, synthesis gas may also contain methane (CH4), as well as small amounts of ethane and propane. Conventional processes for converting carbon components into synthesis gas include steam reforming, partial oxidation, autothermal reforming, and combinations of these processes. According to embodiments of this disclosure, the synthesis gas stream in line 122 can be taken from any suitable source. According to another embodiment of this disclosure, the hydrogen gas stream in line 124 can be taken from any suitable source. In an exemplary embodiment, the hydrogen gas stream in line 124 is taken from a pressure swing adsorption (PSA) unit.

[0029] According to exemplary embodiments of the present disclosure, the methanol synthesis section 111 comprises a first methanol converter 140 and a second methanol converter 160. The synthesis gas stream in line 122 and the hydrogen gas stream in line 124 are fed to the first methanol converter 140 of the methanol synthesis section 111. In embodiments, the synthesis gas stream in line 122 and the hydrogen gas stream in line 124 can be mixed to provide a mixed feed stream 126, which is then fed to the first methanol converter 140. However, the synthesis gas stream in line 122 and the hydrogen gas stream in line 124 may be fed to the first methanol converter 140 separately. The mixed feed stream 126 can be fed to a synthesis gas booster compressor 130 to compress the synthesis gas to a specific pressure and provide a compressed synthesis gas stream in line 132 before being fed to the first methanol converter 140. In an exemplary embodiment, the synthesis gas may be compressed in a synthesis gas booster compressor 130 to a pressure of approximately 6890 kPa (1000 psia) to approximately 8970 kPa (1300 psia). The synthesis gas flow may be heated before being sent to the first methanol converter 140. The compressed synthesis gas flow in line 132 may 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.

[0030] 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 toward the methanol synthesis section 111 has a dicarbonate to monocarbonate molar ratio of 1:2 to 1:4 and a hydrogen to carbonoxide (CO+CO2) molar ratio in the range of about 3:2 to about 3:1.

[0031] Suitable methanol synthesis catalysts may be zinc oxide and copper on an 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 into the synthesis reaction to obtain sufficient conversion.

[0032] 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).

[0033] The methanol synthesis reaction is highly exothermic. Boiler feedwater (BFW) in line 148 is sent to the first methanol converter 140, generating a steam flow in line 142 drawn from the first methanol converter 140. The generation of steam absorbs the heat generated in the methanol synthesis reaction. The steam flow in line 142 is sent to the top separator 145, which separates the steam in line 146 from the water flow in line 147. The water flow in line 147 is replenished by the recirculated BFW in line 149, providing BFW from line 148 for the first methanol converter 140.

[0034] 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 of 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 a heat exchanger 133. The heat-exchanged first reactor effluent in line 135 is cooled in a cooler 131 to provide the cooled first reactor effluent in line 136. The cooled first reactor effluent in line 136 is further cooled in a cooler 137 to provide the further cooled first reactor effluent in line 138. The further cooled first reactor effluent in line 138 is separated in a 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.

[0035] 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 mixed with the feed hydrogen flow in line 153 to provide the mixed first vapor flow in line 156. The mixed first vapor flow in line 156 is compressed in the first compressor 155 to supply compressed first vapor 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 the process and apparatus 101.

[0036] The compressed first vapor flow in line 157 is heat-exchanged with the second reactor outflow flow in line 162 in heat exchanger 163 to provide a heated 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. Suitable methanol synthesis catalysts may be zinc oxide and copper on an alumina support. The synthesis conditions in the second methanol converter 140 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.

[0037] The boiler feedwater (BFW) in line 176 is sent to the second methanol converter 160, which generates a steam flow in line 166 drawn from the second methanol converter 160, absorbing the heat. 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, providing BFW in line 176 for the second methanol converter 160.

[0038] 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 contain methanol, dimethyl ether, ethanol, or a combination thereof. The second reactor outflow in line 162 may be drawn from a side of the second methanol converter 160. The second reactor outflow in line 162 is cooled by heat exchange in a heat exchanger 163 with the compressed first vapor flow in line 157. The heat-exchanged second reactor outflow in line 164 may be cooled in a cooler 165 to provide a cooled and condensed second reactor outflow in line 166. The cooled second reactor outflow in line 166 is separated in a second gas-liquid separator 180 to provide a second vapor flow in line 182 and a second liquid flow in line 184. The second vapor stream in line 182 and the second liquid stream in line 184 can be further processed to recover methanol.

[0039] 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 under a pressure of approximately 6890 kPa (1000 psia) to approximately 8970 kPa (1300 psia).

[0040] 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 183. 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.

[0041] The PSA feed stream in line 183 is processed in PSA unit 185. Typically, a PSA unit includes a series of adsorption beds, each containing one or a combination of adsorbents suitable for adsorbing specific components to be adsorbed. These adsorbents include, but are not limited to, activated alumina, silica gel, activated carbon, zeolite molecular sieve-type materials, or any combination thereof. The adsorbents are organized in any order required by the adsorption process to adsorb impurities or components. In PSA unit 185, the PSA feed gas flows over the adsorbents, adsorbing impurities that are easily adsorbed during the adsorption process, while hydrogen passes through. Pressure swings allow the adsorbed impurities on the adsorbents to desorb into line 186. The purified hydrogen gas leaves the adsorption beds in the PSA top gas stream 124, where the impurities are lean.

[0042] In the PSA unit 185, hydrogen present in the PSA feed stream in line 183 is separated into a hydrogen-rich stream in line 124. As shown in the figure, the purge stream in line 186 from the PSA unit is separated from the hydrogen-rich stream in line 124. The purge stream in line 186 can be used as fuel. In an exemplary embodiment, the hydrogen-rich stream in line 124 can be sent as a hydrogen stream to the synthesis gas booster compressor 130. In this embodiment, the hydrogen stream in line 124 to the synthesis gas booster compressor 130 includes a hydrogen-rich stream.

[0043] Returning to the second gas-liquid separator 180, the second liquid flow in line 184 is drawn 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 mixed with the first liquid flow in line 154 to provide a mixed 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 contains at least 100 ppmw of carbon oxides and / or at least 100 ppmw of C 2+ It may contain oxygenated substances.

[0044] Crude methanol contains methanol, light fractions, and heavier alcohols. Where used and described herein, the terms “crude methanol” or “crude oxygenated feedstock” may include methanol, ethanol, water, light fractions, and fuel off. Light fractions may include ethers, ketones, aldehydes, and dissolved gases such as hydrogen, methane, carbon oxides, and nitrogen. Crude methanol contains fusel oil. Fusel oil in crude methanol typically contains higher alcohols and is commonly burned as fuel in methanol plants. Crude methanol containing fusel oil can be sent to an oxygenated conversion unit for further production of light olefins. According to this disclosure, crude methanol may be sent as feed to an oxygenated conversion unit or an MTO unit.

[0045] According to exemplary embodiments of the present disclosure, crude methanol may have a composition comprising carbon monoxide at a concentration of about 0 to about 1 wt%, carbon dioxide at a concentration of about 0.05 wt% to about 2 wt%, methane at a concentration of about 0.001 wt% to about 2 wt%, hydrogen at a concentration of about 0.05 wt% to about 2 wt%, oxygen at a concentration of about 0 to about 1 wt%, water at a concentration of about 5 wt% to about 18 wt%, nitrogen at a concentration of about 0 to about 1 wt%, methanol at a concentration of about 75 wt% to about 90 wt%, and a heavy alcohol having at least two carbon atoms at a concentration of about 0.05 to about 4 wt%.

[0046] 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 drawn from the crude methanol hold-up tank 195. According to the present invention, the crude methanol flow in line 196 may be sent to the oxygenate conversion unit 200, as shown in Figure 1. The crude methanol hold-up tank 195 is optional.

[0047] Conventionally, the crude methanol stream in line 196 is purified of light gases and heavy oxygenates before being fed into the MTO reactor 202 of the oxygenate conversion unit 200. According to the embodiment of Figure 1, at least 100 ppmw of carbon oxides and / or at least 100 ppmw of C 2+ The crude methanol stream containing oxygenated material is preheated and vaporized in heat exchanger 198, and then sent directly to MTO reactor 202. Heat exchanger 198 may include a series of heat exchangers that use waste heat from the process to preheat and vaporize the crude methanol stream in line 196.

[0048] The superheated crude methanol stream in line 199 is introduced into MTO reactor 202 and contacted with the MTO catalyst under MTO reaction conditions to convert methanol and other oxygenates into olefins and water. The crude methanol stream in line 198 may contain methanol, dimethyl ether, ethanol, or a combination thereof. MTO reactor 202 can fluidize the catalyst under high-speed flow conditions. The MTO catalyst may be a silicoaluminophosphate (SAPO) catalyst. SAPO catalysts and their formulations are generally taught in U.S. Publication Nos. 4,499,327(A), 10,358,394, and 10,384,986. MTO reaction conditions involve 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 hour ("WHSV") in MTO reactor 202 is about 1 to about 15 hours. -1 This is within the specified range. The MTO catalyst is separated from the olefin product stream after the MTO reaction.

[0049] In the MTO process, catalyst particles are repeatedly circulated between the MTO reactor 202 and the MTO regenerator unit 204. During regeneration, coke deposited on the catalyst particles during the reaction in the reaction zone is removed at high temperatures by oxidation in the regenerator unit 204. Removal of the coke deposits restores the activity of the catalyst particles to a point where they can be reused in the MTO reactor 202. The regenerated catalyst is discharged from the regenerator unit 204 in line 206 and recirculated back to the MTO reactor 202. The MTO effluent, including ethylene, propylene, and other olefins along with water and oxygenated materials, is discharged from the MTO reactor 202 in effluent line 207.

[0050] Figure 1 also shows a flue gas recovery process 250 for recovering heat from the high-temperature flue gas in line 209 discharged from the catalyst regenerator 204. The flue gas is very hot and can be used to recover heat from the catalyst regeneration process. The flue gas in line 209 contains catalyst particles that can be removed in a third-stage separator or fine particle removal unit 252, which includes a filter. The flue gas with fine particles removed to a very low concentration may contain carbon monoxide that can be burned into carbon dioxide in the combustor 254. Fuel gas in line 256 or oxygenated material in line 258 can be supplied to the combustor 254 to further increase the temperature of the flue gas. In addition, additional oxygen can be added to line 257 as air or another oxygen-containing stream. The extremely hot flue gas can be sent to a steam generator 260, where high-pressure steam can be generated by heat exchange between water and the flue gas stream. Furthermore, kinetic energy can be recovered from the flue gas stream by passing it through a turbine 262 to provide work. The flue gas recovery process can also be used in the embodiment shown in Figure 3.

[0051] Line 207 transports the MTO products in the MTO effluent to the MTO product unit 210 shown in Figure 2. The MTO product unit 210 can separate light olefins from the oxygenated material for further value extraction. The MTO product unit 210 includes a separation section 21 containing a DME stripper column 350, an extraction distillation column 360, a water stripper column 30, a quenching column 20, and a product separation column 24.

[0052] The high-temperature steam MTO effluent flow in line 207 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 steam reactor effluent is removed by direct contact with a water flow supplied in line 19, which may be taken in from the hot water-rich flow in line 47, thereby removing superheat, neutralizing organic acids, and removing catalyst particles. Furthermore, the circulating water flow in the quenching tower system is used in multiple stages to enhance the recovery of catalyst particles. An additional section in the quenching tower 20 may be provided for caustic injection to remove organic acids such as acetic acid and entrained caustic from the caustic contact section. The quenched olefin flow in line 22 is discharged from the quenching tower 20 and supplied to the product separator tower 24 in the separation section 21.

[0053] The product separator column 24 includes two sections for separating the reactor effluent into a product olefin flow at the top line 40, an intermediate liquid flow in the intermediate line 28, and a water flow in the bottom line 25. The water flow in the bottom line 25 can be separated into a product water flow and a recirculated product 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 column 24 in the recirculated product water flow to cool the quenched reactor effluent in line 22.

[0054] According to embodiments of the present disclosure, the product water flow in line 26 is sent to the water stripper tower 30. The water return flow containing oxygenated by-products from the compression section 80 in the return line 32 may also be passed to the water stripper tower 30. The confluence bottom line 34 can transport the water flow in line 34 from the confluence 29 to the water stripper tower 30. Finally, in the embodiment of Figure 3, the water oxygenated flow in the net oxygenated bottom line 226a may also be delivered to the water stripper tower 30. The water stripper tower 30 may be downstream in communication with the product separator tower 24, the confluence 29, and the compression section 80. The water stripper feed of line 36 can send the flows of lines 26, 32, 34, and 226a to the water stripper tower 30.

[0055] The vapor flow from the first section 24a of the product separator column 24 is sent to the second or upper section 24b of the product separator. The intermediate flow in line 28, containing hydrocarbons, oxygenated by-products, and liquid-phase water, is drawn out at the bottom of the upper section 24b. 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 through a conjugate 29 to separate the hydrocarbon top flow from the aqueous flow in line 34, and this is returned to the product aqueous flow in line 26 and pumped to the water stripper column 30. The top product olefin flow in line 40, containing olefins from the second section 24b of the product separator column 24, is delivered to the compression section 80. According to an exemplary embodiment, the product aqueous flow in line 26, the aqueous flow in line 34, and the water return flow in line 32 are mixed to provide a mixed product aqueous flow in line 36. The mixed product water stream in line 36 is sent to the water stripper tower 30. Alternatively, the product water stream from line 26, the aqueous stream from line 34, and the water return stream from line 32 can be sent separately to the water stripper tower 30. In the alternative embodiment shown in Figure 3, the net oxygenated bottom line 226a can deliver an aqueous oxygenated stream to the mixed product water stream 36 for processing.

[0056] The mixed product water stream in line 36 contains dilute hydrocarbon oxygenates such as DME, methanol, acetaldehyde, acetone, and MEK, and some heavy oxygenates containing two or more carbon molecules. The water stripper tower 30 separates or strips the oxygenates into a methanol and oxygenate-rich stream in the methanol-rich top line 44, i.e., a stream rich in both methanol and at least one other oxygenate containing heavy oxygenates, and a water-rich stream in the bottom line 46. In one embodiment, the temperature of the water stripper tower 30 may be about 115°C (239°F) to about 180°C (356°F) at the bottom of the water stripper tower, and the pressure may be about 75 kPa gauge (11 psig) to about 760 kPa (110 psig) at the top of the water stripper tower 30.

[0057] The mixed product water stream in line 36, which includes the product water stream in line 26, contains dilute hydrocarbon oxygenates such as DME, methanol, acetaldehyde, acetone, and MEK. The water stripper tower 30 separates or strips the oxygenates into a methanol and oxygenate-rich stream in methanol-rich top line 44, i.e., a stream rich in methanol and at least another oxygenate and heavy oxygenate, and a water-rich stream in bottom line 46.

[0058] A portion of the water-rich flow in bottom line 46 is re-boiled and returned to the water stripping column 30. The remainder of the water-rich flow in line 46 is cooled in a condenser to provide the net water-rich flow in line 49. The net water-rich flow in bottom line 49 can be divided into an extractant flow supplied to the extraction distillation column 360 in line 362 and a bottom water-rich recirculation flow in the remaining bottom line 43. The water-rich recirculation flow in line 43 is mixed with the flow in line 374 to give water flow 47. The water flow in line 47 can be supplied to the quenching column 20 in line 19 and the oxygenate absorber in line 102.

[0059] Uncondensed light hydrocarbons can be purged from the receiver top line 41 of the water stripper column 30, while the hydrocarbon lean, methanol and oxygenate-rich flow can be removed at the bottom line 48, and includes methanol, DME, acetaldehyde, acetone, MEK, and heavy oxygenates. A portion of the hydrocarbon lean methanol and oxygenate-rich flow can be returned to the water stripper column 30 as reflux. In one embodiment, the temperature of the water stripper column 30 may be about 115°C (239°F) to about 150°C (302°F) at the bottom of the water stripper column, and the pressure may be about 75 kPa gauge (11 psig) to about 345 kPa (50 psig) at the top of the water stripper column.

[0060] The hydrocarbon lean, methanol, and oxygenate-rich stream can be fed to the extraction distillation column 360 to separate methanol from at least one other oxygenate. However, the hydrocarbon lean methanol and oxygenate-rich stream contains DME, which is readily separated from methanol. Therefore, the hydrocarbon lean, methanol, and oxygenate-rich stream in line 48 can be fed to the DME stripper column 350 to easily remove the DME. The DME stripper column 350 may be downstream-communicated with the water stripper column 30. The DME stripper column 350 can separate or strip the DME into the DME-rich stream in the top line 352, providing the DME lean, methanol, and oxygenate-rich stream into the bottom line 354. The DME-rich stream in the top line 352 may be recycled to the oxygenate conversion section 200 as MTO feed. A portion of the DME lean, methanol, and oxygenate-rich stream can be re-boiled and recycled to the DME stripper column 350. The net DME lean methanol and oxygenate-rich flow in the bottom line 354 can be supplied to the extraction distillation column 360. The extraction distillation column 360 is downstream-communicated with the water stripper column 30 and upstream of the communication with the product separator column 24, ensuring that inert oxygenates do not accumulate in the compression section without a return path to the water stripper column 30. Furthermore, in the embodiment, the extraction distillation column may be downstream-communicated with the DME stripper column 350.

[0061] In one embodiment, the temperature of the DME stripper column 350 may be about 85°C (185°F) to about 120°C (248°F) at the bottom of the DME stripper column, and the pressure may be about 75 kPa gauge (11 psig) to about 414 kPa (60 psig) at the top of the column. The DME stripper column 350 can remove light hydrocarbon purge by utilizing a top condenser and receiver separator in addition to or instead of the top condenser and receiver 45 for the water stripper column 30.

[0062] The DME lean, methanol and oxygenate-rich flow in the DME stripper of the net bottom line 354 is fed into the extractive distillation column 360 to separate methanol from at least one other hydrocarbon oxygenate, preferably all other hydrocarbon oxygenates. The extractant flow of water may also be fed into the extractive distillation column 360 at a position such as the upper quarter of the column, above a position such as the middle quarter of the column, from a position such as the middle quarter of the column into which the DME lean, methanol and oxygenate-rich flow is fed. The extractant flow may be supplied to a line 362, which can be taken out from the water-rich flow in the water stripper bottom line 49.

[0063] The flow rate of the water extractant stream into the extraction distillation column 360 should be approximately 1.5 to 3 times the flow rate of hydrocarbon oxygenates into the extraction distillation column 360 in the DME lean, and approximately 1 to 3 times the flow rate of the methanol and oxygenate-rich stream in the DME stripper bottom line 354, as well as the overall flow rate of the DME, lean, methanol, and oxygenate-rich streams in the bottom line 354, including substantial water.

[0064] The extraction distillation column 360 produces an oxygenate-rich stream containing at least one other hydrocarbon oxygenate, such as acetone, acetaldehyde, MEK, and DME, into the top line 364, and a methanol and water-rich extract stream into the bottom line 366. A portion of the methanol and water-rich stream in the bottom line 366 can be re-boiled and returned to the extraction distillation column 360. The oxygenate-rich stream in the top line 364 can be cooled, partially condensed, and fed to the receiver separator 365. Uncondensed light hydrocarbons can be purged from the receiver top line, while the hydrocarbon lean oxygenate-rich stream, containing heavy oxygenates, DME, acetaldehyde, acetone, and MEK, can be removed in the receiver bottom line 368. The heavy oxygenates fed to the MTO reactor 202 in the crude methanol stream are concentrated into the oxygenate-rich stream in line 368. The oxygenated flow in line 368 can be fed to the combustor 360 in Figure 1 via line 258. A portion of the lean hydrocarbon oxygenated flow can be returned to the extraction distillation column 360 as reflux at a point above where the extractant flow is added to the extraction distillation column 360. A light hydrocarbon purge may be fed to the light olefin recovery. The oxygenated material in line 368 can be combusted to generate heat or steam that can be used elsewhere in the process. Alternatively, the oxygenated material in line 368 may be recovered for further value extraction or to provide a feedstock to the MTO reactor 202. Heavy oxygenated material not removed from the crude methanol flow is also recovered here for combustion.

[0065] At least 99% by weight, preferably at least 99.5% by weight, of hydrocarbon oxygenates other than methanol supplied to the extraction distillation column 360 can be recovered in the oxygenate-rich flow in the top line 364 of the extraction distillation column 360 and in the hydrocarbon lean oxygenate-rich flow in the bottom line 368 of the extraction receiver 365. At least 90% by weight, preferably at least 95% by weight, of methanol can be recovered in the methanol and water-rich flow in the net bottom line 366.

[0066] The extraction distillation column 360 may have operating conditions including a bottom temperature in the range of approximately 75°C (167°F) to approximately 150°C (302°F) and a top pressure in the range of approximately 75 kPa gauge (11 psig) to approximately 200 kPa gauge (29 psig). The extraction distillation column 360 may be downstream-communicated with the top line 44 and bottom line 46 of the water stripper column 30.

[0067] The recovered methanol is an MTO reactant that can be recycled to the MTO reactor 202, but it is undesirable to recycle water with methanol. Therefore, the methanol and water-rich flow in the net bottom line 366 can be fed to the methanol stripper column 370 to separate the methanol-rich flow in the top line 372 from the final water-rich flow in the bottom line 374. The methanol-rich flow in the top line 372 can then be recycled to the MTO reactor 202 without any inert oxygenates that would otherwise not react and accumulate in the MTO product unit 210. A portion of the final water-rich flow in the bottom line 374 can be re-boiled and recycled to the methanol stripper column 370. The final water-rich flow in the net bottom line 374 can be discharged from the water stripper bottom line 46 along with the unrecycled portion of the water-rich flow in the remaining bottom line 49 to provide a water flow 47. A portion of the flow 47 may be drawn out and sent to the wastewater treatment of line 51.

[0068] The product olefin flow in the product top line 40 carries valuable olefin products that need to be recovered. The compression section 80 increases the pressure of the product olefin flow 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 flow into a pressurized first olefin-rich flow 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 oxygenate-rich first aqueous flow in the bottom line 84. The olefin-rich flow 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 is pumped through the manifold line 76 to the return line 32, and the aqueous flow, along with the product aqueous flow in the mixed product aqueous flow in line 36, is returned to the water stripper tower 30.

[0069] The compression section 80 may include a second knockout drum 86 that separates the pressurized first olefin-rich flow into a second pressurized olefin-rich 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 an oxygenate-rich second aqueous flow in the bottom line 88. The second olefin-rich 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 through the manifold line 76 to the return line 32, where the aqueous flow is returned to the water stripper column 30 along with the product aqueous flow in the mixed product aqueous flow in line 36.

[0070] The compression section 80 may include a third knockout drum 90 that separates a pressurized second olefin-rich flow into a third pressurized olefin-rich flow in the top line 91 and an oxygenate-rich third aqueous flow in the bottom line 92. The third olefin-rich flow in the top line 91 may be supplied to the oxygenate absorption column 50. The aqueous flow in the bottom line 92 is pumped through the manifold line 76 to the return line 32, and the aqueous flow is returned to the water stripper column 30 along with the product aqueous flow in the mixed product aqueous flow in line 36.

[0071] 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.

[0072] As shown in Figure 2, in a preferred embodiment, at least a portion of the compression product flow through the top line 91 is brought into contact in the oxygenation absorption tower 50 with a cooled, water-free lean water flow directly taken in from the product separation tower 24, under conditions effective for absorbing the oxygenation without prior removal of it. In an exemplary embodiment, the absorbent flow of line 102, taken from the water-rich flow of line 47, may be sent to the oxygenation absorption tower 50. The contact in the oxygenation absorption tower 50 generates an absorption olefin-rich flow in the top line 54 and an absorption water-rich flow containing a certain amount of effluent oxygenation in the bottom line 52. The operating conditions of the oxygenation absorption tower may include 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).

[0073] 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 oxygenated material 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, mixed with the flow in the stripper top line 71, partially condensed by cooling in the heat exchanger 64, and supplied to the stripper separator 66 in line 65. The stripper separator 66 separates the aqueous flow containing oxygenated material in the boot of line 67 that supplies to the manifold line 76, the vaporized light olefin flow in the top line 68 containing C3-olefins, and the heavy olefin liquid flow containing C4+ olefins in line 69. The heavy olefin liquid flow in line 69 is stripped in the product DME stripper column 70 to remove C3- and lower vapors in the stripper column top line 71 from the heavy olefin liquid flow in the stripper bottom line 93. Most of the oxygenated material is stripped into the stripper column top line 71 and separated when it is recirculated to the stripper separator 66 after cooling. The stripper separator 66 may operate at a temperature of approximately 30°C (86°F) to approximately 60°C (140°F) and a pressure of approximately 1.7 MPa (g) (250 psig) to approximately 2.1 MPa (g) (300 psig).

[0074] The vaporized light olefin stream in the top line 68 is scrubbed conventionally in the caustic scrubber column. However, since the crude methanol stream in line 194 was not purified, carbon dioxide remained in the stream in amounts too large for the caustic scrubber to handle. Therefore, we propose removing most of the carbon dioxide in the amine absorption column 302.

[0075] The evaporating light olefin stream in the top line 68 of the stripper separator 66 can pass through a tray or packed bulk scrubbing column 302, where it is scrubbed with a bulk solvent, such as an aqueous solution, supplied by a scrubbing liquid line 304, to extract and remove acidic gases, including carbon dioxide, into the aqueous solution. Preferred bulk solvents include Selexol®, available from UOP LLC in Des Plaines, Illinois, and amines such as alkanolamines including diethanolamine (DEA), monoethanolamine (MEA), methyl diethanolamine (MDEA), diisopropanolamine (DIPA), and diglycolamine (DGA). Other bulk solvents can be used in place of or in addition to the preferred amines. The lean bulk solvent comes into contact with the vaporizing light olefin stream in the top line 68 of the stripper separator column and absorbs acidic gas contaminants such as carbon dioxide. The resulting lean light olefin flow is removed from the top outlet of the bulk scrubbing column 302 in the bulk scrubber top line 306, and the rich bulk solvent is removed from the bottom at the bottom outlet of the bulk scrubbing column 302 in the bulk scrubbing bottom line 308. The spent bulk solvent from the bottom may be regenerated and recycled back into the bulk scrubbing column 302 in the scrubbing liquid line 304. The spent bulk scrubbing solvent can be regenerated from carbon dioxide capture from flue gas from combustion heaters in the plant, or from the MTO regenerator 204 along with other used solvents. The lean light olefin flow may exit the bulk scrubber column 302 via the bulk scrubber top line 306 and be supplied to the caustic scrubber column 73.

[0076] The bulk scrubbing column 304 can operate with a gas inlet temperature of approximately 38°C (100°F) to approximately 66°C (150°F) and a top pressure of approximately 3 MPa (gauge) (435 psig) to approximately 20 MPa (gauge) (2900 psig). Preferably, the bulk scrubbing column 104 can operate at temperatures of approximately 40°C (104°F) to approximately 125°C (257°F) and pressures of approximately 1200 to approximately 1600 kPa. The temperature of the vaporized light olefin flow in the top line 68 to the bulk scrubbing column 304 may be approximately 20°C (68°F) to approximately 80°C (176°F), and the temperature of the bulk solvent flow in the scrubbing liquid line 302 may be approximately 20°C (68°F) to approximately 70°C (158°F).

[0077] The lean light olefin flow in the bulk scrubber top line 306 then scrubs in the caustic scrubber column 73 through countercurrent contact with the caustic solution in line 42, absorbing residual acidic gases such as carbon dioxide from the vaporized light olefin product flow, which exits the caustic scrubber 73 at the top line 74. The acidic gas-rich caustic solution exits the scrubber 73 at line 44 and is supplied to the water stripper manifold 76.

[0078] The scrubbed light olefin vapor in the top line 74 may be cooled with a propylene refrigerant in a cryogenic cooler 75 to liquefy a portion of the light olefin product flow, separated in a dry separator 46 to provide an aqueous flow from the boot, and in the manifold line 76 and the top line 77. 2- A vapor stream of light olefin products containing hydrocarbons and gases, and C 3+ The vaporized light olefin product stream in the bottom line 78, which contains hydrocarbons, is led to the vaporized light olefin product stream in the top line 77. The vaporized light olefin product stream in the top line 77 is dried in the dryer 79a to provide the vaporized product olefin product stream in line 112. The liquidized light olefin product stream in the bottom line 78 is dried in the dryer 79b to provide the liquid product olefin product stream in line 114. The product olefin product streams in lines 112 and 114 can be drawn out and further processed.

[0079] According to another embodiment of the present disclosure, an integrated process and apparatus for producing light olefins comprises a methanol synthesis unit 101 having a methanol purification section 208, as shown in Figure 3. Many of the elements in Figure 3 have the same configuration and the same reference numerals as in Figure 2. Elements in Figure 3 that correspond to elements in Figure 2 but have a different configuration have the same reference numerals as in Figure 2, but are denoted with a prime symbol (').

[0080] According to one embodiment, the crude methanol stream in line 196 can be sent to the methanol purification section 208 to separate by-products and / or trace components and provide a methanol product stream for the oxygenate conversion unit 200.

[0081] According to an exemplary embodiment, the crude methanol stream in line 196 may be sent to a methanol purification section 210, which includes at least two distillation columns, namely a first distillation column 210 and a second distillation column 220. 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 stream in line 214 may be sent to a fuel section or possibly used as fuel in a combustor 254 in line 256 in Figure 1. From the first top receiver 215, the top receiver liquid flow is drawn out in line 216 and sent to the top of the first distillation column 210.

[0082] The bottom flow of the first distillation column containing methanol in line 218 is drawn out for further separation. The bottom flow of the first distillation column in line 218 is separated into a first reboiling flow in line 218b and a first distillation column outflow flow in line 218a. The first reboiling flow in line 218b is reboiled in reboiler 219 before being sent to the first distillation bottom. According to an exemplary embodiment, the first distillation column 210 is operated at a pressure of about 172 kPa (25 psia) to about 1379 kPa (200 psia). According to another exemplary embodiment, the first distillation column is operated at a temperature of about -17°C (0°F) to about 177°C (350°F).

[0083] The first distillation column outflow in line 218a is C, which should be removed from the crude methanol stream. 2+ It contains heavy oxygenated substances such as alcohols, ketones, and aldehydes. 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 bottom flow in line 226. The second distillation column top is in the gas phase. Instead of condensing it, a methanol feed stream is taken in from the second distillation column top line 222 and supplied to the MTO reactor 202 in line 199'. Using a superheater (not shown), the temperature of the methanol feed stream in line 199' can be raised before it fills the MTO reactor 202. A portion of the top flow in line 222 may be condensed in line 228 for reflux into the column. From the heat exchanger 223, the partially condensed top flow of the second distillation column in line 224 is sent to the second top receiver 225. In the second top receiver 225, the condensed portion of the top flow of the second distillation column in line 224 is recirculated to the second distillation column 220.

[0084] The second distillation column bottom flow in line 226 is drawn out of the column. The second distillation column bottom flow in line 226 is separated into a second reboiling flow in line 226b and a 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. According to an exemplary embodiment, the second distillation column operates at a pressure of about 3 kPa (5 psia) to about 862 kPa (125 psia). According to a further exemplary embodiment, the second distillation column operates at a temperature of about 38°C (100°F) to about 149°C (300°F). The second distillation column outflow flow in line 226a includes heavy oxygenated material and water, and aqueous oxygenated material.

[0085] In the embodiment shown in Figure 1, heavy oxygenates in the crude methanol stream are processed in the MTO outflow stream at line 207 and separated into an oxygenate-rich stream at line 368. In the embodiment shown in Figure 3, the heavy oxygenates are concentrated in the aqueous oxygenation stream in the net second distillation column bottom line 226a. The aqueous oxygenation stream of line 226a is transported to line 36 and stripped in the water stripper column 30, where the heavy oxygenates are stripped from the water, for example, from the product water stream of line 26, to provide an oxygenation stream, and are also concentrated to become the rich oxygenation stream of line 368.

[0086] According to embodiments of the present disclosure, the methanol purification section 201 may also include a third distillation column (not shown) for further removal of heavy oxygenates from the crude methanol stream. According to exemplary embodiments, the third distillation column may operate at a pressure of about 35 kPa (5 psia) to about 345 kPa (50 psia). According to another exemplary embodiment of the present disclosure, the third distillation column may be an atmospheric column operating at approximately atmospheric pressure. According to yet another exemplary embodiment, the second distillation column operates at a temperature of about 38°C (100°F) to about 122°C (250°F).

[0087] If a third column is also used, the outflow from the second distillation column in line 226a is separated in the third distillation column to provide a top flow and bottom flow containing methanol. The top flow, along with the methanol product flow from line 222 in MTO feed line 199', can be sent to the MTO reactor 202.

[0088] 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.

[0089] A first embodiment of the present invention is a carbon oxide of at least 100 ppmw or at least 100 ppmw C 2+A process for producing olefins from carbon oxides, comprising providing a crude methanol stream containing oxygenated material, and loading the crude methanol stream into an MTO reactor to convert methanol to olefins and produce an MTO effluent stream. Embodiments of the present invention are one or all of the preceding embodiments in this paragraph to the first embodiment in this paragraph, further comprising converting carbon oxides into a crude methanol stream. Embodiments of the present disclosure are one or all of the preceding embodiments in this paragraph to the first embodiment in this paragraph, further comprising separating the MTO effluent stream into a heavy oxygenated stream and a product olefin stream. Embodiments of the present invention are one or all of the preceding embodiments in this paragraph to the first embodiment in this paragraph, further comprising burning the heavy oxygenated stream. Embodiments of the present invention are one or all of the preceding embodiments in this paragraph to the first embodiment in this paragraph, further comprising burning the heavy oxygenated stream in a CO combustor. Embodiments of the present invention are any or all of the embodiments from the preceding embodiments in this paragraph to the first embodiment in this paragraph, further comprising burning a heavy oxygenated flow with flue gas from an MTO regenerator in a CO boiler. Embodiments of the present disclosure are any or all of the embodiments from the preceding embodiments in this paragraph to the first embodiment in this paragraph, further comprising absorbing carbon dioxide from a light olefin flow into a bulk solvent to provide a lean light olefin flow. Embodiments of the present disclosure are any or all of the embodiments from the preceding embodiments in this paragraph to the first embodiment in this paragraph, further comprising absorbing carbon dioxide from a lean light olefin flow into a caustic flow to provide a light olefin product flow. Embodiments of the present invention are any or all of the embodiments from the preceding embodiments in this paragraph to the first embodiment in this paragraph, further comprising separating aqueous oxygenated material from a crude methanol flow to provide an aqueous oxygenated flow before filling the crude methanol flow into an MTO reactor.Embodiments of the present invention are any or all of the embodiments described in the preceding paragraph to the first embodiment described in this paragraph, further comprising: rapidly cooling an MTO outflow logistics to provide a rapidly cooled olefin logistics; separating the rapidly cooled olefin logistics to provide a product olefin logistics and a product aqueous logistics; and stripping oxygenated materials from the product aqueous logistics and aqueous oxygenated logistics to provide an oxygenated logistics. Embodiments of the present invention are any or all of the embodiments described in the preceding paragraph to the first embodiment described in this paragraph, further comprising separating a heavy oxygenated logistics from the oxygenated logistics and burning it.

[0090] A second embodiment of the present invention is a carbon oxide of at least 100 ppmw or C 2+A process for producing olefins from carbon oxides, comprising: providing a crude methanol stream containing oxygenated material; separating light gases from the crude methanol stream to provide an oxygenated methanol stream; separating aqueous oxygenated material from the oxygenated methanol stream to provide a methanol stream; and loading the methanol stream into an MTO reactor to convert methanol to olefins and produce an MTO effluent stream. Embodiments of the present invention are any or all of the preceding embodiments in this paragraph to the second embodiments in this paragraph, further comprising converting carbon oxides into a crude methanol stream. Embodiments of the present invention are any or all of the preceding embodiments in this paragraph to the second embodiments in this paragraph, further comprising separating aqueous oxygenated material from the oxygenated methanol stream to provide a vapor methanol stream; and loading the vapor methanol stream into an MTO reactor. Embodiments of the present invention are any or all of the embodiments from the preceding to the second embodiments in this paragraph, further comprising: rapidly cooling the MTO outflow logistics to provide a rapidly cooled olefin flow; separating the rapidly cooled olefin flow to provide a product olefin flow and a product water flow; and stripping oxygenated materials from the product water flow and aqueous oxygenated logistics to provide an oxygenated logistics. Embodiments of the present invention are any or all of the embodiments from the preceding to the second embodiments in this paragraph, further comprising: separating the heavy oxygenated logistics from the oxygenated logistics and burning it. Embodiments of the present invention are any or all of the embodiments from the preceding to the second embodiments in this paragraph, further comprising: burning the heavy oxygenated logistics in a CO combustor. Embodiments of the present invention are any or all of the embodiments from the preceding to the second embodiments in this paragraph, further comprising: burning the heavy oxygenated logistics with flue gas from an MTO regenerator in a CO boiler.

[0091] Without further detail, it is expected that a person skilled in the art can use the foregoing description to the fullest extent without departing from the spirit and scope of the Disclosure, readily identify the essential characteristics of the Disclosure, and make various changes and modifications to the Disclosure to suit various uses and conditions. Accordingly, the prior preferred specific embodiments should be interpreted as merely illustrative examples and not in any way limiting the remainder of the Disclosure, but are intended to cover various modifications and equivalent configurations that fall within the scope of the appended claims.

[0092] 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 olefins from carbon oxides, At least 100 ppmw of carbon oxide or at least 100 ppmw of C ２＋ To provide a crude methanol stream containing oxygenated material, and A process comprising filling an MTO reactor with the crude methanol stream to convert methanol to olefins and generate an MTO effluent stream.

2. The process according to claim 1, further comprising converting the carbon oxide into the crude methanol stream.

3. The process according to claim 1, further comprising separating the MTO outflow into a heavy oxygenated flow and a product olefin flow.

4. The process according to claim 3, further comprising burning the heavy oxygenated material.

5. The process according to claim 4, further comprising burning the heavy oxygenated material in a CO combustor.

6. The process according to claim 5, further comprising burning the heavy oxygenated logistics together with flue gas from an MTO regenerator in a CO boiler.

7. The process according to claim 1, further comprising absorbing carbon dioxide from the light olefin flow into a bulk solvent to provide a lean light olefin flow.

8. The process according to claim 7, further comprising absorbing carbon dioxide from the lean light olefin flow into a caustic flow to provide a light olefin product flow.

9. The process according to claim 1, further comprising separating aqueous oxygenated material from the crude methanol stream to provide an aqueous oxygenated stream before filling the MTO reactor with the crude methanol stream.

10. The process according to claim 9, further comprising: rapidly cooling the MTO outflow logistics to provide a rapidly cooled olefin flow; separating the rapidly cooled olefin flow to provide a product olefin flow and a product aqueous flow; and stripping oxygenated materials from the product aqueous flow and the aqueous oxygenated logistics to provide an oxygenated logistics.

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