Device and method for preparing light olefins from alcohol

Abstract

The present invention belongs to the field of light olefins and relates to a device and method for producing light olefins from alcohol. The device comprises a first reaction system, a quenching system, a compression system, a separation system, a hydrogenation system, and a second reaction system connected in sequence; the discharge pipeline of carbon four and above heavy hydrocarbons of the separation system is connected to the hydrogenation system; the discharge pipeline of the hydrogenation system is connected to the second reaction system; the second reaction system comprises an intermediate product preheater / preheater group, an intermediate product reaction facility, and an optional cracking gas quencher / quencher group arranged in sequence; the discharge pipeline of the second reaction system is connected to the quenching system. Compared with the existing alcohol-to-light olefin process of the same scale, the device and method of the present invention can reduce the types of by-products and increase the production of ethylene and propylene. The present invention can be directly applied to newly built alcohol-to-light olefins devices, and can also be applied to the capacity expansion and transformation of existing methanol-to-light olefins systems with a small amount of additional investment.

Description

Technical Field

[0001] The present invention belongs to the field of low-carbon olefins, and specifically relates to a device and method for producing olefins from alcohols. The device and method of the present invention can reduce the types of by-products, lower the raw material consumption per unit olefin, and achieve increased production of ethylene and propylene. Background Art

[0002] With economic development, demand for light olefins is increasing. Simultaneously, due to the dramatic fluctuations in oil prices, non-hydrocarbon cracking technologies are rapidly developing. Furthermore, with the rapid increase in shale gas production and the maturation of coal-to-syngas technology, the application of methanol-to-olefins (MTO) or dimethyl ether-to-olefins (DTO) technologies, which produce methanol or dimethyl ether from natural gas or coal-to-syngas, has attracted widespread attention. MTO and DTO can produce light olefins with high selectivity and can flexibly adjust the propylene / ethylene ratio over a wide range.

[0003] Prior art technology uses a dedicated fluidized bed reactor and regenerator. The raw methanol enters the fluidized bed reactor and is dehydrated to produce a product gas containing ethylene and propylene. The product gas is cooled and compressed before entering a subsequent separation system to separate the product ethylene and propylene, as well as the byproducts fuel gas, ethane, propane, C4, and C5+.

[0004] The second existing technology uses a dedicated fluidized bed reactor and regenerator to convert the raw methanol into a product gas containing ethylene and propylene. The product gas is cooled and then sent out. This technology does not have a compression and subsequent separation system.

[0005] Existing technology three utilizes proprietary catalysts and fluidized bed reactors to convert methanol into product gas. The product gas undergoes compression, pre-deethanization, demethanization, C4 absorption and other processes to obtain product ethylene, propylene and by-product fuel gas, ethane, propane, C4 and C5+.

[0006] The fourth existing technology uses a fixed bed reactor to convert olefins in C4 and above raw materials into propylene. This method has the following problems: 1. Frequent bed regeneration; 2. Low ethylene:propylene ratio in the reaction product; 3. Alkanes do not participate in the reaction, and the raw material utilization rate is low.

[0007] Therefore, the by-products of C4 and above heavy hydrocarbons in the existing alcohol-to-olefins process are not reasonably utilized. Summary of the Invention

[0008] The inventors of the present invention have found that the C4 and above heavy hydrocarbons in the existing alcohol-to-olefins process are all sent out as by-products (such as Figure 1The present invention provides an apparatus and method for producing light olefins from alcohols. By using this method, heavy hydrocarbons with a carbon content of four or more are no longer by-products but serve as intermediates, thereby increasing the production of ethylene and propylene and improving the ethylene:propylene ratio in the product.

[0009] A first aspect of the present invention provides an alcohol-to-low-carbon olefin production device, comprising a first reaction system, a quenching system, a compression system, a separation system, a hydrogenation system, and a second reaction system connected in sequence;

[0010] The separation system is provided with a C4 and above heavy hydrocarbon discharge pipeline; the C4 and above heavy hydrocarbon discharge pipeline of the separation system is connected to the hydrogenation system;

[0011] The hydrogenation system includes a preheating heat exchanger / heat exchanger group, a hydrogenation reactor, and an optional cooler / cooler group arranged in sequence; the discharge pipeline of the hydrogenation system is connected to the second reaction system;

[0012] The second reaction system includes an intermediate product preheater / preheater group, an intermediate product reaction facility, and an optional cracking gas quencher / quencher group arranged in sequence; the discharge pipeline of the second reaction system is connected to the quenching system.

[0013] The second aspect of the present invention provides a method for preparing light olefins from alcohols, which is carried out in the above-mentioned device, comprising: raw material methanol enters a first reaction system, a quenching system, a compression system and a separation system in sequence; the intermediate product C4 and above heavy hydrocarbons obtained in the separation system enters a hydrogenation system for hydrogenation saturation, and then enters a second reaction system for steam cracking reaction to obtain a cracked gas containing ethylene and / or propylene; the cracked gas and the product gas obtained in the first reaction system are combined and enter the quenching system, and are subsequently separated to obtain an ethylene product and / or a propylene product.

[0014] The present invention provides a novel alcohol-to-light olefins process that reduces the number of byproducts compared to typical methanol-to-olefins processes, enabling the production of more ethylene and propylene products from the same methanol feedstock. This process offers significant economic benefits while also reducing overall plant energy consumption. Specifically, the device can be directly applied to newly constructed alcohol-to-light olefins plants or, with minimal additional investment, to expand the capacity of existing methanol-to-light olefins systems, demonstrating promising application prospects.

[0015] Other features and advantages of the present invention will be described in detail in the following detailed description. BRIEF DESCRIPTION OF THE DRAWINGS

[0016] The above and other objects, features and advantages of the present invention will become more apparent through a more detailed description of exemplary embodiments of the present invention with reference to the accompanying drawings, wherein like reference numerals generally represent like components throughout the exemplary embodiments of the present invention.

[0017] Figure 1 A schematic diagram of the process in the prior art is shown.

[0018] Figure 2 A process schematic diagram of an embodiment of the present invention is shown.

[0019] Figure 3 This is a specific flow chart of Example 1 and Example 2 of the present invention, and the separation system adopts a front-deethanization process.

[0020] Figure 4 This is a specific flow chart of Example 3 of the present invention, in which the separation system adopts a front-deethanization process.

[0021] Description of Reference Numerals

[0022] A. First Response System

[0023] A01, methanol conversion reactor and regenerator

[0024] B.Quick cooling system

[0025] B01, quench heat exchanger / heat exchanger group; B02, quench tower

[0026] C. Compression system

[0027] C01, compressor; C02, oxide water washing tower and alkali washing tower

[0028] D. Separation system

[0029] D01, deethanizer unit; D02, demethanizer unit; D03, C2 hydrogenation unit; D04, ethylene fractionator unit; D05, depropanizer unit; D06, propylene fractionator unit; D07, depentanizer unit; D10, OCC unit

[0030] E. Hydrogenation system

[0031] E01, hydrogenation reactor;

[0032] F. Second reaction system

[0033] F01, intermediate product preheater / preheater group; F02, intermediate product reaction facility; F03, cracking gas quencher / quencher group.

[0034] 1. Methanol; 2. Product gas; 3. Process gas; 4. Fractionated stream; 5. Pressurized fractionated stream; 6. Process gas from which acidic gases have been removed; 7. Deethanizer overhead material; 8. Methane hydrogen; 9. Deethanizer bottom material; 10. Ethane and ethylene; 11. Ethylene product; 12. By-product ethane; 13. Deethanizer bottom material; 14. C3 material; 15. Propylene product; 16. By-product propane; 17. Depropanizer bottom material; 20. Intermediate products (heavy hydrocarbons of C4 and above); 21. Hydrogen; 22. Intermediate products after hydrogenation; 23. Preheated intermediate products; 24. Cracked gas; 25. Cooled cracked gas; 26. Steam; 27. Organic sulfur; 28. OCC unit feed; 29. Crude propylene; 30. OCC outflow heavy hydrocarbons of C4 and above. DETAILED DESCRIPTION

[0035] The preferred embodiments of the present invention will be described in more detail below. Although the preferred embodiments of the present invention are described below, it should be understood that the present invention can be implemented in various forms and should not be limited to the embodiments set forth herein.

[0036] The present invention provides an alcohol-to-low-carbon olefin device, comprising a first reaction system, a quenching system, a compression system, a separation system, a hydrogenation system, and a second reaction system connected in sequence;

[0037] The separation system is provided with a C4 and above heavy hydrocarbon discharge pipeline; the C4 and above heavy hydrocarbon discharge pipeline of the separation system is connected to the hydrogenation system;

[0038] The hydrogenation system includes a preheating heat exchanger / heat exchanger group, a hydrogenation reactor, and an optional cooler / cooler group arranged in sequence; the discharge pipeline of the hydrogenation system is connected to the second reaction system;

[0039] The second reaction system includes an intermediate product preheater / preheater group, an intermediate product reaction facility, and an optional cracking gas quencher / quencher group arranged in sequence; the discharge pipeline of the second reaction system is connected to the quenching system.

[0040] According to the present invention, preferably, the first reaction system includes a methanol conversion reactor and a regenerator.

[0041] According to the present invention, preferably, the quenching system includes a quenching heat exchanger / heat exchanger group and a quenching tower which are arranged in sequence.

[0042] According to the present invention, preferably, the compression system comprises a compressor, an oxide water washing tower and an alkali washing tower which are arranged in sequence.

[0043] According to the present invention, preferably, the separation system includes a deethanizer unit, a demethanizer unit, a C2 hydrogenation unit, an ethylene fractionator unit, a depropanizer unit, a propylene fractionator unit, and optionally a debutanizer unit and / or a depentanizer unit. The debutanizer unit and the depentanizer unit can be selectively provided based on product requirements. For example, when a C4 component is required, a debutanizer unit is required; when a C5 component is required, a depentanizer unit is required.

[0044] In the separation unit, there can be many sources of C4 and above heavy hydrocarbons. Therefore, the setting of the C4 and above heavy hydrocarbon discharge pipeline can also be very flexible. As long as the hydrocarbon flow discharge pipeline (or its branch pipeline) with a carbon number of 4 or more obtained by the separation system can be used as the C4 and above heavy hydrocarbon discharge pipeline. Specifically, the C4 and above heavy hydrocarbon discharge pipeline of the separation system can be connected to at least one of the following according to the requirements of the device to produce by-products: the bottom discharge pipeline of the depropanizer unit, the top discharge pipeline of the optional debutanizer unit (if provided), the bottom discharge pipeline of the optional debutanizer unit (if provided), the top discharge pipeline of the optional depentanizer unit (if provided), and the bottom discharge pipeline of the optional depentanizer unit (if provided).

[0045] Specifically, when the separation system is provided with only a depropanizer unit, the C4 and above heavy hydrocarbons discharge line of the separation system directly serves as the bottom discharge line of the depropanizer unit, or is connected to a branch of the bottom discharge line of the depropanizer unit. In this case, the C4 and above heavy hydrocarbons produced from the depropanizer kettle include C4, C5, and C6+ components.

[0046] When the separation system is provided with a depropanizer unit and a debutanizer unit (or a depentanizer unit), at least one branch of the bottom discharge line of the depropanizer unit is connected to the debutanizer unit (or the depentanizer unit), that is, the bottom discharge line of the depropanizer unit can be connected only to the debutanizer unit (or the depentanizer unit), or it can be divided into at least two branches, one of which is connected to the debutanizer unit (or the depentanizer unit), and the other branch is connected to the C4 and above heavy hydrocarbon discharge line; the top discharge line (or a branch of the top discharge line) and / or the bottom discharge line (or a branch of the bottom discharge line) of the debutanizer unit (or the depentanizer unit) are connected to the C4 and above heavy hydrocarbon discharge line. At this time, the C4 and above heavy hydrocarbons extracted from the depropanizer bottom include C4, C5 and C6+ components; the debutanizer bottom material includes C5 and C6+ components, and the top material is C4 component; the depentanizer bottom material includes C6+ components, and the top material is C4 and C5 components.

[0047] When the separation system is provided with a depropanizer unit, a debutanizer unit and a depentanizer unit, at least one branch of the bottom discharge line of the depropanizer unit is connected to the debutanizer unit, that is, the bottom discharge line of the depropanizer unit can be connected only to the debutanizer unit, or it can be divided into at least two branches, one of which is connected to the debutanizer unit, and the other is connected to the C4 and above heavy hydrocarbon discharge line; at least one of the bottom discharge lines of the debutanizer unit is connected to the depentanizer unit (or can be connected only to the depentanizer unit), the top discharge line (or a branch of the top discharge line) of the debutanizer unit, another optional bottom discharge line, and at least one of the top discharge line (or a branch of the top discharge line) and the bottom discharge line (or a branch of the bottom discharge line) of the depentanizer unit can be connected to the C4 and above heavy hydrocarbon discharge line. At this point, the C4 and higher heavy hydrocarbons produced from the depropanizer bottoms include C4, C5, and C6+ components; the debutanizer bottoms stream includes C5 and C6+ components, with the overhead stream consisting of C4 components; the depentanizer overhead stream consists of C5 components, with the bottoms stream consisting of C6 and higher components. The C4 and higher heavy hydrocarbons discharge line can be connected to the depropanizer bottom discharge line, the debutanizer top / bottom discharge lines, the depentanizer top / bottom discharge lines, or any of their branches.

[0048] The above is only an exemplary list of the configuration of the C4 and above heavy hydrocarbon discharge pipelines. Those skilled in the art can combine and change the above methods according to the design concept of the present invention.

[0049] According to the present invention, the separation system can adopt a sequential process, a front deethanization process or a front depropanization process according to the different arrangement orders of the three characteristic distillation tower units (demethanizer unit, deethanizer unit and depropanizer unit).

[0050] According to a specific embodiment of the present invention, the separation system adopts a sequential process: the compression system is connected to the demethanizer unit, the top of the demethanizer unit is connected to a methane hydrogen discharge pipeline, and the bottom discharge pipeline is connected to the deethanizer unit; the top discharge pipeline of the deethanizer unit is connected to the carbon two hydrogenation unit and the ethylene distillation unit in sequence, and the bottom discharge pipeline is connected to the depropanizer unit; the top of the ethylene distillation unit is connected to an ethylene discharge pipeline, and the bottom is connected to a by-product ethane discharge pipeline; the top discharge pipeline of the depropanizer unit is connected to a propylene distillation unit, and the bottom discharge pipeline is optionally connected to a debutanizer unit and / or a depentanizer unit; the top of the propylene distillation unit is connected to a propylene discharge pipeline, and the bottom is connected to a by-product propane discharge pipeline.

[0051] According to another specific embodiment of the present invention, the separation system adopts a front deethanization process: the compression system is connected to the deethanizer unit, the top discharge pipeline of the deethanizer unit is connected to the demethanizer unit, and the bottom discharge pipeline is connected to the depropanizer unit; the top of the demethanizer unit is connected to a methane hydrogen discharge pipeline, and the bottom discharge pipeline is sequentially connected to a carbon two hydrogenation unit and an ethylene distillation tower unit; the top of the ethylene distillation tower unit is connected to an ethylene discharge pipeline, and the bottom is connected to a by-product ethane discharge pipeline; the top discharge pipeline of the depropanizer unit is connected to a propylene distillation tower unit, and the bottom discharge pipeline is optionally connected to a debutanizer unit and / or a depentanizer unit; the top of the propylene distillation tower unit is connected to a propylene discharge pipeline, and the bottom is connected to a by-product propane discharge pipeline.

[0052] According to another specific embodiment of the present invention, the separation system adopts a front depropanization process: the compression system is connected to the depropanizer unit, the top discharge pipeline of the depropanizer unit is connected to the demethanizer unit, and the bottom discharge pipeline is optionally connected to the debutanizer unit and / or the depentanizer unit; the top of the demethanizer unit is connected to a methane hydrogen discharge pipeline, and the bottom discharge pipeline is connected to the deethanizer unit; the top discharge pipeline of the deethanizer unit is connected to a carbon two hydrogenation unit and an ethylene distillation unit in sequence, and the bottom discharge pipeline is connected to a propylene distillation unit; the top of the ethylene distillation unit is connected to an ethylene discharge pipeline, and the bottom is connected to a by-product ethane discharge pipeline; the top of the propylene distillation unit is connected to a propylene discharge pipeline, and the bottom is connected to a by-product propane discharge pipeline;

[0053] In the above three processes, the C4 and above heavy hydrocarbon discharge line of the separation system can be connected to at least one of the following according to the requirements of the device for producing by-products: the bottom discharge line of the depropanizer unit, the top discharge line of the optional debutanizer unit (if set), the bottom discharge line of the optional debutanizer unit (if set), the top discharge line of the optional depentanizer unit (if set), and the bottom discharge line of the optional depentanizer unit (if set).

[0054] The C4 and above heavy hydrocarbon discharge pipeline of the present invention can be flexibly set according to the requirements of the device output by-products, but from the perspective of making full use of the by-products as much as possible, it is preferred to set only the depropanizer unit, and use the bottom discharge pipeline of the depropanizer unit directly as the C4 and above heavy hydrocarbon discharge pipeline. If the existing device is already provided with a debutanizer unit and / or a depentanizer unit, it is no longer fed thereto, but the bottom discharge pipeline of the depropanizer unit is directly used as the C4 and above heavy hydrocarbon discharge pipeline, that is, all the depropanizer bottom materials are recycled. In the case of an OCC unit, it is preferred to set a depentanizer unit, and the bottom discharge pipeline of the depropanizer unit is only connected to the depentanizer unit, and the top discharge pipeline of the depentanizer unit is connected to the OCC unit, and the heavy hydrocarbon discharge pipeline of the OCC unit is merged with the bottom discharge pipeline of the depentanizer unit as the C4 and above heavy hydrocarbon discharge pipeline. That is, the depentanizer bottom materials and the heavy hydrocarbons of the OCC unit are fully recycled.

[0055] According to the present invention, preferably, the intermediate product reaction facility of the second reaction system is selected from at least one of a tubular cracking reactor, a cracking furnace and a catalytic fluidized bed reactor.

[0056] According to one embodiment, an OCC unit is provided after the aforementioned MTO device. In this mode, the separation system includes a deethanizer unit, a demethanizer unit, a C2 hydrogenation unit, an ethylene distillation tower unit, a depropanizer unit, a propylene distillation tower unit, and a debutanizer unit and / or a depentanizer unit; the top discharge line of the debutanizer unit and / or the depentanizer unit is connected to the OCC unit, and the heavy hydrocarbon (usually C4 or higher hydrocarbon) discharge line of the OCC unit is merged with the bottom discharge line of the debutanizer unit and / or the depentanizer unit as the C4 and higher heavy hydrocarbon discharge line.

[0057] The present invention also provides a method for preparing light olefins from alcohols, which is carried out in the above-mentioned device, comprising: raw material methanol sequentially entering a first reaction system, a quenching system, a compression system and a separation system; intermediate products of carbon four and above heavy hydrocarbons obtained in the separation system enter a hydrogenation system for hydrogenation saturation, and then enter a second reaction system for steam cracking reaction to obtain cracked gas containing ethylene and / or propylene; the cracked gas and the product gas obtained in the first reaction system are combined and enter the quenching system, and are subsequently separated to obtain ethylene products and / or propylene products.

[0058] According to the present invention, preferably, the reaction temperature of the cracking in the intermediate product reaction facility of the second reaction system is 600-950°C, preferably 750-900°C; the reaction pressure is 0.01-0.5 MPaG, preferably 0.02-0.35 MPaG.

[0059] According to the present invention, the hydrogenated intermediate product is preheated before entering the intermediate product reaction facility of the second reaction system; the temperature after preheating can be 30-700°C, preferably 500-680°C.

[0060] According to a preferred embodiment of the present invention, the hydrogenated intermediate product is preheated using the heat of the product gas after methanol conversion in the first reaction system, and there is thermal coupling between the second reaction system and the first reaction system.

[0061] According to another preferred embodiment of the present invention, the hydrogenated intermediate product is preheated using the heat of the regenerator flue gas in the first reaction system, and there is thermal coupling between the second reaction system and the first reaction system.

[0062] According to the method of the present invention, preferably, the intermediate product reaction facility of the second reaction system is a tubular cracking reactor, which is arranged in the auxiliary combustion chamber of the methane conversion reactor of the first reaction system, and the heat required for the cracking reaction is provided by the fuel, and there is thermal coupling between the second reaction system and the first reaction system.

[0063] According to the present invention, in order to reduce the hydrocarbon partial pressure and increase the ethylene / propylene yield, the hydrogenated intermediate product is mixed with steam in a certain proportion and then enters the intermediate product reaction facility of the first reaction system; the mass ratio of the steam to the hydrogenated intermediate product is 0.1 to 1.0:1, preferably 0.25 to 0.4:1.

[0064] According to the present invention, in order to suppress coking of the reaction facility, organic sulfur is preferably added to the hydrogenated intermediate product; the amount of organic sulfur added is generally such that the organic sulfur content in the system is 50 to 1000 ppmw, preferably 150 to 300 ppmw.

[0065] According to the method of the present invention, the outlet logistics of the intermediate product reaction facility of the second reaction system can be directly incorporated into the product gas obtained by methanol conversion in the first reaction system, or it can be cooled before being incorporated; if cooling is required, it can be cooled through a one-stage or two-stage rapid cooling heat exchanger, and the temperature after cooling is generally 200-700°C, preferably 300-650°C.

[0066] When the apparatus includes an OCC unit, the method of the present invention further comprises: the top discharge of the debutanizer unit and / or the depentanizer unit of the separation system enters the OCC unit for reaction, and the heavy hydrocarbons in the resulting product are combined with the bottom discharge of the debutanizer unit and / or the depentanizer unit and then enter the hydrogenation system. Specifically, the heavy hydrocarbon discharge pipeline can be directly connected to the hydrogenation system, or it can be combined with the bottom discharge pipeline of the debutanizer unit and / or the depentanizer unit and then connected to the hydrogenation system.

[0067] According to a specific embodiment of the present invention, Figure 2 As shown, the alcohol-to-low-carbon olefins apparatus comprises: a first reaction system A, a quenching system B, a compression system C, a separation system D, a hydrogenation system E, and a second reaction system F. The raw methanol is superheated and vaporized before entering the methanol conversion reactor of the first reaction system A. The product gas obtained after the reaction enters the quenching system B. In the quenching zone, the product gas is cooled at various stages before entering the compression system C from the top of the separation tower. In the compression zone, the product gas undergoes multi-stage compression and pressure boosting by a compressor, and acidic gases and organic oxides are removed before entering the separation system D. The separation zone consists of multiple groups of distillation towers, where the components are accurately separated to ultimately produce the products ethylene and propylene, as well as various by-products. Heavy hydrocarbons with C4 or higher, as intermediate products, first pass through the hydrogenation system E to convert unsaturated hydrocarbons into saturated hydrocarbons, and then are heated to a certain temperature before being fed into the second reaction system F. In the second reaction system F, the heavy hydrocarbons with C4 or higher are cracked at an appropriate temperature to produce a cracked gas containing ethylene and propylene. The cracked gas is combined with the product gas at the outlet of the first reaction system A.

[0068] Typical methanol to olefins process flow chart Figure 1 As shown, compared to a typical methanol-to-olefins process, the process of the present invention produces C4 and higher heavy hydrocarbons as intermediate products rather than byproducts, significantly increasing the yields of ethylene and propylene and improving the ethylene:propylene ratio. As used herein, the term "intermediate products" refers to C4 and higher heavy hydrocarbons, including the depropanizer bottom discharge, as well as the optional debutanizer top and bottom discharges, the depentanizer top and bottom discharges, or portions thereof.

[0069] The logistics entering the compressor in the existing typical process is defined as (P-xianyou).

[0070] In one embodiment of the present invention, a hydrogenation system and a second reaction system are set up on the basis of a typical process and a pre-deethanization process is adopted. The cracked gas at its outlet is combined with the product gas obtained by methanol conversion in the first reaction system. The logistics entering the compressor of the present invention is defined as (P-faming).

[0071] Table 1

[0072] P-faming / P-xianyou CO 1.15～1.30 <![CDATA[CO2]]> 0.91～0.96 hydrogen 2.30～3.00 methane 1.60～2.05 Acetylene 225～285 Ethylene 1.00～1.03 Propylene 0.90～0.93 <![CDATA[C3H4]]> 150～190

[0073] Note: The above table is a comparison of the concentration multiples of the same component in P-faming and P-xianyou.

[0074] According to the changes in the above components, the equipment or local processes of the separation system of the present invention are preferably further improved.

[0075] Relative to Figure 1In the typical existing process shown, when the front deethanization process is adopted, the flux of the distillation section of the demethanizer unit of the present invention is preferably increased; specifically, the flux of the distillation section of the demethanizer unit needs to be increased to 1.5 to 2.0 times.

[0076] According to the present invention, since the acetylene composition varies greatly, the carbon dihydrogenation reactor in the carbon dihydrogenation unit of the separation system is preferably operated continuously (applicable to the three processes).

[0077] In addition to the above process conditions, the operating conditions of other equipment in the present invention can adopt conventional process conditions in the art.

[0078] According to the present invention, the byproducts ethane and propane from the ethylene and propylene fractionation units can also be utilized as intermediate products. In this case, the bottom discharge line of the ethylene and / or propylene fractionation units is directly connected to the second reaction system. The ethane and propane separated by the ethylene and propylene fractionation units are directly preheated and cracked to increase ethylene and propylene production.

[0079] The present invention is further described below by way of examples.

[0080] Example 1

[0081] This embodiment is used to illustrate the device for preparing light olefins from alcohols of the present invention. Figure 3 As shown, it includes a first reaction system A, a quenching system B, a compression system C, a separation system D, a hydrogenation system E and a second reaction system F connected in sequence;

[0082] The first reaction system A includes a methanol conversion reactor and a regenerator A01 (fluidized bed reactor);

[0083] The quenching system B includes a quenching heat exchanger / heat exchanger group B01 and a quenching tower B02 arranged in sequence;

[0084] The compression system C includes a compressor C01, an oxide water washing tower and an alkali washing tower C02 arranged in sequence;

[0085] The separation system D includes a deethanizer unit D01, a demethanizer unit D02, a carbon two hydrogenation unit D03, an ethylene distillation unit D04, a depropanizer unit D05, and a propylene distillation unit D06; the discharge pipelines of the oxide water washing tower and the alkali washing tower C02 are connected to the deethanizer unit D01, the top discharge pipeline of the deethanizer unit D01 is connected to the demethanizer unit D02, and the bottom discharge pipeline is connected to the depropanizer unit D05; the top of the demethanizer unit D02 is connected to the methane The hydrogen discharge pipeline and the bottom discharge pipeline are connected to the C2 hydrogenation unit D03 and the ethylene distillation tower unit D04 in sequence; the top of the ethylene distillation tower unit D04 is connected to the ethylene product discharge pipeline, and the bottom is connected to the by-product ethane discharge pipeline; the top discharge pipeline of the depropanizer tower unit D05 is connected to the propylene distillation tower unit D06, and the bottom discharge pipeline is connected to the C4 and above heavy hydrocarbon discharge pipeline; the top of the propylene distillation tower unit D06 is connected to the propylene product discharge pipeline, and the bottom is connected to the by-product propane discharge pipeline;

[0086] The C4 and above heavy hydrocarbon discharge pipeline of the separation system D is connected to the hydrogenation system E, which includes a preheating heat exchanger / heat exchanger group (not shown) and a hydrogenation reactor E01 arranged in sequence; unsaturated hydrocarbons are hydrogenated into saturated hydrocarbons in the hydrogenation reactor E01; the discharge pipeline of the hydrogenation system E is connected to the second reaction system F;

[0087] The second reaction system F includes an intermediate product preheater / preheater group F01, an intermediate product reaction facility F02, and a cracking gas quencher / quencher group F03 arranged in sequence; the discharge pipeline of the second reaction system F is connected to the quench heat exchanger / heat exchanger group B01 of the quenching system B.

[0088] Example 2

[0089] This embodiment is used to illustrate the method for preparing low-carbon olefins from alcohols of the present invention. Figure 3 The picture shows a 1.8 million tons methanol / year unit.

[0090] 225t / h methanol 1 is converted into product gas 2 in the methanol conversion reactor and regenerator A01. The product gas 2 is cooled to 200-300℃ by the quench heat exchanger / heat exchanger group B01. The obtained process gas 3 enters the quench tower B02 for further cooling to 40-50℃. The fractionated distillate stream 4 obtained by fractionation enters the compressor C01. In the compressor system, it is first compressed to 1.0-1.8MPaG through three stages. A cooler and an interstage tank are provided between each stage. Each stage is cooled to 35-42℃. The condensed liquid phase of each stage is separated to obtain the pressurized fractionated stream 5 After passing through the oxide water scrubber and alkali scrubber C02 (the operating pressure of the oxide water scrubber is 1.0-1.8 MPaG and the operating temperature is 35-42°C; the operating pressure of the alkali scrubber is 1.0-1.8 MPaG and the operating temperature is 42-47°C), oxygen-containing compounds and acid gases are removed and then enter the fourth stage of the compressor for compression. After being compressed to 2.5-3.7 MPaG, the material is cooled to 12-19°C in the fourth stage discharge tank of the process gas compressor to obtain gas phase and liquid phase logistics respectively. The gas and liquid phases are dried in the dryer respectively and then enter the deethanizer unit D01.

[0091] The operating pressure of the deethanizer in the deethanizer unit D01 is 2.4-3.5 MPaG, the top temperature is -22°C--15°C, and the kettle temperature is 91-100°C. The deethanizer top material 7 (composition: H2, C1, C2) is partially condensed, the liquid phase is used as the deethanizer reflux, and the gas phase enters the demethanizer unit D02. The deethanizer kettle material 13 (composition: C 3+ ) enters the depropanizer unit D05. The demethanizer in the demethanizer unit D02 operates at a pressure of 2.4 to 3.5 MPaG, a top temperature of -39°C to -22°C, and a kettle temperature of -19°C to -11°C. The demethanizer overhead material (composition: H2, C1, and a small amount of C2) is partially condensed, with the liquid phase serving as demethanizer reflux. The gaseous material (methane hydrogen 8) is discharged as by-product fuel gas. The demethanizer kettle material 9 (composition: C2) passes through the C2 hydrogenation unit D03 to remove acetylene and is dried to remove trace water before entering the ethylene distillation unit D04. The operating pressure of the ethylene distillation tower in the ethylene distillation tower unit D04 is 1.3-2.0 MPaG, the top temperature is -39°C--29°C, and the kettle temperature is -17°C--9°C. The liquid phase condensed at the top of the tower is used as reflux for the ethylene distillation tower, and the non-condensable gas in the gas phase (if any) is returned to the compressor stage. The ethylene product 11 is obtained by side-drawing, and the by-product ethane 12 is obtained in the bottom of the tower. The operating pressure of the depropanizer in the depropanizer unit D05 is 0.6-0.9 MPaG, the top temperature is 9-18°C, and the kettle temperature is 76-88°C. The top material of the depropanizer tower (composition: C3) is fully condensed, a portion of which is refluxed, and a portion is used as C3 material 14 to enter the propylene distillation tower unit D06. The depropanizer tower bottom material 17 (composition: C 4+) as an intermediate product (heavy hydrocarbons of C4 and above) 20 and sent to hydrogenation system E01. The propylene distillation tower in propylene distillation tower unit D06 utilizes a dual-tower system. The top operating pressure of the first propylene distillation tower is 1.6-2.2 MPa, the top temperature is 45-50°C, and the kettle temperature is 55-60°C. The top operating pressure of the second propylene distillation tower is 1.5-2.1 MPa, the top temperature is 43-49°C, and the kettle temperature is 45-50°C. The top of the second propylene distillation tower is fully condensed, with a portion refluxed and a portion withdrawn as propylene product 15. The by-product propane 16 is obtained in the kettle of the first propylene distillation tower.

[0092] In order to increase the production of ethylene and propylene by utilizing heavy hydrocarbons of C4 and above, the intermediate product (heavy hydrocarbons of C4 and above) 20 is preheated to 100-350°C and enters the hydrogenation reactor E01 for hydrogenation reaction to convert unsaturated hydrocarbons into saturated hydrocarbons. After exiting the hydrogenation reactor E01, the pressure is reduced to 0.5-1.0 MPaG and preheated to 500-680°C by the intermediate product preheater / preheater group F01. The preheated intermediate product 23 is mixed with steam 26 (0.25-0.4:1) and organic sulfur 27 (the content of organic sulfur in the system is 150-300 ppmw) in proportion and then enters the intermediate product reaction facility F02 for cracking. The cracked gas 24 at the outlet of the intermediate product reaction facility F02 is cooled to 300-650°C in the cracked gas quench cooler / quench cooler group F03 and then combined with the product gas 2 of the first reaction system A in the quench heat exchanger / heat exchanger group B01 as the cooled cracked gas 25, and enters the quench system B for subsequent separation to obtain ethylene product 11 and propylene product 15.

[0093] According to the method of this embodiment, by cracking the intermediate product C4 and above heavy hydrocarbons, compared with the existing alcohol to light olefin process of the same scale, the ethylene production increased by 4.921t / h, the propylene production increased by 1.818t / h, and the ethylene:propylene ratio increased from 1.049 to 1.126, increasing the production of ethylene and propylene by about 8.83wt%, and the annual production of ethylene and propylene increased by 53,900 tons. Calculated at 8,000 yuan / t (ethylene + propylene), the annual added benefit is 431.2 million yuan, and the cost invested in implementing the above process can be recovered within 3 to 6 months. It can be seen that the present invention has very significant economic benefits. Moreover, the energy consumption of dienes (ethylene and propylene) decreased by about 1%, which shows that the present invention can also reduce the energy consumption of the entire device. In addition, the present invention can be directly applied to newly built alcohol to light olefin devices, and can also be applied to the expansion and transformation of existing methanol system devices with a small amount of additional investment, and has good application prospects in the expansion and transformation of existing devices to remove bottlenecks.

[0094] Example 3

[0095] This embodiment is used to illustrate the method for preparing low-carbon olefins from alcohols of the present invention. Figure 4The 1.8 million tons of methanol / year plant shown in FIG. The difference from Example 1 is that the bottom discharge line of the depropanizer unit D05 is connected to the depentanizer unit D07, the top discharge line of the depentanizer unit D07 is connected to the OCC unit D10, the top discharge line of the OCC unit D10 is connected to the front of the C02 alkaline washing tower, and the bottom discharge line of the OCC unit D10 is combined with the bottom discharge line of the depentanizer unit D07 and then connected to the hydrogenation reactor E01.

[0096] 225t / h methanol 1 is converted into product gas 2 in the methanol conversion reactor and regenerator A01. The product gas 2 is cooled to 200-300℃ by the quench heat exchanger / heat exchanger group B01. The obtained process gas 3 enters the quench tower B02 for further cooling to 40-50℃. The fractionated distillate stream 4 obtained by fractionation enters the compressor C01. In the compressor system, it is first compressed to 1.0-1.8MPaG through three stages. A cooler and an interstage tank are provided between each stage. Each stage is cooled to 35-42℃. The condensed liquid phase of each stage is separated to obtain the pressurized fractionated stream 5 After passing through the oxide water scrubber and alkali scrubber C02 (the operating pressure of the oxide water scrubber is 1.0-1.8 MPaG and the operating temperature is 35-42°C; the operating pressure of the alkali scrubber is 1.0-1.8 MPaG and the operating temperature is 42-47°C), oxygen-containing compounds and acid gases are removed and then enter the fourth stage of the compressor for compression. After being compressed to 2.5-3.7 MPaG, the material is cooled to 12-19°C in the fourth stage discharge tank of the process gas compressor to obtain gas phase and liquid phase logistics respectively. The gas and liquid phases are dried in the dryer respectively and then enter the deethanizer unit D01.

[0097] The operating pressure of the deethanizer in the deethanizer unit D01 is 2.4-3.5 MPaG, the top temperature is -22°C--15°C, and the kettle temperature is 91-100°C. The deethanizer top material 7 (composition: H2, C1, C2) is partially condensed, the liquid phase is used as the deethanizer reflux, and the gas phase enters the demethanizer unit D02. The deethanizer kettle material 13 (composition: C 3+) enters the depropanizer unit D05. The demethanizer in the demethanizer unit D02 operates at a pressure of 2.4 to 3.5 MPaG, a top temperature of -39°C to -22°C, and a kettle temperature of -19°C to -11°C. The demethanizer overhead material (composition: H2, C1, and a small amount of C2) is partially condensed, with the liquid phase serving as demethanizer reflux. The gaseous material (methane hydrogen 8) is discharged as by-product fuel gas. The demethanizer kettle material 9 (composition: C2) passes through the C2 hydrogenation unit D03 to remove acetylene and is dried to remove trace water before entering the ethylene distillation unit D04. The operating pressure of the ethylene distillation tower in the ethylene distillation tower unit D04 is 1.3~2.0MPaG, the top temperature is -39℃~-29℃, and the kettle temperature is -17℃~-9℃. The liquid phase condensed at the top of the tower is used as the reflux of the ethylene distillation tower, and the gas phase non-condensable gas (if any) is returned to the compressor section. The side line is taken to obtain ethylene product 11, and the by-product ethane 12 is obtained in the bottom of the tower. The operating pressure of the depropanizer in the depropanizer unit D05 is 0.6-0.9 MPaG, the top temperature is 9-18°C, and the kettle temperature is 76-88°C. The top material of the depropanizer (composition: C3) is fully condensed, a portion is refluxed, and a portion is used as C3 material 14 to enter the propylene distillation unit D06. The propylene distillation tower in the propylene distillation unit D06 adopts a double tower. The top operating pressure of the 1# propylene distillation tower is 1.6-2.2 MPa, the top temperature is 45-50°C, and the kettle temperature is 55-60°C. The top operating pressure of the 2# propylene distillation tower is 1.5-2.1 MPa, the top temperature is 43-49°C, and the kettle temperature is 45-50°C. The top of the 2# propylene distillation tower is fully condensed, a portion is refluxed, and a portion is withdrawn as propylene product 15. The by-product propane 16 is obtained in the kettle of the 1# propylene distillation tower. 4+ ) enters the depentanizer unit D07. The operating pressure of the depentanizer unit D07 is 0.2-0.5 MPaG, the top temperature is 40-60°C, and the kettle temperature is 130-170°C. The depentanizer overhead material (composition: mixed C4C5) is fully condensed, part of which is refluxed and part of which is used as the OCC unit feed 28 to enter the OCC unit D10. The depentanizer kettle material 20 (composition: C 6+ ) is sent to the hydrogenation system E01 as an intermediate product. In the OCC unit, after the OCC unit feed 28 is reacted, the crude propylene 29 is returned to the CO2 alkaline scrubber, and the heavy hydrocarbons above C4 30 flowing out of the OCC merge with the depentanizer bottom material 20.

[0098] In order to increase the production of ethylene and propylene by utilizing heavy hydrocarbons of C4 and above, the depentanizer bottom material 20 is mixed with heavy hydrocarbons of C4 and above 30 flowing out of the OCC and then preheated to 100-350°C as an intermediate product (heavy hydrocarbons of C4 and above) and enters the hydrogenation reactor E01 for hydrogenation reaction to convert unsaturated hydrocarbons into saturated hydrocarbons. After exiting the hydrogenation reactor E01, the pressure is reduced to 0.5-1.0 MPaG and preheated to 500-680°C in the intermediate product preheater / preheater group F01. The preheated intermediate product 23 is mixed with steam 26 (0.25-0.4:1) and organic sulfur 27 (the content of organic sulfur in the system is 150-300 ppmw) in proportion and then enters the intermediate product reaction facility F02 for cracking. The cracked gas 24 at the outlet of the intermediate product reaction facility F02 is cooled to 300-650°C in the cracked gas quench cooler / quench cooler group F03 and then combined with the product gas 2 of the first reaction system A in the quench heat exchanger / heat exchanger group B01 as the cooled cracked gas 25, and enters the quench system B for subsequent separation to obtain ethylene product 11 and propylene product 15.

[0099] According to the method of the present embodiment, by the cracking of intermediate products of carbon four and above heavy hydrocarbons, compared with the existing same-scale alcohol-to-low-carbon olefin process, the ethylene output increased by 1.96t / h, the propylene output increased by 0.73t / h, and the production of ethylene and propylene increased by about 3.15wt%, and the annual production of ethylene and propylene increased by a total of 21,500 tons. According to 8,000 yuan / t (ethylene+propylene), the annual newly added benefits are 172 million yuan, and the cost invested in realizing the above-mentioned process can be recovered within one year. It can be seen that the present invention has very significant economic benefits. Moreover, the energy consumption of dienes (ethylene and propylene) decreased by about 0.5%, which shows that the present invention can also reduce the energy consumption of the entire device. In addition, the present invention can be directly applied to a newly-built alcohol-to-low-carbon olefin device, and can also be applied to the capacity expansion and transformation of an existing methanol system device with a small amount of additional investment, and has a good application prospect in the capacity expansion and transformation of the existing device to remove the bottleneck.

[0100] While various embodiments of the present invention have been described above, the above description is intended to be illustrative, not exhaustive, and not limited to the disclosed embodiments. Many modifications and variations will be apparent to those skilled in the art without departing from the scope and spirit of the described embodiments.

Claims

1. A method for preparing light olefins from alcohols, the method being carried out in an apparatus for preparing light olefins from alcohols, the apparatus comprising a first reaction system, a quenching system, a compression system, a separation system, a hydrogenation system, and a second reaction system connected in sequence; The first reaction system includes a methanol conversion reactor and a regenerator; The separation system is provided with a C4 and above heavy hydrocarbon discharge pipeline; the C4 and above heavy hydrocarbon discharge pipeline of the separation system is connected to the hydrogenation system; The hydrogenation system includes a preheating heat exchanger / heat exchanger group, a hydrogenation reactor, and an optional cooler / cooler group arranged in sequence; the discharge pipeline of the hydrogenation system is connected to the second reaction system; The second reaction system includes an intermediate product preheater / preheater group, an intermediate product reaction facility, and an optional cracking gas quencher / quencher group, which are arranged in sequence; the discharge pipeline of the second reaction system is connected to the quenching system; The intermediate product reaction facility of the second reaction system is a tubular cracking reactor, which is arranged in the auxiliary combustion chamber of the methane conversion reactor of the first reaction system. The heat required for the cracking reaction is provided by fuel, and there is thermal coupling between the second reaction system and the first reaction system; The separation system consists of a deethanizer unit, a demethanizer unit, a C2 hydrogenation unit, an ethylene distillation unit, a depropanizer unit, and a propylene distillation unit; The separation system adopts a front deethanization process: the compression system is connected to the deethanizer unit, the top discharge pipeline of the deethanizer unit is connected to the demethanizer unit, and the bottom discharge pipeline is connected to the depropanizer unit; the top of the demethanizer unit is connected to a methane hydrogen discharge pipeline, and the bottom discharge pipeline is connected to a C2 hydrogenation unit and an ethylene distillation unit in sequence; the top of the ethylene distillation unit is connected to an ethylene discharge pipeline, and the bottom is connected to a by-product ethane discharge pipeline; the top discharge pipeline of the depropanizer unit is connected to a propylene distillation unit, and the bottom discharge pipeline is connected to a C4 and above heavy hydrocarbon discharge pipeline; the top of the propylene distillation unit is connected to a propylene discharge pipeline, and the bottom is connected to a by-product propane discharge pipeline; The method comprises: The raw material methanol enters the first reaction system, the quenching system, the compression system and the separation system in sequence. The intermediate product C4 and above heavy hydrocarbons obtained in the separation system enter the hydrogenation system for hydrogenation saturation, and then enter the second reaction system for steam cracking reaction to obtain cracked gas containing ethylene and propylene. The cracked gas and the product gas obtained in the first reaction system are combined and enter the quenching system for subsequent separation to obtain ethylene product and propylene product. The flux of the distillation section of the demethanizer unit is increased to 1.5 to 2.0 times; the C2 hydrogenation reactor in the C2 hydrogenation unit is operated continuously; The hydrogenated intermediate product is preheated before entering the intermediate product reaction facility of the second reaction system; the hydrogenated intermediate product is preheated using the heat of the product gas after methanol conversion in the first reaction system and / or the heat of the regenerator flue gas in the first reaction system, and there is thermal coupling between the second reaction system and the first reaction system; the reaction temperature of the cracking in the intermediate product reaction facility of the second reaction system is 600-950°C; the reaction pressure is 0.01-0.5 MPaG; and the temperature after preheating is 30-700°C; The hydrogenated intermediate product is mixed with steam and then enters the intermediate product reaction facility of the second reaction system; the mass ratio of the steam to the hydrogenated intermediate product is 0.25 to 0.4:1; Adding organic sulfur to the hydrogenated intermediate product; the amount of organic sulfur added is such that the organic sulfur content in the system is 50 to 1000 ppmw; The outlet logistics of the intermediate product reaction facility of the second reaction system is cooled, and the temperature after cooling is 200-700°C.

2. The method according to claim 1, wherein The quenching system includes a quenching heat exchanger / heat exchanger group and a quenching tower arranged in sequence; The compression system comprises a compressor, an oxide water washing tower and an alkali washing tower which are arranged in sequence.

3. The method according to claim 1, wherein The reaction temperature of the cracking in the intermediate product reaction facility of the second reaction system is 750-900° C.; the reaction pressure is 0.02-0.35 MPaG.

4. The method according to claim 1, wherein The temperature after preheating is 500-680°C.

5. The method according to claim 1, wherein The amount of organic sulfur added is such that the content of organic sulfur in the system is 150 to 300 ppmw.

6. The method according to claim 1, wherein The temperature after cooling is 300-650°C.

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