Method for converting ethanol, method for producing c2-3 olefins, and method for regenerating catalysts

Abstract

To provide a method for converting ethanol by reducing the operating load of a refining system, a method for producing C2-3 olefins and a method for regenerating catalysts.SOLUTION: There is provided a method for continuously converting a raw material containing ethanol into a reaction gas containing an olefin having 3 or more carbon atoms using two or more fixed bed reactors provided with a zeolite catalyst, the method comprising a reaction step of supplying the raw material to at least one reactor of the fixed bed reactors and bringing the raw material into contact with the zeolite-containing catalyst to obtain the reaction gas, wherein the fixed bed reactor that supplies the raw materials is switched within a time period during which the propylene production rate in the reaction step is varied by 15% or less with respect to the initial production rate.SELECTED DRAWING: Figure 1

Description

[Technical Field]

[0001] The present invention relates to a method for the conversion of ethanol, 2-3 The present invention relates to a method for producing olefins and a method for regenerating a catalyst. [Background technology]

[0002] Hydrocarbons such as lower olefins and aromatic compounds are important raw materials in the chemical industry, and the development and improvement of production methods for propylene, in particular, which is expected to see increased demand, has been actively pursued.Among these, a commonly known method for producing propylene is to contact naphtha and olefins with a catalyst using zeolite as the active species.

[0003] In addition to olefins, the recent rise in environmental awareness has led to the production of chemical products such as lower olefins and aromatic compounds using biomass-derived alcohols as raw materials, which has attracted attention. In particular, since ethanol is a compound for which a production method from biomass raw materials has been established, there are high hopes for the early development of an efficient method for converting ethanol.

[0004] For example, Patent Documents 1 and 2 disclose methods for producing phosphorus-modified zeolite catalysts that promote the conversion reaction from ethanol to propylene. [Prior art documents] [Patent documents]

[0005] [Patent Document 1] China Published Patent Publication No. 107649173A [Patent Document 2] Patent No. 6229274 Summary of the Invention [Problem to be solved by the invention]

[0006] In Patent Documents 1 and 2, a raw material containing ethanol is converted into a reaction gas containing ethylene, propylene, and aromatic compounds using a zeolite catalyst. The reaction gas obtained by this reaction is purified to produce high-purity ethylene and propylene. Various purification systems are conceivable, including a distillation column designed solely for purifying the reaction gas, and a system in which the reaction gas is introduced into the purification system of an existing industrial facility, such as a naphtha cracker, for purification. These purification systems, particularly when implemented on an industrial scale, require continuous operation. Furthermore, even a slight change in the composition ratio can necessitate changes to the purification system conditions, the number of distillation columns, and the like, which increases the operational load of the purification system.

[0007] Therefore, the present invention provides an ethanol conversion method that reduces the operational load of the purification system, 2-3 The present invention aims to provide a method for producing olefins and a method for regenerating a catalyst. [Means for solving the problem]

[0008] As a result of intensive studies to achieve the above object, the present inventors have found that the above object can be achieved by using two or more fixed-bed reactors to obtain a reaction gas containing olefins having 3 or more carbon atoms from a raw material containing ethanol, and switching the fixed-bed reactor to which the raw material is supplied within a time period in which the fluctuation of the amount of propylene produced from the initial production amount is within a predetermined value. That is, the present invention includes the following embodiments. <1> A method for continuously converting a raw material containing ethanol into a reaction gas containing olefins having 3 or more carbon atoms using two or more fixed-bed reactors each equipped with a zeolite catalyst, comprising: a reaction step of supplying the raw material to at least one of the fixed-bed reactors, bringing the raw material into contact with the zeolite-containing catalyst, and obtaining the reaction gas; A method for converting ethanol, comprising switching the fixed-bed reactor to which the raw material is supplied within a time period in which the propylene production rate in the reaction step varies by 15% or less from the initial production rate. <2> When switching the fixed-bed reactor, the fluctuation of the total effective raw material supply mass flow rate of the raw material to the fixed-bed reactor related to the switching is within 5%. <1> The method for producing ethanol according to claim 1. <3> the reaction gas contains ethylene, When switching the fixed-bed reactor, the fluctuation in the ethylene production rate from the fixed-bed reactor due to the switching is limited to within 20%. <1> or <2> 1. A method for converting ethanol according to claim 1. <4> In switching the fixed bed reactor, after starting to supply the raw material to the fixed bed reactor to be switched to, the supply of the raw material to the fixed bed reactor in operation is stopped. <1> ~ <3> 1. The method for converting ethanol according to any one of the preceding claims. <5> Three or more fixed bed reactors are used, In the reaction step, the raw material is introduced into at least two reactors among the fixed bed reactors. <1> ~ <4> 1. The method for converting ethanol according to any one of the preceding claims. <6> a catalyst regeneration step of introducing an oxygen-containing regeneration gas into a reactor different from the fixed-bed reactor used in the reaction step, and burning coke deposited on the catalyst in the reaction step at a temperature in the reactor of 650°C or less; <1> ~ <5> 1. The method for producing ethanol according to any one of the preceding claims. <7> a hydrocarbon purging step of flowing an inert gas through the fixed-bed reactor to reduce the hydrocarbon concentration in the fixed-bed reactor to 5.0% by volume or less, prior to the catalyst regeneration step; <6> 1. A method for converting ethanol according to claim 1. <8> an oxygen purging step of supplying an inert gas to the fixed-bed reactor after the catalyst regeneration step to reduce the oxygen concentration in the fixed-bed reactor to 2.5% by volume or less; <6> or <7> 1. A method for converting ethanol according to claim 1. <9> an inert gas purging step of supplying a purge gas mainly containing hydrocarbons to the fixed bed reactor after the oxygen purging step, and reducing the concentration of the inert gas in the reactor to 10.0 vol % or less; <6> ~ <8> 1. The method for converting ethanol according to any one of the preceding claims. <10> The time T R and the time T required for the catalyst regeneration step in one fixed bed reactor. C But, T R ≧T C fulfill <6> ~ <9> 1. The method for converting ethanol according to any one of the preceding claims. <11> In the catalyst regeneration step, the coke is combusted at a reactor temperature of 450°C or higher. <6> ~ <10> 1. The method for converting ethanol according to any one of the preceding claims. <12> The oxygen contained in the regeneration gas is 20% by volume or less. <6> ~ <11> The method for changing ethanol according to any one of the above. <13> The operating rate of the fixed bed reactor for the reaction step is 50% or more. <6> ~ <12> 1. The method for converting ethanol according to any one of the preceding claims. <14> The silica / alumina molar ratio of the zeolite contained in the zeolite-containing catalyst is 20 to 3000. <1> ~ <13> 1. The method for converting ethanol according to any one of the preceding claims. <15> The zeolite-containing catalyst contains elemental silver or elemental phosphorus. <1> ~ <14> 1. The method for converting ethanol according to any one of the preceding claims. <16> The feed rate of the raw material per fixed bed reactor is 15,000 kg / hr or less. <1> ~ <15> 1. The method for converting ethanol according to any one of the preceding claims. <17> <1> ~ <16> C further comprises a purification step of purifying ethylene and propylene from the reaction gas obtained by the ethanol conversion method according to any one of the above. 2-3 A method for producing olefins. <18> The method includes a combining step of introducing hydrocarbons having a carbon number of 4 or more into a cracking furnace and combining the cracked components or a fraction derived therefrom with the reaction gas or a fraction derived therefrom to obtain a combined fraction, feeding the combined fraction to the purification step; <17> C described in 2-3 A method for producing olefins. <19> 1. A method for regenerating a catalyst in a process for continuously converting an ethanol-containing raw material into a reaction gas containing olefins having 3 or more carbon atoms using two or more fixed-bed reactors each equipped with a zeolite catalyst, comprising: A method for regenerating a catalyst, comprising a catalyst regeneration step of introducing an oxygen-containing regeneration gas into a fixed-bed reactor equipped with a catalyst to which coke has adhered, and combusting the coke at a temperature in the reactor of 650°C or less. [Effects of the Invention]

[0009] According to the present invention, there is provided a method for converting ethanol into ethanol, which reduces the operational load of the purification system. 2-3 A method for producing olefins and a method for regenerating a catalyst can be provided. [Brief explanation of the drawings]

[0010] [Figure 1] FIG. 1 shows a schematic diagram of one embodiment of an apparatus for converting ethanol. [Figure 2] FIG. 2 shows a schematic diagram of one embodiment of a fixed-bed, single-stage adiabatic reactor. [Figure 3] FIG. 3 shows the operation process of each reactor in Example 1. [Figure 4] FIG. 4 shows the run time and production rate results for Example 1. [Figure 5] FIG. 5 shows the results of the run time and production rate for Comparative Example 1. [Figure 6] FIG. 6 shows the operation process of each reactor in Example 2. [Figure 7] FIG. 7 shows the run time and production rate results for Example 2. [Figure 8] FIG. 8 shows the operation process of each reactor in Example 3. [Figure 9] FIG. 9 shows the run time and production rate results for Example 3. [Figure 10] FIG. 10 shows the operation process of each reactor in Example 4. [Figure 11] FIG. 11 shows the run time and production rate results for Example 4. [Figure 12] FIG. 12 shows the operation process of each reactor in Example 5. [Figure 13] FIG. 13 shows the run time and production rate results for Example 5. [Figure 14] FIG. 14 shows the operation process of each reactor in Example 6. [Figure 15] FIG. 15 shows the run time and production rate results for Example 6. [Figure 16] FIG. 16 shows the run time and production rate results for Example 7. [Figure 17] FIG. 17 shows the operating status of the reactor in Example 8. [Figure 18] FIG. 18 shows the run time and production rate results for Example 8. [Figure 19] FIG. 19 shows the run time and production rate results for Example 10. [Figure 20] FIG. 20 shows the operation process of each reactor in Example 11. [Figure 21] FIG. 21 shows the run time and production rate results for Example 11. [Figure 22] FIG. 22 is a schematic diagram of the purification equipment used in Example 12. DETAILED DESCRIPTION OF THE INVENTION

[0011] The present invention will be specifically described below. Note that the present invention is not limited to the following embodiments (present embodiments), and can be practiced in various modifications within the scope of the gist thereof. In this specification, a numerical range indicated using "to" indicates a range that includes the numerical values before and after "to" as the minimum and maximum values, respectively. In the numerical ranges described in stages in this specification, the upper limit or lower limit of a numerical range in one stage can be arbitrarily combined with the upper limit or lower limit of a numerical range in another stage. "Mainly comprises" means that the total mass of the components described as "mainly comprises" exceeds 50 mass% of the whole.

[0012] [How to convert ethanol] The ethanol conversion method according to this embodiment includes the steps of: A method for continuously converting a raw material containing ethanol into a reaction gas containing olefins having 3 or more carbon atoms using two or more fixed-bed reactors each equipped with a zeolite catalyst, comprising: a reaction step of supplying the raw material to at least one of the fixed-bed reactors, bringing the raw material into contact with the zeolite-containing catalyst, and obtaining the reaction gas; A method for converting ethanol, comprising switching the fixed-bed reactor to which the raw material is supplied within a time period in which the propylene production rate in the reaction step varies by 15% or less from the initial production rate. According to the present embodiment, the ethanol conversion method C reduces the operational load of the purification system. 2-3 It is possible to provide an olefin production method and a catalyst regeneration method. In other words, by performing catalyst regeneration under specific composition and conditions, catalyst performance degradation can be controlled, and by simultaneously performing the reaction step and the catalyst regeneration step using two or more fixed-bed reactors, it is possible to maintain an appropriate production capacity over the long term and suppress fluctuations in the production rate of the target compound over the course of the reaction. Furthermore, according to the above-described embodiment, ethanol can be converted into target compounds such as propylene and aromatic compounds with high yield.

[0013] In this embodiment, two or more fixed bed reactors equipped with a zeolite catalyst are used to continuously convert a raw material containing ethanol into a reaction gas containing olefins having three or more carbon atoms. The ethanol conversion method according to this embodiment includes the steps of: The method includes a reaction step of supplying a raw material to at least one of the fixed-bed reactors, bringing the raw material into contact with the zeolite-containing catalyst, and obtaining the reaction gas. Then, the fixed bed reactor to which the raw material is supplied is switched within a time period in which the propylene production rate in the reaction step varies by 15% or less from the initial production rate.

[0014] In addition, the ethanol conversion method according to this embodiment preferably includes the steps of: The process includes a catalyst regeneration step in which an oxygen-containing regeneration gas is introduced into a reactor different from the fixed-bed reactor used in the reaction step, and coke that has adhered to the catalyst in the reaction step is combusted at a temperature in the reactor of 650°C or less. That is, the reaction is carried out in a fixed-bed reactor in the reaction step, while the catalyst regeneration step involves introducing an oxygen-containing regeneration gas into a reactor other than the fixed-bed reactor to regenerate the catalyst, thereby preparing a fixed-bed reactor to be used in the reaction step, thereby making it possible to carry out the above-mentioned reaction step continuously for a long period of time. Furthermore, by burning the coke that adheres to the catalyst during the reaction process at a temperature inside the reactor of 650°C or less, deterioration of the catalyst can be suppressed, enabling continuous operation for a longer period of time.

[0015] First, the ethanol conversion device used in this embodiment will be described. FIG. 1 is a schematic diagram of an ethanol conversion apparatus. The ethanol conversion apparatus 100 includes two or more fixed-bed reactors each equipped with a zeolite catalyst. More specifically, the ethanol conversion apparatus 100 includes a reactor 1a and a reactor 1b for carrying out a reaction step. The reactors will be described in detail later. In the ethanol conversion apparatus 100, the raw material supply line 2 is connected to the reactor 1a via a valve V2a and to the reactor 1b via a valve V2b. The reaction gases obtained in reactor 1a and reactor 1b are introduced into the purification equipment through reaction gas supply line 3. Reactor 1a is connected to the purification equipment via valve V3a and reaction gas supply line 3. Reactor 1b is connected to the purification equipment via valve V3b and reaction gas supply line 3. The ethanol conversion apparatus 100 has a regeneration gas supply line 4 that supplies regeneration gas to reactor 1a and reactor 1b. The regeneration gas supply line 4 is connected to reactor 1a via valve V4a and to reactor 1b via valve V4b. The regeneration gas used in reactors 1a and 1b may be sent to a waste facility through a regeneration gas discharge line 5. Reactor 1a is connected to a waste facility via valve V5a and regeneration gas discharge line 5. Reactor 1b is connected to a waste facility via valve V5b and regeneration gas discharge line 5. Each step will be described in detail below.

[0016] <Reaction process> The ethanol conversion method according to this embodiment includes a reaction step of supplying a raw material containing ethanol to at least one of the fixed-bed reactors, bringing the raw material into contact with a zeolite-containing catalyst, and converting the raw material into a reaction gas containing olefins having three or more carbon atoms.

[0017] (Reactor) In this embodiment, the reaction step is carried out by filling a reactor having a fixed bed (hereinafter also referred to as a "fixed bed reactor") with a zeolite-containing catalyst. The fixed bed reactor may be either an adiabatic reactor or an isothermal reactor. Among these, it is preferable to use a fixed bed adiabatic reactor from the viewpoint of excellent operability. If necessary, a heating device for heating the raw materials may be provided before the reactor.

[0018] In this embodiment, the ethanol conversion method is carried out using two or more fixed-bed reactors. Among these, from the viewpoint of excellent stability of the production rate, it is preferable to use three or more reactors, and it is more preferable to use four or more reactors. Furthermore, from the viewpoint of excellent ease of operation, it is preferable to use 2N (N is a natural number) reactors.

[0019] The operating rate is the ratio of the number of reactors simultaneously used in the reaction step to the total number of reactors used in the ethanol conversion method. From the viewpoint of excellent production efficiency, the operating rate is preferably 50% or more, and more preferably 60% or more.

[0020] (Fixed-bed adiabatic reactor) For fixed-bed adiabatic reactors, the description in Adiabatic Fixed-Bed Reactors (Elsevier, 2014, Ch. 1, P. 4, L. 5-24 ISBN: 978-0-12-801306-9) can be referred to. Among fixed-bed adiabatic reactors, a fixed-bed single-stage adiabatic reactor having only one fixed catalyst bed is more preferred.

[0021] 2 is a schematic diagram of a fixed-bed, single-stage adiabatic reactor. The fixed-bed, single-stage adiabatic reactor 10a includes a reaction casing 12 with a heat insulator 121 attached to its outer periphery, a catalyst bed 13, a reactor inlet 14, and a reactor outlet 15. The reaction casing 12 is provided with the heat insulator 121 on its outer periphery, thereby preventing heat from escaping from within the reactor. In the production method according to this embodiment, the temperature within the reactor can be controlled by the heat generated and absorbed by the reaction.

[0022] The catalyst bed 13 is filled with a catalyst, which will be described later. A first sheathed thermocouple 161 is provided in the catalyst bed 13 immediately before it contacts the catalyst bed inlet 131. A second sheathed thermocouple 162 is provided in the catalyst bed 13 immediately after it passes through the catalyst bed outlet 132. These thermocouples measure the temperature of the mixed raw material immediately before it contacts the catalyst bed inlet 131 and the reaction gas immediately after it passes through the catalyst bed outlet 132. The positions of these thermocouples can be changed as necessary. The catalyst bed 13 may be of a multi-stage type, but is preferably of a single stage type as shown in FIG. 2.

[0023] In the fixed-bed single-stage adiabatic reactor 10 a , a raw material is introduced through a reactor inlet 14 and brought into contact with a catalyst bed 13 , and a reaction gas is taken out through a reactor outlet 15 .

[0024] (Zeolite-containing catalyst) The reaction step according to this embodiment uses a zeolite-containing catalyst. The zeolite-containing catalyst exhibits catalytic activity for converting ethanol and olefins into target compounds such as propylene. A common problem in conventional olefin production using ethanol as a feedstock with zeolite is the degradation of the zeolite due to steam during the reaction step and catalyst regeneration step. According to the ethanol conversion method according to this embodiment, catalyst performance degradation is controlled by performing catalyst regeneration under specific composition and conditions, and suitable production capacity is maintained over the long term by simultaneously performing the reaction step and catalyst regeneration step using two or more fixed-bed reactors. This suppresses fluctuations in the production rate of the target compound over the course of the reaction, making it easier to maintain the yield of the target compound over a long period of time.

[0025] The zeolite-containing catalyst may be a molded body containing zeolite and a binder.

[0026] In the conversion method according to this embodiment, the zeolite-containing catalyst preferably contains a zeolite having a 10-membered oxygen ring structure. Examples of zeolites include ZSM-5, ZSM-8, ZSM-11, ZSM-12, ZSM-21, ZSM-23, ZSM-35, and ZSM-38. Among these, MFI zeolites are preferred, and ZSM-5 is more preferred, from the viewpoint of excellent catalytic performance (catalytic activity and durability against coking).

[0027] There are no particular limitations on the method for synthesizing the zeolite according to this embodiment, but it can be produced by optimizing various conditions in a conventionally known method for hydrothermal synthesis of MFI type zeolite.

[0028] Generally, efficient methods for obtaining MFI zeolite by hydrothermal synthesis include hydrothermal synthesis using an appropriate organic structure-directing agent (SDA), hydrothermal synthesis using hydrothermally synthesized MFI zeolite as seed crystals, or hydrothermal synthesis using a seed slurry in the crystalline stage. Examples of organic structure-directing agents used here include ammonium salts, urea compounds, amines, and alcohols. It is known that not only organic SDA but also inorganic cations and anions contribute to the structure, and zeolite synthesis depends on the combined functions of each component. In the hydrothermal synthesis method of MFI zeolite described above, a suitable catalyst can be obtained by appropriately optimizing synthesis conditions such as the type of raw materials and additives (SDA), the amount of additives, pH, silica / alumina molar ratio, medium, and the ratio of cations and anions present in the raw material feed composition, as well as the synthesis temperature and synthesis time. Specific examples include the synthesis method using a seed slurry described in Japanese Patent No. 5426983 and the method exemplified in The Hydrothermal Synthesis of Zeolites (Chemcal Reviews, 2003, 103, 663-702).

[0029] In addition, commercially available zeolites can also be used as long as they are MFI zeolites having the above-mentioned specific physical properties and composition.

[0030] (Composition of Zeolite-Containing Catalyst) The silica / alumina (SiO2 / Al2O3) molar ratio of the zeolite contained in the zeolite-containing catalyst according to this embodiment can be appropriately selected, but from the viewpoint of excellent catalytic activity and propylene selectivity, it is preferably 20 to 2000, more preferably 200 to 1800, and even more preferably 400 to 1500. The silica / alumina molar ratio of the zeolite can be measured by a known method, for example, by completely dissolving the zeolite in an alkaline aqueous solution and analyzing the resulting solution by plasma emission spectroscopy or the like.

[0031] The mass ratio of silicon (Si) to aluminum (Al) (Si / Al mass ratio) contained in the zeolite-containing catalyst according to this embodiment can be selected as appropriate. From the viewpoint of excellent catalyst durability, the mass ratio is preferably 125 to 3000, and more preferably 500 to 2500. The Si / Al mass ratio can be measured by a known method, for example, by completely dissolving the zeolite-containing catalyst in an alkaline aqueous solution and analyzing the resulting solution by plasma emission spectroscopy or the like. By controlling the Si / Al mass ratio of the zeolite-containing catalyst, the affinity of the zeolite-containing catalyst with water can be reduced, and the reaction step can be carried out while maintaining catalytic performance even under hydrothermal conditions. However, the factors for this are not limited to those mentioned above.

[0032] (Doping elements in zeolite-containing catalysts) The zeolite-containing catalyst in this embodiment may contain phosphorus or at least one doping element selected from the group consisting of elements belonging to Group 11 of the periodic table (hereinafter, these elements are collectively referred to as "dope element"). Among the elements belonging to Group 11 of the periodic table, the zeolite-containing catalyst preferably contains silver.

[0033] Examples of elements belonging to Group 11 of the periodic table include copper, silver, and gold. By including an element belonging to Group 11 of the periodic table, dealumination of the zeolite is suppressed, thereby improving the durability of the catalyst. Among these elements belonging to Group 11 of the periodic table, silver is preferred from the viewpoint of excellent support efficiency.

[0034] The content of the doping element contained in the zeolite-containing catalyst may be 2.0 mass% or less, preferably 0.01 to 2.0 mass%, and more preferably 0.05 to 1.0 mass% from the viewpoint of further suppressing dealumination, relative to the total amount of the zeolite-containing catalyst. In this embodiment, the content of the doping element in the zeolite-containing catalyst is a value measured using an X-ray fluorescence analyzer. The content of phosphorus, copper, silver, or gold can be measured using a commercially available X-ray fluorescence analyzer under normal conditions according to the instruction manual. For example, when using a Rigaku product (trade name: "RIX3000"), the measurement conditions can be P-Kα radiation, a tube voltage of 50 kV, and a tube current of 50 mA.

[0035] In this embodiment, metal nitrate salts such as copper nitrate and silver nitrate may be used as raw materials for the elements belonging to Group 11 of the periodic table contained in the zeolite-containing catalyst. A zeolite-containing catalyst containing an element belonging to Group 11 of the periodic table can be obtained by ion-exchanging a zeolite-containing catalyst containing sodium as a counter cation with a metal nitrate salt and then sintering the catalyst. Ion-exchanging sodium, the counter cation in the zeolite, with a metal nitrate salt can be performed by immersing the zeolite or the zeolite-containing catalyst in an aqueous solution of the metal nitrate salt, followed by washing with water. The ion exchange rate can be improved by repeating the immersion and washing steps multiple times. Alternatively, a zeolite-containing catalyst containing an element belonging to Group 11 of the periodic table can be obtained by treating a proton-type or ammonium-type zeolite-containing catalyst with an aqueous solution of a metal nitrate salt.

[0036] In this embodiment, phosphoric acid and / or a phosphate (hereinafter also referred to as "phosphorus raw material") may be used as a source of phosphorus contained in the zeolite-containing catalyst. Phosphate is preferred as the phosphorus raw material, and among phosphates, compounds showing a solubility of 1 g or more per 100 g of water at 25°C are more preferred. Examples of phosphoric acid include phosphoric acid and pyrophosphoric acid, and examples of phosphates include ammonium phosphate salts such as ammonium phosphate, diammonium hydrogen phosphate, ammonium dihydrogen phosphate, and ammonium sodium hydrogen phosphate, potassium hydrogen phosphate, aluminum hydrogen phosphate, sodium phosphate, and potassium phosphate. Among these, ammonium phosphate salts having relatively high solubility in water are preferred, and at least one selected from the group consisting of ammonium phosphate, diammonium hydrogen phosphate, and ammonium dihydrogen phosphate is more preferred. These may be used alone or in combination of two or more.

[0037] Examples of the form of phosphorus include phosphorus polymers (e.g., polyphosphoric acid), phosphorus oxides (e.g., P2O5), and compounds in which phosphorus is added to aluminum in zeolite. A combination of these may also be included. When the zeolite contains aluminum, phosphorus has the effect of suppressing dealumination of the zeolite. In the method according to this embodiment, water is generated in the reactor, and the raw material ethanol may also contain water, creating a high-temperature steam atmosphere inside the reactor that causes dealumination. The inclusion of phosphorus in the zeolite-containing catalyst suppresses dealumination of the zeolite, further improving the durability of the catalyst.

[0038] (Method for molding zeolite-containing catalyst) The zeolite-containing catalyst according to this embodiment can be produced by molding a zeolite having the specific physical properties and composition described above, for example, as follows. The molding method is not particularly limited, and a common method can be used. Specific examples include a method of compressing and molding the catalyst components, an extrusion molding method, and a spray-dry molding method that is optimal for a fluidized bed reaction system. A binder can be used for molding. The binder is not particularly limited, and for example, silica, alumina, and kaolin can be used alone or in combination. Among these, it is preferable to use a silica binder from the viewpoint of excellent coking resistance. Commercially available binders can be used for these binders. The mass ratio of zeolite to binder is preferably in the range of 10 / 90 to 90 / 10, and more preferably in the range of 20 / 80 to 80 / 20.

[0039] (Pretreatment step for zeolite-containing catalyst) In the ethanol conversion method according to this embodiment, a pretreatment step may be performed on the zeolite-containing catalyst prior to contacting the catalyst with the feedstock. A preferred pretreatment step is a steaming step in which the catalyst is heated in the presence of steam at a temperature of 450°C or higher. Pretreatment tends to significantly inhibit catalyst degradation and improve selectivity. In the above method, the treatment is preferably performed at a temperature of 450°C to 900°C in a mixed gas of air or an inert gas such as nitrogen and steam (water vapor), with the atmosphere being not particularly limited, under conditions of a water vapor partial pressure of 0.01 atmosphere or higher. The heat treatment temperature is more preferably 500°C to 700°C. The pretreatment step can be performed using a fixed-bed reactor for converting ethanol. To facilitate the operation of the pretreatment step, it is preferable to provide a line for introducing steam into the fixed-bed reactor, a line for discharging the steam from the fixed-bed reactor, and valves for operating the opening and closing of these lines (not shown).

[0040] (Shape of zeolite-containing catalyst) In the ethanol conversion method according to the present embodiment, zeolite-containing catalysts having various shapes, such as cylindrical and ring-shaped, can be used, but cylindrical zeolite-containing catalysts are preferred for ease of handling. Among these, cylindrical zeolite-containing catalysts having a diameter of 1.6 mm or more and a length of 0.1 to 10.0 cm are preferred for catalyst strength.

[0041] (raw materials) In the reaction step according to this embodiment, a raw material containing ethanol is used. From the viewpoint of environmental friendliness, the ethanol is preferably derived from biomass. Note that biomass refers to organic resources other than fossil resources originating from plants and animals, and biomass-derived refers to a compound produced using biomass as a raw material.

[0042] From the viewpoint of excellent production efficiency of the target compound, the ethanol content in the raw material is preferably 30% by mass or more, more preferably 40% by mass or more, and even more preferably 50% by mass or more.

[0043] The raw material may contain ethylene in addition to ethanol. Ethylene produced by various methods can be used. For example, ethylene obtained by thermal decomposition of naphtha and / or ethane, direct or oxidative dehydrogenation of ethane, or dehydration of ethanol can be used. Among these, ethylene obtained by dehydrating biomass-derived ethanol is preferred from the viewpoint of environmental friendliness. The molar ratio of ethylene / ethanol in the raw material is preferably 0.20 to 2.50, more preferably 0.40 to 2.0, from the viewpoint of excellent ease of reaction control.

[0044] The raw material may further contain a hydrocarbon having 4 to 6 carbon atoms. Examples of hydrocarbons having 4 to 6 carbon atoms include olefins having 4 to 6 carbon atoms and saturated hydrocarbons having 4 to 6 carbon atoms. Among these, olefins having 4 to 6 carbon atoms, like ethylene and ethanol, can give target compounds such as propylene when contacted with a zeolite-containing catalyst. Examples of hydrocarbons having 4 to 6 carbon atoms include butene, pentene, hexene, butane, pentane, and hexane. In the mixed raw material, the molar ratio of olefins having 4 to 6 carbon atoms to ethylene is preferably 3.0 or less, more preferably 1.0 or less, and even more preferably 0.14 to 1.0. The term "olefin" above includes linear, branched, and cyclic olefins as well as cycloparaffins.

[0045] In the ethanol conversion method according to this embodiment, the raw material may further contain an oxygenated compound having 1 to 6 carbon atoms other than ethanol. Similar to ethylene and ethanol, the oxygenated compound having 1 to 6 carbon atoms can be brought into contact with a catalyst to give a target compound such as propylene. Examples of oxygenated compounds having 1 to 6 carbon atoms other than ethanol include methanol, propanol, dimethyl ether, and diethyl ether. In the raw material, the molar ratio of the oxygenated compound having 1 to 6 carbon atoms to ethylene is preferably 1.0 or less, and more preferably 0.5 or less.

[0046] Ethylene, ethanol, olefins having 4 to 6 carbon atoms, and oxygen-containing compounds having 1 to 6 carbon atoms other than ethanol are also collectively referred to as "effective raw materials."

[0047] In addition to the above-mentioned effective raw materials, the feedstock may contain saturated aliphatic hydrocarbons such as paraffin, olefins having 7 or more carbon atoms, and oxygen-containing compounds having 7 or more carbon atoms. These saturated aliphatic hydrocarbons, olefins having 7 or more carbon atoms, and oxygen-containing compounds having 7 or more carbon atoms can be converted into target compounds by contacting them with a catalyst, similar to ethylene and ethanol, but have lower reactivity than the above-mentioned effective raw materials.

[0048] In addition to the above-mentioned raw materials that can be converted to propylene by the reaction process, the raw material may also contain inert gases such as nitrogen. The raw material may also contain hydrogen or methane as a diluent gas, but it is preferable not to dilute with hydrogen. Hydrogen is sometimes used to suppress catalyst coking deterioration, but at the same time, hydrogenation reactions of the produced propylene and other compounds occur, which has the adverse effect of reducing the propylene purity (propylene / (propylene + propane)) [mol / mol].

[0049] The total content of ethylene and ethanol in the raw materials is preferably 30 to 100% by mass, more preferably 40 to 100% by mass, and even more preferably 50 to 100% by mass, based on the effective raw material supply mass, provided that the mass of ethanol converted into ethylene is used in calculating the ethanol mass and the effective raw material supply mass.

[0050] The total content of olefins having 4 to 6 carbon atoms in the raw material is preferably 65 mass % or less, more preferably 10 to 55 mass %, based on the effective raw material supply mass.

[0051] The raw material may contain diethyl ether, but preferably does not contain diethyl ether. The diethyl ether content of the effective raw material supply mass is preferably 10 mass% or less, more preferably 5 mass% or less, and even more preferably 1 mass% or less.

[0052] In the ethanol conversion method according to this embodiment, water can be included in the mixed feedstock. The ethylene and ethanol contained in the feedstock are produced by various production methods, and therefore contain "water generated during the production process." Here, "water generated during the production process" refers to water that is generated during the production process of ethylene and / or ethanol and has not been removed. In the ethanol conversion method according to this embodiment, water can be included in the mixed feedstock in addition to "water generated during the production process." Water has the effect of suppressing coking degradation by lowering the olefin partial pressure and improving the yield of lower olefins. On the other hand, from the viewpoint of excellent production efficiency of the target compound per feedstock flow rate, it is preferable not to include water in the mixed feedstock in addition to "water generated during the production process."

[0053] (Reaction conditions) The reaction temperature in the reaction step is preferably 400 to 600°C, more preferably 450 to 580°C, and even more preferably 480 to 550°C, from the viewpoint of achieving a high yield of the target compound.

[0054] When an adiabatic reactor is used, the reaction temperature is as follows. The inlet temperature of the catalyst bed is the temperature of the raw material immediately before the raw material fluid comes into contact with the catalyst bed packed in the adiabatic reactor. The outlet temperature of the catalyst bed is the temperature of the reaction gas immediately after it has passed through the catalyst bed. The temperatures of the raw material and reaction gas here refer to temperatures between 0d and 0.8d in a plane perpendicular to the fluid flow direction, where 0 is the center of the reactor and d is the distance from the center of the reactor to the inner wall surface of the reactor. The average inlet and outlet reaction temperature is calculated by measuring the inlet temperature and outlet temperature of the catalyst bed as shown in Figure 1 and using the formula: [inlet temperature of catalyst bed + outlet temperature of catalyst bed] / 2 (hereinafter, also referred to simply as "reaction temperature"). When an adiabatic reactor is used, the inlet temperature of the catalyst bed is preferably 480°C to 550°C, and the outlet temperature of the catalyst bed is preferably 480°C to 550°C. The temperature difference between the outlet temperature of the catalyst bed and the inlet temperature of the catalyst bed is preferably -80K to 80K, more preferably -60K to 60K.

[0055] In the reaction step, the temperature inside the reactor can be easily controlled by introducing heated raw materials into the reactor. As a method for heating the raw materials, from the viewpoint of excellent thermal efficiency, it is preferable to perform heat exchange between the reaction gas obtained from the outlet of the reactor and the raw materials introduced into the inlet of the reactor. In addition to the heat exchange with the reaction gas described above, the raw materials can also be heated in a heating furnace or the like.

[0056] The reaction pressure in the reaction step is preferably 0.01 to 4.0 MPaG, more preferably 0.05 to 3.0 MPaG. The reaction time in the reaction step is preferably 24 hours or more, more preferably 36 hours or more, and even more preferably 48 hours or more, in order to facilitate the operation of the process.

[0057] The effective raw material feed mass per reactor is preferably 0.1 to 1000 h in terms of the mass-based hourly space velocity (WHSV) of the zeolite-containing catalyst. -1 and more preferably 0.1 to 100 hours -1 , more preferably 0.5 to 50 hours-1 In the reaction step, WHSV is calculated by converting ethanol into ethylene as shown in the following formula: From the viewpoint of excellent productivity of the target compound, the effective raw material supply mass is preferably 1 kg / hr or more, more preferably 10 kg / hr or more, and even more preferably 1 t / hr or more. WHSV(hr -1 ) = Effective raw material supply mass (kg / hr) / Amount of catalyst (kg) Effective raw material supply mass (kg / hr) = ethylene flow rate (kg / hr) + ethylene equivalent ethanol flow rate (kg / hr) + carbon number 4-6 olefin flow rate (kg / hr) + carbon number 1-6 oxygenated compound flow rate other than ethanol (kg / hr) Ethylene-equivalent ethanol flow rate (kg / hr) = ethanol flow rate (kg / hr) × ethylene molecular weight (g / mol) / ethanol molecular weight (g / mol)

[0058] From the viewpoint of excellent valve operability, the amount of raw material supplied per reactor is preferably 15,000 kg / hr or less, and more preferably 10,000 kg / hr or less.

[0059] (switching) In the ethanol conversion method according to this embodiment, the fixed-bed reactor supplying the raw material is switched within a time period in which the propylene production rate in the reaction step fluctuates by 15% or less from the initial production rate. The fixed-bed reactor is switched within a time period in which the fluctuating rate is preferably 12% or less, more preferably 10% or less, and even more preferably 8% or less.

[0060] The propylene production rate is the rate in the fixed bed reactor in use. The propylene production rate is measured as the amount produced per second and calculated using the following formula: Production rate (kg / hr) = Production volume per second (kg / s) x 60 x 60 Production volume per second (kg / s) = Product concentration in reaction gas (wt%) x Production volume of reaction gas per second (kg / s)

[0061] "Initial production rate" means the rate after the fixed bed reactor is started up or restarted.

[0062] In the reaction process, the raw material is brought into contact with a zeolite-containing catalyst in a fixed-bed reactor to obtain a reaction gas. However, over time, coke adheres to the zeolite catalyst, causing a decrease in the propylene production rate. When the composition is introduced into a refining facility such as a naphtha cracker, the decrease in production rate causes the propylene concentration in the supplied composition to change over time, which requires changes to the operating conditions of the refining facility, making the facility operation more complicated. Therefore, by switching the fixed-bed reactor to which the raw material is supplied within a time period that keeps the propylene production rate within the above-mentioned change range, the fluctuation range of the propylene production rate can be kept within a predetermined range.

[0063] When switching between fixed-bed reactors, it is preferable to keep the total fluctuation of the effective raw material feed mass flow rate of the raw material to the fixed-bed reactor involved in the switching within 5%. By keeping the fluctuation of the raw material feed rate within a certain range, it is possible to keep the fluctuation of the reaction gas production rate within a certain range, making it easier to keep the production of the target compound constant. It is preferable to keep the total fluctuation of the effective raw material feed mass flow rate of the raw material to the fixed-bed reactor involved in the switching within 3%, and more preferably within 1%.

[0064] Here, "fixed-bed reactors involved in the switchover" refers to both the fixed-bed reactor to be discontinued and the fixed-bed reactor to be started up for use. "Total effective raw material feed mass flow rate of raw materials" refers to the total value of effective raw material feed mass flow rates of raw materials fed to the fixed-bed reactor to be discontinued and the fixed-bed reactor to be started up for use. "Fluctuation" refers to the fluctuation before and after the switchover, based on the feed amount to the fixed-bed reactor to be discontinued before the switchover.

[0065] From the viewpoint of excellent operability in subsequent purification, the fluctuation in the ethylene production rate during switching is preferably within 20%, more preferably within 15%, and even more preferably within 10%.

[0066] Switching between fixed-bed reactors involves stopping and starting the supply of raw materials to each reactor, which poses a problem of the production amount of the target compound being prone to instantaneous fluctuations. In particular, if the production of the target compound stops instantaneously, the balance of pressure and other factors within the process is disrupted, resulting in a significant decrease in operating efficiency. In order to prevent the production of the target compound from stopping instantaneously, it is preferable to stop the supply of raw materials to the operating fixed-bed reactor after starting the supply of raw materials to the fixed-bed reactor to be switched to.

[0067] In the ethanol conversion method according to this embodiment, a buffer tank for temporarily storing the ethanol conversion fraction or a fraction derived therefrom can be provided between the reactor and the purification equipment as a measure to prevent fluctuations in the production amount of the target compound. The size of the buffer tank is preferably equal to or greater than the volume of the reaction gas produced per 0.5 hours, more preferably equal to or greater than the volume of the reaction gas produced per hour, and even more preferably equal to or greater than the volume of the reaction gas produced per 6 hours, under the reaction temperature and pressure conditions, in order to suppress fluctuations due to switching. Furthermore, in order to suppress fluctuations in the reaction process, the volume of the buffer tank is preferably equal to or greater than 1 / 96, more preferably equal to or greater than 1 / 48, and even more preferably equal to or greater than 1 / 24 of the volume of the reaction gas produced per reaction time.

[0068] In the ethanol conversion method according to this embodiment, the number of times the reactor is switched per day is referred to as the switching frequency (times / day). From the viewpoint of excellent stability in the production rate of each compound, the switching frequency is preferably 0.25 times / day or more, and more preferably 0.50 times / day or more. Furthermore, from the viewpoint of excellent ease of process operation, the switching frequency is preferably 8.0 times / day or less, and more preferably 4.0 times / day or less.

[0069] In the ethanol conversion method according to this embodiment, after switching, the fixed-bed reactor that has been discontinued is preferably subjected to the above-mentioned catalyst regeneration step.

[0070] <Catalyst regeneration process> The ethanol conversion method according to this embodiment preferably includes a catalyst regeneration step in which an oxygen-containing regeneration gas is introduced into a reactor separate from the fixed-bed reactor used in the reaction step, and coke deposited on the catalyst during the reaction step is combusted at a reactor temperature of 650°C or lower. In the reaction in which ethanol is converted using a zeolite catalyst to produce hydrocarbons such as propylene and aromatic compounds, coking and structural degradation of the zeolite progress during the reaction. While the reaction produces equimolar water as a by-product with the feed ethanol, steam at temperatures above 400°C causes structural collapse of the zeolite catalyst due to dealumination, reducing the catalytic performance of the zeolite. Even in catalysts where structural degradation has been suppressed by techniques such as element doping, degradation of catalytic performance due to coke deposition during the reaction is unavoidable. Furthermore, because combustion of coke deposited on the catalyst involves the generation of steam and localized heat generation, structural degradation of the catalyst also progresses during catalyst regeneration. In other words, when converting ethanol by contacting a feed containing ethanol with zeolite, catalytic activity declines with repeated cycles of the reaction and regeneration steps, resulting in a low yield of the target compound. In particular, when a reactor having a fixed bed is used, catalyst replacement or additional charging is not possible during the reaction, and therefore deterioration of catalyst performance leads to a significant decrease in productivity. According to the catalyst regeneration process of this embodiment, coke adhering to the zeolite-containing catalyst can be efficiently burned. Furthermore, according to the catalyst regeneration process of this embodiment, the performance of the zeolite-containing catalyst can be regenerated while suppressing structural deterioration of the zeolite catalyst.

[0071] In one aspect, the present embodiment is a method for regenerating a catalyst. The present embodiment is a method for regenerating a catalyst in a process for continuously converting an ethanol-containing raw material into a reaction gas containing olefins having 3 or more carbon atoms using two or more fixed-bed reactors each equipped with a zeolite catalyst, the method comprising: The method includes a catalyst regeneration step of introducing an oxygen-containing regeneration gas into a fixed-bed reactor equipped with a catalyst to which coke has adhered, and combusting the coke at a temperature inside the reactor of not more than 650° C. This catalyst regeneration method makes it possible to regenerate the performance of a zeolite-containing catalyst while suppressing structural deterioration of the zeolite catalyst.

[0072] In the catalyst regeneration step according to this embodiment, the zeolite-containing catalyst is regenerated in the reactor without being removed from the reactor. By carrying out the catalyst regeneration step in the reactor, the zeolite-containing catalyst can be immediately supplied to the reaction step after the catalyst regeneration step. The catalyst regeneration step is preferably carried out while other reactors are being supplied to the reaction step.

[0073] In the catalyst regeneration step according to this embodiment, the temperature inside the reactor is preferably 650°C or lower. While increasing the temperature inside the reactor can reduce the time required for catalyst regeneration, maintaining a temperature of 650°C or lower suppresses structural deterioration of the zeolite and maintains catalytic activity after regeneration. In the catalyst regeneration step, the coke is preferably combusted at a reactor temperature of 450°C or higher to achieve a superior coke combustion rate. The reactor temperature in the catalyst regeneration step is preferably 450 to 650°C, more preferably 500 to 625°C, and even more preferably 540 to 600°C. Furthermore, since this facilitates control of heat generation associated with coke combustion, it is preferable to increase the temperature from the start to the end of the catalyst regeneration step. For example, it is preferable to perform a low-temperature treatment at 450 to 550°C, followed by a high-temperature treatment at 500 to 650°C. The terms "low-temperature treatment" and "high-temperature treatment" used herein refer to relatively low-temperature and high-temperature treatments, with the "low-temperature treatment" taking precedence.

[0074] The regeneration gas used in the catalyst regeneration step can be air or a mixture of air or oxygen and an inert gas. The oxygen concentration in the regeneration gas is preferably 20% by volume or less, more preferably 10% by volume or less, and even more preferably 0.050 to 2.0% by volume. By setting the oxygen concentration in the regeneration gas within a suitable range, the rate of the coke combustion reaction in the catalyst regeneration step can be controlled and heat generation associated with combustion can be suppressed. Furthermore, since this makes it easier to control heat generation associated with coke combustion, it is preferable to increase the oxygen concentration from the start to the end of the catalyst regeneration step. For example, it is preferable to perform a low-concentration treatment at an oxygen concentration of 0.4% by volume or less, followed by a high-concentration treatment at an oxygen concentration of 0.5% by volume or more.

[0075] The catalyst regeneration step may include a hydrocarbon purging process in which an inert gas is passed through the reactor before the regeneration gas is passed through the reactor to reduce the hydrocarbon concentration in the reactor to 5.0% by volume or less. By discharging the hydrocarbons from the reactor before the regeneration gas is passed through, heat generation associated with the reaction between the hydrocarbons remaining in the reactor and oxygen can be suppressed. The hydrocarbon concentration in the reactor after the hydrocarbon purging process is preferably 5.0% by volume or less, more preferably 3.0% by volume or less, and even more preferably 1.0% by volume or less. Examples of inert gases used in the hydrocarbon purging process include nitrogen and argon.

[0076] The catalyst regeneration step may include an oxygen purging step in which an inert gas is passed through the reactor after coke combustion to reduce the oxygen concentration in the reactor to 2.5% by volume or less. By including the oxygen purging step in the catalyst regeneration step, heat generation associated with the reaction between the oxygen remaining in the reactor and the raw materials supplied when the reaction step is restarted can be suppressed. The oxygen concentration in the reactor after the oxygen purging step is preferably 2.5% by volume or less, more preferably 2.0% by volume or less, and even more preferably 1.5% by volume or less. Examples of inert gases used in the oxygen purging step include nitrogen and argon.

[0077] The catalyst regeneration process may include an inert gas purging process after the oxygen purging process, in which a purge gas mainly containing hydrocarbons is supplied to the reactor to reduce the inert gas concentration in the reactor to 10.0% by volume or less. The purge gas may be the raw material or the reaction gas. The inert gas-containing purge gas used in the inert gas purging process is discarded without being introduced into the purification process. By reducing the inert gas concentration in the reactor after regeneration and before starting the supply of raw materials, it is possible to prevent the inert gas from being introduced into the purification system, which would otherwise reduce the operational efficiency of the purification system. The inert gas concentration in the reactor after the inert gas purging process is preferably 10.0% by volume or less, more preferably 5.0% by volume or less, and even more preferably 2.5% by volume or less.

[0078] In the catalyst regeneration process of this embodiment, the catalyst regeneration process is completed when the difference between the carbon dioxide concentration in the regeneration gas supplied to the reactor and the carbon dioxide concentration in the regeneration gas discharged from the reactor becomes 0.01% by volume or less. Hereinafter, the time required from the start of introduction of the regeneration gas into the reactor to the completion of the catalyst regeneration process will be referred to as the "time required for the catalyst regeneration process T C The time required for the catalyst regeneration process T C In the catalyst regeneration step, the time T R and the time required for the catalyst regeneration process T C But, T R ≧T C It is preferable to satisfy T R ≧T C By satisfying the above condition, the reactor can be used efficiently in the reaction step. In addition, the time required for the catalyst regeneration step in a fixed-bed reactor can be shortened by increasing the oxygen concentration of the regeneration gas and the temperature inside the reactor.

[0079] As described above, in the ethanol conversion method according to this embodiment, the reaction step is carried out continuously using two or more fixed-bed reactors. This will be explained in more detail using the schematic diagram of one embodiment of an ethanol conversion apparatus shown in FIG. 1. In the ethanol conversion method according to this embodiment, at least one fixed-bed reactor (here, fixed-bed reactor 1a) is operated to carry out the reaction step. Although the reaction performance deteriorates over time due to coke deposition on the zeolite catalyst, the fixed-bed reactor to which the raw material is supplied is switched from fixed-bed reactor 1a to fixed-bed reactor 1b within a time period in which the propylene production rate in the reaction step fluctuates by 15% or less from the initial production rate. Thereafter, fixed-bed reactor 1a is subjected to a catalyst regeneration step. Furthermore, within the time period during which the propylene production rate in the fixed-bed reactor 1b in the reaction step fluctuates by 15% or less from the initial production rate, the fixed-bed reactor to which the raw material is supplied is switched again from the fixed-bed reactor 1b to the fixed-bed reactor 1a. Thereafter, the fixed-bed reactor 1b is subjected to a catalyst regeneration step. As described above, when focusing on one fixed bed reactor, the reaction step and the catalyst regeneration step can be repeatedly carried out to continuously obtain reaction gas.

[0080] In the ethanol conversion system shown in Figure 1, all valves are closed. First, valves V2a and V3a are opened, and reaction gas is supplied to fixed-bed reactor 1a, where the reaction process is carried out. The propylene production rate in the reaction process is switched within a time period in which fluctuations of the initial production rate do not exceed 15%. After the switchover, valves V2a and V3a are closed, and valves V2b and V3b are opened. During the fixed-bed reactor switchover, the feedstock supply rates to fixed-bed reactors 1a and 1b are adjusted by valve operation, thereby keeping the total feedstock mass flow rate to all reactors within 5% before and after the switchover. To suppress fluctuations in the feedstock supply rate, for example, when stopping the feedstock supply to fixed-bed reactor 1a, the closing operation of valve V2a is initiated while valve V2b is being opened. In this switching operation, when the fully open state of the valve is represented by 100%, it is preferable to start the closing operation of the valve 2a before the opening degree of the valve V2b reaches 10%.

[0081] Next, regeneration gas is introduced into the fixed-bed reactor 1a to carry out the catalyst regeneration process. At this time, valves V4a and V5a are opened. During the catalyst regeneration process, valve V3a is closed to separate the reactor from the purification equipment. The regeneration gas used in the catalyst regeneration is introduced into the disposal equipment through regeneration gas disposal line 5.

[0082] Although not shown, it is preferable to provide a double valve in order to easily avoid contact between combustible gases such as raw materials and reaction gases and the regeneration gas and to improve operability.

[0083] Although not shown, each reactor is preferably provided with an inert gas supply line for supplying an inert gas, an inert gas waste line for introducing the inert gas into waste equipment, and a valve for operating the opening and closing of the line.

[0084] In the ethanol conversion method according to this embodiment, three or more fixed-bed reactors may be used, and the raw material may be introduced into at least two of the fixed-bed reactors in the reaction step. In this case, the propylene production rate refers to the total production rate of the two reactors operating in the reaction step. At least one fixed-bed reactor supplying the raw material may be switched within a time period in which the total propylene production rate in the reaction step fluctuates by 15% or less from the initial total production rate.

[0085] [Product: Reaction gas containing olefins with 3 or more carbon atoms] In the ethanol conversion method according to this embodiment, a reaction gas containing olefins having 3 or more carbon atoms is obtained by contacting a raw material with a zeolite-containing catalyst. The reaction gas may contain ethylene. The reaction gas may contain hydrogen, aliphatic hydrocarbons having 1 to 3 carbon atoms, aliphatic hydrocarbons having 4 to 6 carbon atoms, aromatic compounds, and hydrocarbons having 9 or more carbon atoms.

[0086] In the ethanol conversion method according to this embodiment, aliphatic hydrocarbons having 1 to 3 carbon atoms, aliphatic hydrocarbons having 4 to 6 carbon atoms, aromatic compounds, and hydrocarbons having 9 or more carbon atoms are referred to as target compounds.

[0087] The propylene content in the reaction gas is preferably 5 to 50 mass %, more preferably 8 to 40 mass %, and even more preferably 10 to 30 mass %.

[0088] The ethylene content in the reaction gas is preferably 10 to 50 mass %, more preferably 15 to 45 mass %, and even more preferably 20 to 40 mass %.

[0089] The mass ratio of propylene to ethylene (P / E ratio) in the reaction gas is preferably 0.2 to 1.2, more preferably 0.4 to 1.0, and even more preferably 0.5 to 0.8.

[0090] The content of the aromatic compound in the reaction gas is preferably 1 to 50 mass %, more preferably 2 to 30 mass %, and even more preferably 3 to 15 mass %.

[0091] In the ethanol conversion method according to this embodiment, a separation step is provided before the purification step, and various separation operations are carried out to efficiently purify the target compound. Among these, it is preferable to include a step of separating the reaction gas obtained in the above-described reaction step into a fraction A mainly containing hydrocarbons having 2 to 3 carbon atoms and a fraction B mainly containing hydrocarbons having 4 to 6 carbon atoms using a separation device. By separating into each fraction in this manner, the target compound can be efficiently separated. Examples of separation devices used in the separation step include a pressure booster, a distillation column, a quench column, and a decanter.

[0092] The term "mainly comprises" in various fractions means that the total mass of the components described as "mainly comprises" exceeds 50 mass% of each fraction. These separation steps can be carried out by combining various known methods such as distillation and extraction.

[0093] [Refining process] The ethanol conversion method according to this embodiment may include a purification step of purifying ethylene and propylene from the reaction gas. 2-3 The olefin production method includes a purification step of purifying ethylene and propylene from the reaction gas obtained by the ethanol conversion method according to this embodiment. 2-3 "Olefins" are ethylene and propylene. The purification of ethylene and propylene can be carried out, for example, by cryogenic separation.

[0094] Alternatively, the reaction gas obtained in the above-described reaction step or a fraction separated from the reaction gas may be introduced into a purification system of an ethylene plant, and the purification step may be carried out using the equipment of the ethylene plant. The ethylene plant refers to a plant that introduces hydrocarbons having 4 or more carbon atoms into a cracking furnace to produce a cracking fraction, and then produces ethylene and other olefins, including propylene, from the cracking fraction. That is, the ethanol conversion method according to this embodiment includes a combining step in which the cracking components or a fraction derived therefrom obtained by introducing hydrocarbons having 4 or more carbon atoms into a cracking furnace are combined with the reaction gas or a fraction derived therefrom to obtain a combined fraction, and the combined fraction can be introduced into a purification step to purify ethylene and propylene.

[0095] FIG. 22 is a schematic diagram of the purification process in an ethanol conversion apparatus according to this embodiment. The reaction gas produced in reactor 1 is introduced into ethylene purification column 2 after undergoing a separation process. In the ethylene purification column, a fraction containing mainly ethylene is obtained from the top, and a fraction containing mainly olefins having 3 or more carbon atoms is obtained from the bottom. The fraction containing mainly olefins having 3 or more carbon atoms is introduced into propylene purification column 3, and a fraction containing mainly propylene is obtained from the top. Furthermore, the fraction containing mainly ethylene or propylene can be further purified to improve the purity of each compound. An example of such a purification process is distillation. [Example]

[0096] The present embodiment will be described in more detail below with reference to examples, but the present embodiment is not limited to the following examples.

[0097] [Method for measuring catalyst properties] The various physical properties of the catalyst were measured as follows.

[0098] (1) Silica / alumina molar ratio of zeolite Zeolite was completely dissolved in sodium hydroxide solution to prepare a solution. The amounts of Si and Al contained in the solution were measured using an inductively coupled plasma (ICP) emission spectrometer (Rigaku, product name "JY138") according to standard methods, and the silica / alumina molar ratio was calculated from the results. The measurement conditions were: high-frequency power: 1 kW, plasma gas: 13 L / min, sheath gas: 0.15 L / min, nebulizer gas: 0.25 L / min, Si measurement wavelength: 251.60 nm, Al measurement wavelength: 396.152 nm.

[0099] (2) Al content and Si content in the zeolite-containing catalyst A solution in which the zeolite-containing catalyst was completely dissolved in a sodium hydroxide solution was prepared, and the amount of aluminum (Al) in the zeolite-containing catalyst was measured and calculated in the same manner as in (1).

[0100] (3) The content of doping elements in the zeolite-containing catalyst The content of the doping element in the zeolite-containing catalyst was measured by a conventional method using an X-ray fluorescence analyzer (manufactured by Rigaku, trade name "RIX3000").

[0101] (4) Zeolite structural type The structural type of the zeolite in the zeolite-containing catalyst was identified by measuring the X-ray diffraction pattern of the zeolite using an X-ray analyzer (manufactured by Bruker, product name "D8 Advance") and referring to the diffraction patterns of known zeolites. The measurement conditions were as follows: Cu cathode Tube voltage: 40kV Bulb current: 40mA Scan speed: 6deg / min

[0102] [How to convert ethanol] (Reactor) In the following Examples and Comparative Examples, evaluations were carried out using, as necessary, multiple fixed-bed single-stage adiabatic reactors as shown in Figure 2. The rate of reactors simultaneously used in the reaction step among all reactors used was calculated as the operating rate using the following formula. Operating rate (%) = Number of reactors simultaneously used in the reaction process / Total number of reactors

[0103] (temperature measurement) The temperatures at the catalyst bed inlet and outlet were measured using thermocouples inserted from outside the reactor. Specifically, as shown in Figure 2, in a plane perpendicular to the fluid flow direction, the center of the reactor was set to 0, and the distance from the center of the reactor to the inner wall surface of the reactor was set to d. The temperature was measured at 0.5d to 0.6d. Note that the effect of heat radiation due to the insertion of the thermocouples was negligible. If necessary, the thermocouple was moved in the fluid flow direction to measure the lowest temperature inside the reactor.

[0104] (raw materials) In the following examples, a raw material containing ethanol was used. In addition to ethanol, the raw material may contain ethylene, an olefin having four or more carbon atoms, and water. The effective raw material supply mass flow rate was calculated using the following formula. Effective raw material supply mass flow rate (kg / hr) = ethylene flow rate (kg / hr) + ethylene-equivalent ethanol flow rate (kg / hr) + olefin flow rate with carbon number 4 or more (kg / hr) Ethylene-equivalent ethanol flow rate (kg / hr) = ethanol flow rate (kg / hr) × ethylene molecular weight (g / mol) / ethanol molecular weight (g / mol)

[0105] (Reaction evaluation of reaction process) According to the following examples and comparative examples, the reaction was carried out so that the average inlet and outlet reaction temperature was 510°C. A portion of the gas at the reactor outlet was sampled every 3 hours from the start of the reaction and introduced into a gas chromatograph (TCD, FID detector) to analyze the reaction gas composition. The average value of the GC analysis results from the start of the reaction to the end of the reaction was calculated. The average inlet and outlet reaction temperature was calculated according to the following formula. Average inlet / outlet reaction temperature (℃) = [catalyst bed inlet temperature (℃) + catalyst bed outlet temperature (℃)] / 2

[0106] [Gas chromatograph analysis conditions] (Reaction gas analysis) Equipment: Shimadzu GC-2030 Column: Custom capillary column SPB-1 (inner diameter 0.25 mm, length 60 m, film thickness 3.0 μm) manufactured by SUPELCO, USA Sample gas volume: 1 mL (sampling line is kept at 200°C to 300°C) Temperature program: maintained at 40°C for 12 minutes, then heated to 200°C at 5°C / min, and maintained at 200°C for 22 minutes. Split ratio: 200:1 Carrier gas (nitrogen) flow rate: 120 mL / min FID detector: Air supply pressure 50 kPa (approx. 500 mL / min), hydrogen supply pressure 60 kPa (approx. 50 mL / min) Measurement method: A TCD detector and an FID detector were connected in series, and composition analysis was performed based on data detected by the TCD detector for hydrogen, carbon dioxide, and carbon monoxide, and data detected by the FID detector for oxygen-containing compounds such as hydrocarbons and ethanol.The concentration of each component in the reaction gas was determined using a calibration curve method, and the mass per unit time (mass flow rate) produced by the reaction was calculated.The maximum and minimum production rates for each compound were calculated based on the production rates after a steady state was reached in each example and comparative example, that is, after all reactors were used for the reaction process, catalyst regeneration process, and switching process.

[0107] <Manufacturing speed> The various production rates were measured by measuring the product production amount per second from the product concentration in the reaction gas and the reaction gas production amount per second, and calculated by the following formula. The product concentration in the reaction gas was measured and updated once every 3 hours. Production rate (kg / hr) = Production amount per second (kg / s) × 60 × 60 Production amount per second (kg / s) = Product concentration in the reaction gas (wt%) × Reaction gas production amount per second (kg / s)

[0108] <P / E ratio> The P / E ratio represents the mass ratio of propylene to ethylene in the reaction gas obtained from the reaction process and was calculated by the following formula. P / E ratio (-) = Propylene production rate (kg / hr) / Ethylene production rate (kg / hr)

[0109] <Ethylene fluctuation rate> The ethylene fluctuation rate represents the magnitude of the fluctuation of the ethylene production rate in the reaction process and was calculated by the following formula. Ethylene fluctuation rate (%) = (Maximum ethylene production rate (kg / hr) - Minimum ethylene production rate (kg / hr)) / Minimum ethylene production rate (kg / hr) × 100

[0110] <Propylene fluctuation rate> The propylene fluctuation rate represents the magnitude of the fluctuation of the propylene production rate in the reaction process and was calculated by the following formula. Propylene fluctuation rate (%) = (Maximum propylene production rate (kg / hr) - Minimum propylene production rate (kg / hr)) / Minimum propylene production rate (kg / hr) × 100

[0111] <Aromatic fluctuation rate> The aromatic fluctuation rate represents the magnitude of the fluctuation of the aromatic production rate in the reaction process and was calculated by the following formula. Aromatic fluctuation rate (%) = (Maximum aromatic production rate (kg / hr) - Minimum aromatic production rate (kg / hr)) / Minimum aromatic production rate (kg / hr) × 100

[0112] [Production Example 1: Preparation of Zeolite-Containing Catalyst 1] A clay obtained from 70 parts by mass of proton-type ZSM-5 (silica / alumina molar ratio 980), a medium pore size zeolite, and 30 parts by mass of silica (water content adjusted using colloidal silica and fumed silica) was kneaded and then extrusion-molded to obtain an extruded molded body with a diameter of 2.1 mm and a length of 4 to 6 mm. The obtained molded body was calcined at 600°C for 5 hours to obtain zeolite-containing catalyst 1.

[0113] [Production Example 2: Preparation of Zeolite-Containing Catalyst 2] Clay obtained from 70 parts by mass of proton-type ZSM-5 (silica / alumina molar ratio 980), a medium pore size zeolite, and 30 parts by mass of silica (water content adjusted using colloidal silica and fumed silica) was kneaded and then extruded to obtain an extruded molded body with a diameter of 2.1 mm and a length of 4 to 6 mm. The obtained molded body was calcined at 600°C for 5 hours to obtain a catalyst precursor. The obtained catalyst precursor was stirred in a 0.01N silver nitrate aqueous solution for 1 hour, then filtered and washed three times, and calcined at 600°C for 5 hours to obtain zeolite-containing catalyst 2. At this time, the silver content in zeolite-containing catalyst 2 was 0.16% by mass.

[0114] [Reference example 1] A feedstock containing 19.2 mass% ethylene, 48.1 mass% ethanol, 18.1 mass% water, and 14.6 mass% 1-butene was fed to a reactor packed with zeolite-containing catalyst 1 at a WHSV of 4.0, a pressure of 0.15 MPa, and a catalyst bed inlet temperature of 510°C, and the reaction was carried out for 48 hours. At the start of the reaction, the ethylene yield was 25.0 mass%, the propylene yield was 16.8 mass%, and the aromatics yield was 2.4 mass%. After 48 hours, the ethylene yield was 28.8 mass%, the propylene yield was 15.5 mass%, and the aromatics yield was 1.6 mass%.

[0115] [Reference example 2] Reference Example 2 was carried out in the same manner as Reference Example 1, except that zeolite-containing catalyst 2 was used. At the start of the reaction, the ethylene yield was 25.9 mass%, the propylene yield was 16.7 mass%, and the aromatics yield was 2.4 mass%, and after 48 hours had passed, the ethylene yield was 29.0 mass%, the propylene yield was 15.7 mass%, and the aromatics yield was 1.6 mass%.

[0116] [Example 1] In Example 1, a reaction process was carried out continuously for 144 hours by switching between two reactors consisting of reactor 1 and reactor 2 as follows. The operating status of each reactor in this example is shown in Figure 3.

[0117] 1)0~48 hours (Reaction step) The raw materials were supplied to reactor 1, and a 48-hour reaction process was initiated in the same manner as in Reference Example 1. At this time, the effective raw material supply mass flow rate was adjusted to 10.0 kg / hr. From the 47.5-hour point, the inside of reactor 2 was replaced with the raw materials. (switching) The feed rate of raw material to reactor 2 was increased, and at the same time, the feed rate of raw material to reactor 1 was decreased, so that the total effective raw material feed mass flow rate supplied to the two reactors was adjusted to 10.0 kg / hr. After 48 hours, the feed rate of raw material to reactor 1 was stopped, and the feed rate of raw material to reactor 2 was set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0118] 2)48~96 hours (Reaction step) At the 48-hour point, a 48-hour reaction step was started in the reactor 2 in the same manner as in Reference Example 1. At this time, the effective raw material supply mass flow rate was adjusted to 10.0 kg / hr. (Catalyst regeneration process) The zeolite-containing catalyst used in the reaction step was packed in reactor 1, and the atmosphere in reactor 1 was replaced with nitrogen gas until the hydrocarbon concentration was 5.0% by volume or less. Then, the catalyst regeneration step was carried out under the following conditions 1 to 5. Procedure 1) Temperature: 480°C, oxygen concentration: 1%, 1 hour Step 2) Temperature: 520°C, oxygen concentration: 1%, 3 hours Step 3) Temperature: 550°C, oxygen concentration: 1%, 3 hours Step 4) Temperature: 550°C, oxygen concentration: 5%, 1 hour Step 5) Temperature: 580°C, oxygen concentration: 5%, 2 hours In step 5, since carbon dioxide due to coke combustion was no longer detected, it was determined that catalyst regeneration was complete. At this point, the catalyst regeneration process took 10 hours. After step 5 was completed, the reactor 1 was purged with nitrogen gas until the oxygen concentration in the reactor 1 was 2.5% by volume or less. After 95.5 hours, the inside of the nitrogen-purged reactor 1 was purged with raw materials.

[0119] (switching) The feed rate of raw material to reactor 1 was increased, and at the same time, the feed rate of raw material to reactor 2 was decreased, so that the total effective raw material feed mass flow rate supplied to the two reactors was adjusted to 10.0 kg / hr. At the 96-hour point, the feed of raw material to reactor 2 was stopped, and the effective raw material feed mass flow rate to reactor 1 was set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0120] 3)96~144 hours (Reaction step) At the 96-hour point, a 48-hour reaction step was started in the reactor 1 in the same manner as in Reference Example 1. At this time, the effective raw material supply mass flow rate was adjusted to 10.0 kg / hr. (Catalyst regeneration process) With the zeolite-containing catalyst used in the reaction step still packed in reactor 2, the reactor 2 was purged with nitrogen gas until the hydrocarbon concentration therein was 5.0% by volume or less, and then the catalyst regeneration step was carried out. After the coke combustion removal was completed, the reactor 2 was purged with nitrogen gas until the oxygen concentration therein was 2.5% by volume or less. The production rates of ethylene, propylene, and aromatic compounds obtained from the reaction process carried out continuously as described above are shown in Figure 4. Over 144 hours, the propylene fluctuation rate was 8.9 mass%, and the P / E ratio fluctuated by a maximum of 0.13. Furthermore, this example shows that by carrying out the above-mentioned catalyst regeneration process, the catalytic performance of a zeolite-containing catalyst whose activity has been reduced by being subjected to a reaction process can be restored to the same level as before being subjected to the reaction process. This is presumably because the catalyst regeneration process can burn off coke adhering to the catalyst, and the catalyst regeneration process of this example suppresses catalyst deterioration due to structural collapse, but the reasons for this are not limited to these.

[0121] [Comparative Example 1] In Comparative Example 1, the reaction process was carried out continuously using only reactor 1. Raw materials were supplied to reactor 1, and the reaction process was carried out for 144 hours in the same manner as in Reference Example 1. The resulting production rates of ethylene, propylene, and aromatic compounds are shown in Figure 5. At this time, the effective raw material supply mass flow rate was adjusted to 10.0 kg / hr. In this Comparative Example, the propylene fluctuation rate was 23.4 mass% throughout the 144 hours, and the P / E ratio fluctuated by a maximum of 0.29. Comparison between the Example and Comparative Example 1 revealed that when the reaction was carried out continuously using only one reactor, the ethylene conversion rate and P / E ratio fluctuated greatly, and the load on the purification system increased.

[0122] [Example 2] In Example 2, a reaction process was carried out continuously for 144 hours by switching between three reactors consisting of reactor 1, reactor 2, and reactor 3 as follows. The operating status of each reactor in this example is shown in Figure 6.

[0123] 1)0~24 hours (Reaction step) Raw materials were supplied to reactor 1 and reactor 2, and the reaction process was started in the same manner as in Reference Example 1, with reactor 1 being operated for 48 hours and reactor 2 being operated for 24 hours. At this time, the total effective raw material supply mass flow rate supplied to reactor 1 and reactor 2 was adjusted to be 10.0 kg / hr. For reactor 3, the interior of the reactor was replaced with raw materials from the 23.5 hour point. The reaction process in reactor 2 was stopped after 24 hours only at the start of the reaction, but after 24 hours from the start of the reaction after the steady state was reached, the reaction process in each reactor was carried out for 48 hours.

[0124] (switching) The feed rate of raw material to reactor 3 was increased, and at the same time, the feed rate of raw material to reactor 2 was decreased, so that the total effective raw material feed mass flow rate supplied to the three reactors was adjusted to 10.0 kg / hr. At the 24-hour point, the feed of raw material to reactor 2 was stopped, and the total effective raw material feed mass flow rate supplied to reactor 1 and reactor 3 was set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0125] 2)24~48 hours (Reaction step) At the 24-hour point, a 48-hour reaction process was started in Reactor 3 in the same manner as in Reference Example 1. At this time, the total effective raw material feed mass flow rate supplied to Reactor 1 and Reactor 3 was adjusted to 10.0 kg / hr. (Catalyst regeneration process) The zeolite-containing catalyst used in the reaction step was packed in reactor 2, and the hydrocarbon concentration in reactor 2 was replaced with nitrogen gas until it reached 5.0% by volume or less. Then, the catalyst regeneration step was carried out under the following conditions 1 to 5. Procedure 1) Temperature: 480°C, oxygen concentration: 1%, 1 hour Step 2) Temperature: 520°C, oxygen concentration: 1%, 3 hours Step 3) Temperature: 550°C, oxygen concentration: 1%, 3 hours Step 4) Temperature: 550°C, oxygen concentration: 5%, 1 hour Step 5) Temperature: 580°C, oxygen concentration: 5%, 2 hours In step 5, since carbon dioxide due to coke combustion was no longer detected, it was determined that catalyst regeneration was complete. At this point, the catalyst regeneration process took 10 hours. After step 5 was completed, the reactor 2 was purged with nitrogen gas until the oxygen concentration in the reactor 2 was 2.5% by volume or less. After 47.5 hours, the inside of the nitrogen-purged reactor 2 was purged with raw materials.

[0126] (switching) The feed rate of raw material to reactor 2 was increased, and at the same time, the feed rate of raw material to reactor 1 was decreased, so that the total effective raw material feed mass flow rate supplied to the three reactors was adjusted to 10.0 kg / hr. After 48 hours, the feed rate of raw material to reactor 1 was stopped, and the total effective raw material feed mass flow rate to reactors 2 and 3 was adjusted to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0127] 3)48~72 hours (Reaction step) At the 48-hour point, a 48-hour reaction step was started in Reactor 2 in the same manner as in Reference Example 1. At this time, the effective raw material feed mass flow rates supplied to Reactor 2 and Reactor 3 were adjusted to 10.0 kg / hr.

[0128] (Catalyst regeneration process) With the zeolite-containing catalyst used in the reaction step still packed in reactor 1, the reactor 1 was purged with nitrogen gas until the hydrocarbon concentration therein was 5.0% by volume or less, and then the catalyst regeneration step was carried out. After the coke combustion removal was completed, the reactor 1 was purged with nitrogen gas until the oxygen concentration therein was 2.5% by volume or less. The nitrogen-purged reactor 1 was purged with raw material from the 71.5-hour point.

[0129] (switching) The feed rate of raw material to reactor 1 was increased, and at the same time, the feed rate of raw material to reactor 3 was decreased, so that the total effective raw material feed mass flow rate supplied to the three reactors was adjusted to 10.0 kg / hr. 72 hours after the start of the reaction, the feed of raw material to reactor 3 was stopped, and the effective raw material feed mass flow rates to reactors 1 and 2 were set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0130] 3)72~144 hours From 72 to 144 hours, as in the first 72 hours, two reactors were used for the reaction step and one reactor was used for the catalyst regeneration step, thereby continuously carrying out the reaction step. The production rates of ethylene, propylene, and aromatic compounds obtained from the reaction process carried out continuously as described above are shown in Figure 7. Over 144 hours, the propylene fluctuation rate was 4.3 mass%, and the P / E ratio fluctuated by a maximum of 0.13.

[0131] [Example 3] In Example 3, a reaction process was carried out continuously for 144 hours by switching between four reactors consisting of reactor 1, reactor 2, reactor 3, and reactor 4 as follows. The operating status of each reactor in this example is shown in Figure 8.

[0132] 1)0~12 hours (Reaction step) Raw materials were supplied to reactor 1, reactor 2, and reactor 3, and reaction steps were initiated in the same manner as in Reference Example 1: 36 hours in reactor 1, 24 hours in reactor 2, and 12 hours in reactor 3. At this time, the total effective raw material supply mass flow rate supplied to reactors 1, 2, and 3 was adjusted to 10.0 kg / hr. From the 11.5-hour point, the interior of reactor 4 was replaced with raw materials. Only at the start of the reaction, the reaction step in reactor 2 was stopped after 24 hours, and the reaction step in reactor 3 after 12 hours. However, after 24 hours from the start of the reaction after stabilization, the reaction step in each reactor was carried out for 36 hours.

[0133] (switching) The feed rate of raw material to reactor 4 was increased, and at the same time, the feed rate of raw material to reactor 3 was decreased, so that the total effective raw material feed mass flow rate supplied to the four reactors was adjusted to 10.0 kg / hr. At the 12-hour point, the feed of raw material to reactor 3 was stopped, and the total effective raw material feed mass flow rate supplied to reactors 1, 2, and 4 was adjusted to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0134] 2)12~24 hours (Reaction step) At the 12-hour point, a 36-hour reaction process was started in reactor 4 in the same manner as in Reference Example 1. At this time, the total mass flow rate of the effective raw material supplied to reactor 1, reactor 2, and reactor 4 was adjusted to 10.0 kg / hr.

[0135] (Catalyst regeneration process) The zeolite-containing catalyst used in the reaction step was packed in the reactor 3, and the hydrocarbon concentration in the reactor 3 was replaced with nitrogen gas until it reached 5.0% by volume or less. Then, the catalyst regeneration step was carried out under the following conditions 1 to 5. Procedure 1) Temperature: 480°C, oxygen concentration: 1%, 1 hour Step 2) Temperature: 520°C, oxygen concentration: 1%, 3 hours Step 3) Temperature: 550°C, oxygen concentration: 1%, 3 hours Step 4) Temperature: 550°C, oxygen concentration: 5%, 1 hour Step 5) Temperature: 580°C, oxygen concentration: 5%, 2 hours In step 5, since carbon dioxide due to coke combustion was no longer detected, it was determined that catalyst regeneration was complete. At this point, the catalyst regeneration process took 10 hours. After step 5 was completed, the reactor 3 was purged with nitrogen gas until the oxygen concentration in the reactor 3 was 2.5% by volume or less. After 23.5 hours, the inside of the nitrogen-purged reactor 3 was purged with raw materials.

[0136] (switching) The feed rate of raw material to reactor 3 was increased, and at the same time, the feed rate of raw material to reactor 2 was decreased, so that the total effective raw material feed mass flow rate supplied to the four reactors was adjusted to 10.0 kg / hr. At the 24-hour point, the feed of raw material to reactor 2 was stopped, and the total effective raw material feed mass flow rate to reactors 1, 3, and 4 was adjusted to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0137] 3)24~36 hours (Reaction step) At the 24-hour point, a 36-hour reaction process was started in Reactor 3 in the same manner as in Reference Example 1. At this time, the effective raw material feed mass flow rates supplied to Reactor 1, Reactor 3, and Reactor 4 were adjusted to 10.0 kg / hr. (Catalyst regeneration process) With the zeolite-containing catalyst used in the reaction step still packed in reactor 2, the reactor 2 was purged with nitrogen gas until the hydrocarbon concentration therein was 5.0% by volume or less, and then the catalyst regeneration step was carried out. After the coke combustion removal was completed, the reactor 2 was purged with nitrogen gas until the oxygen concentration therein was 2.5% by volume or less. The nitrogen-purged reactor 2 was purged with raw material from the 71.5-hour point.

[0138] (switching) The feed rate of raw material to reactor 2 was increased, and at the same time, the feed rate of raw material to reactor 1 was decreased, so that the total effective raw material feed mass flow rate supplied to the four reactors was adjusted to 10.0 kg / hr. 36 hours after the start of the reaction, the feed of raw material to reactor 1 was stopped, and the effective raw material feed mass flow rates to reactors 2, 3, and 4 were set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%. 3)36~144 hours From 36 to 144 hours, similarly to the period up to 36 hours, three reactors were used for the reaction step and one reactor was used for the catalyst regeneration step, thereby continuously carrying out the reaction step. The production rates of ethylene, propylene, and aromatic compounds obtained from the reaction process carried out continuously as described above are shown in Figure 9. Throughout 144 hours, the propylene fluctuation rate was 2.1 mass%, and the P / E ratio fluctuated by a maximum of 0.03. This example demonstrates that the use of multiple reactors to increase the switching frequency can suppress fluctuations in the production rate of the target compound.

[0139] [Example 4] In Example 4, a reaction process was carried out continuously for 144 hours by switching between four reactors consisting of reactor 1, reactor 2, reactor 3, and reactor 4 as follows. The operating status of each reactor in this example is shown in Figure 10.

[0140] 1)0~24 hours (Reaction step) Raw materials were supplied to reactor 1 and reactor 2, and the reaction process was started in the same manner as in Reference Example 1, for 48 hours in reactor 1 and for 24 hours in reactor 2. At this time, the total effective raw material supply mass flow rate supplied to reactor 1 and reactor 2 was adjusted to 10.0 kg / hr. From the 23.5 hour point, the inside of reactor 4 was replaced with raw materials. Note that the reaction process in reactor 2 was stopped after 24 hours only at the start of the reaction, but from 24 hours after the start of the reaction, the reaction process in each reactor was carried out for 48 hours.

[0141] (switching) The feed rate of raw material to reactor 4 was increased, and at the same time, the feed rate of raw material to reactor 2 was decreased, so that the total effective raw material feed mass flow rate supplied to the three reactors was adjusted to 10.0 kg / hr. At the 24-hour point, the feed rate of raw material to reactor 2 was stopped, and the total effective raw material feed mass flow rate supplied to reactor 1 and reactor 4 was adjusted to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0142] 2)24~48 hours (Reaction step) At the 24-hour point, a 48-hour reaction process was started in reactor 4 in the same manner as in Reference Example 1. At this time, the total mass flow rate of the effective raw material supplied to reactor 1 and reactor 4 was adjusted to 10.0 kg / hr.

[0143] (Catalyst regeneration process) The zeolite-containing catalyst used in the reaction step was packed in reactor 2, and the hydrocarbon concentration in reactor 2 was replaced with nitrogen gas until it reached 5.0% by volume or less. Then, the catalyst regeneration step was carried out under the following conditions 1 to 5. Procedure 1) Temperature: 480°C, oxygen concentration: 1%, 1 hour Step 2) Temperature: 520°C, oxygen concentration: 1%, 3 hours Step 3) Temperature: 550°C, oxygen concentration: 1%, 3 hours Step 4) Temperature: 550°C, oxygen concentration: 5%, 1 hour Step 5) Temperature: 580°C, oxygen concentration: 5%, 2 hours In step 5, since carbon dioxide due to coke combustion was no longer detected, it was determined that catalyst regeneration was complete. At this point, the catalyst regeneration process took 10 hours. After step 5 was completed, the reactor 2 was purged with nitrogen gas until the oxygen concentration in the reactor 2 was 2.5% by volume or less. After 23.5 hours, the inside of the nitrogen-purged reactor 2 was purged with raw materials.

[0144] (switching) The feed rate of the raw material to reactor 3 was increased, and at the same time, the feed rate of the raw material to reactor 1 was decreased, so that the total effective raw material feed mass flow rate supplied to the three reactors was adjusted to 10.0 kg / hr. After 48 hours, the feed of raw material to reactor 1 was stopped, and the total effective raw material feed mass flow rate to reactors 3 and 4 was set to 10.0 kg / hr. This reactor switching operation was completed within one minute.

[0145] 3)48~72 hours (Reaction step) At the 48-hour point, a 48-hour reaction process was started in Reactor 3 in the same manner as in Reference Example 1. At this time, the effective raw material feed mass flow rates supplied to Reactor 3 and Reactor 4 were adjusted to 10.0 kg / hr. (Catalyst regeneration process) With the zeolite-containing catalyst used in the reaction step still packed in reactor 1, the reactor 1 was purged with nitrogen gas until the hydrocarbon concentration therein was 5.0% by volume or less, and then the catalyst regeneration step was carried out. After the coke combustion removal was completed, the reactor 1 was purged with nitrogen gas until the oxygen concentration therein was 2.5% by volume or less. The nitrogen-purged reactor 2 was purged with raw material from the 47.5-hour point.

[0146] (switching) The feed rate of raw material to reactor 2 was increased, and at the same time, the feed rate of raw material to reactor 4 was decreased, so that the total effective raw material feed mass flow rate supplied to the three reactors was adjusted to 10.0 kg / hr. 72 hours after the start of the reaction, the feed of raw material to reactor 4 was stopped, and the effective raw material feed mass flow rates to reactors 2 and 3 were set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0147] 3)72~144 hours From 72 to 144 hours, as in the first 72 hours, two reactors were used for the reaction step, and two reactors were used for the catalyst regeneration step or were on standby, thereby continuously carrying out the reaction step. The production rates of ethylene, propylene, and aromatic compounds obtained from the reaction process carried out continuously as described above are shown in Figure 11. Throughout 144 hours, the propylene fluctuation rate was 4.3 mass%, and the P / E ratio fluctuated by a maximum of 0.07. A comparison between this example and Example 3 reveals that by extending the reaction step time, the switching frequency can be reduced, facilitating the production operation.

[0148] [Example 5] In Example 5, the reaction process was carried out continuously for 144 hours by switching between eight reactors consisting of reactor 1 to reactor 8 as follows. The operating status of each reactor in this example is shown in Figure 12.

[0149] 1)0~12 hours (Reaction step) The raw material was supplied from reactor 1 to reactor 4, and the reaction process was started in the same manner as in Reference Example 1: 48 hours in reactor 1, 36 hours in reactor 2, 24 hours in reactor 3, and 12 hours in reactor 4. At this time, the total effective raw material supply mass flow rate supplied from reactor 1 to reactor 4 was adjusted to 10.0 kg / hr. From the 11.5 hour point, the reactor 8 was substituted with the raw material. Only at the start of the reaction, the reaction process in reactor 2 was stopped after 36 hours, the reaction process in reactor 3 after 24 hours, and the reaction process in reactor 4 after 12 hours, but after 12 hours from the start of the reaction, the reaction process in each reactor was carried out for 48 hours.

[0150] (switching) The feed rate of raw material to reactor 8 was increased, and at the same time, the feed rate of raw material to reactor 4 was decreased, so that the total effective raw material feed mass flow rate supplied to the five reactors was adjusted to 10.0 kg / hr. At the 12-hour point, the feed of raw material to reactor 4 was stopped, and the total effective raw material feed mass flow rate supplied to reactors 1, 2, 3, and 8 was adjusted to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0151] 2)12~24 hours (Reaction step) At the 12-hour point, a 48-hour reaction process was started in reactor 8 in the same manner as in Reference Example 1. At this time, the total effective raw material feed mass flow rate supplied to reactors 1, 2, 3 and 8 was adjusted to 10.0 kg / hr. (Catalyst regeneration process) The zeolite-containing catalyst used in the reaction step was packed in reactor 4, and the hydrocarbon concentration in reactor 4 was replaced with nitrogen gas until it reached 5.0% by volume or less. Then, the catalyst regeneration step was carried out under the following conditions 1 to 5. Procedure 1) Temperature: 480°C, oxygen concentration: 1%, 1 hour Step 2) Temperature: 520°C, oxygen concentration: 1%, 3 hours Step 3) Temperature: 550°C, oxygen concentration: 1%, 3 hours Step 4) Temperature: 550°C, oxygen concentration: 5%, 1 hour Step 5) Temperature: 580°C, oxygen concentration: 5%, 2 hours In step 5, since carbon dioxide due to coke combustion was no longer detected, it was determined that catalyst regeneration was complete. At this point, the catalyst regeneration process took 10 hours. After step 5 was completed, the reactor 4 was purged with nitrogen gas until the oxygen concentration in the reactor 4 was 2.5% by volume or less. After 23.5 hours, the inside of the nitrogen-purged reactor 4 was purged with raw materials.

[0152] (switching) The feed rate of raw material to reactor 7 was increased, and at the same time, the feed rate of raw material to reactor 3 was decreased, so that the total effective raw material feed mass flow rate supplied to the five reactors was adjusted to 10.0 kg / hr. At the 24-hour point, the feed of raw material to reactor 3 was stopped, and the total effective raw material feed mass flow rate to reactors 1, 2, 7, and 8 was adjusted to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0153] 3)24~36 hours (Reaction step) At the 48-hour point, a 48-hour reaction step was started in reactor 7 in the same manner as in Reference Example 1. At this time, the effective raw material feed mass flow rates supplied to reactors 1, 2, 7, and 8 were adjusted to 10.0 kg / hr.

[0154] (Catalyst regeneration process) With the zeolite-containing catalyst used in the reaction step still packed in the reactor 3, the reactor 3 was purged with nitrogen gas until the hydrocarbon concentration therein was 5.0% by volume or less, and then the catalyst regeneration step was carried out. After the coke combustion removal was completed, the reactor 3 was purged with nitrogen gas until the oxygen concentration therein was 2.5% by volume or less. The nitrogen-purged reactor 3 was purged with the raw material from the 35.5-hour point.

[0155] (switching) The feed rate of raw material to reactor 6 was increased, and at the same time, the feed rate of raw material to reactor 2 was decreased, so that the total effective raw material feed mass flow rate supplied to the five reactors was adjusted to 10.0 kg / hr. 36 hours after the start of the reaction, the feed of raw material to reactor 2 was stopped, and the effective raw material feed mass flow rates to reactors 1, 6, 7, and 8 were set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0156] 3)36~144 hours From 36 to 144 hours, as in the first 36 hours, four reactors were used for the reaction process, and four reactors were used for the catalyst regeneration process or were put on standby, thereby continuously carrying out the reaction process. The production rates of ethylene, propylene, and aromatic compounds obtained from the reaction process carried out continuously as described above are shown in Figure 13. Throughout 144 hours, the propylene fluctuation rate was 2.1 mass%, and the P / E ratio fluctuated by a maximum of 0.01.

[0157] [Example 6] In Example 6, the eight reactors consisting of reactor 1 to reactor 8 were divided into two groups, reactor group 1 consisting of reactor 1 to reactor 4 and reactor group 2 consisting of reactor 5 to reactor 8, and switched between these groups as follows, to carry out a continuous reaction process for 144 hours. The operating status of each reactor in this example is shown in Figure 14.

[0158] 1)0~48 hours (Reaction step) The raw materials were supplied to Reactor Group 1, and a 48-hour reaction process was initiated in the same manner as in Reference Example 1. At this time, the effective raw material supply mass flow rate was adjusted to 10.0 kg / hr. In Reactor Group 2, the interior of the reactor was replaced with the raw materials from the 47.5-hour point.

[0159] (switching) The feed rate of raw material to reactor group 2 was increased, and at the same time, the feed rate of raw material to reactor group 1 was decreased, so that the total effective raw material feed mass flow rate supplied to the eight reactors was adjusted to 10.0 kg / hr. At the 48-hour point, the feed rate of raw material to reactor group 1 was stopped, and the feed rate of raw material to reactor group 2 was set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0160] 2)48~96 hours (Reaction step) At the 48-hour point, a 48-hour reaction step was started in reactor group 2 in the same manner as in Reference Example 1. At this time, the mass flow rate of the effective raw material supply was adjusted to 10.0 kg / hr. (Catalyst regeneration process) The zeolite-containing catalyst used in the reaction step was packed in the reactor group 1, and the atmosphere in the reactor group 1 was replaced with nitrogen gas until the hydrocarbon concentration became 5.0% by volume or less. Then, the catalyst regeneration step was carried out under the following conditions 1 to 5. Procedure 1) Temperature: 480°C, oxygen concentration: 1%, 1 hour Step 2) Temperature: 520°C, oxygen concentration: 1%, 3 hours Step 3) Temperature: 550°C, oxygen concentration: 1%, 3 hours Step 4) Temperature: 550°C, oxygen concentration: 5%, 1 hour Step 5) Temperature: 580°C, oxygen concentration: 5%, 2 hours In step 5, since carbon dioxide due to coke combustion was no longer detected, it was determined that catalyst regeneration was complete. At this point, the catalyst regeneration process took 10 hours. After step 5 was completed, reactor group 1 was purged with nitrogen gas until the oxygen concentration in reactor group 1 was 2.5% by volume or less. From the 95.5-hour mark, the interior of reactor group 1 was purged with nitrogen and the raw material was purged.

[0161] (switching) The feed rate of raw material to reactor group 1 was increased, and at the same time, the feed rate of raw material to reactor group 2 was decreased, so that the total effective raw material feed mass flow rate supplied to the eight reactors was adjusted to 10.0 kg / hr. At the 96-hour point, the feed rate of raw material to reactor group 2 was stopped, and the effective raw material feed mass flow rate to reactor group 1 was set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0162] 3)96~144 hours (Reaction step) At the 96-hour point, a 48-hour reaction step was started in Reactor Group 1 in the same manner as in Reference Example 1. At this time, the effective raw material supply mass flow rate was adjusted to 10.0 kg / hr. (Catalyst regeneration process) With the zeolite-containing catalyst used in the reaction step still packed in reactor group 2, the reactor group 2 was purged with nitrogen gas until the hydrocarbon concentration therein reached 5.0% by volume or less, and then the catalyst regeneration step was carried out. After the coke combustion removal was completed, the reactor group 2 was purged with nitrogen gas until the oxygen concentration therein reached 2.5% by volume or less. The production rates of ethylene, propylene, and aromatic compounds obtained from the reaction process carried out continuously as described above are shown in Figure 15. Over 144 hours, the propylene fluctuation rate was 8.9 mass%, and the P / E ratio fluctuated by a maximum of 0.13.

[0163] [Example 7] This example was carried out in the same manner as in Example 1, except that zeolite-containing catalyst 2 was used in the same manner as in Reference Example 2. The production rates of ethylene, propylene, and aromatic compounds obtained in the reaction step are shown in Figure 16. Over 144 hours, the propylene fluctuation rate was 6.3 mass%, and the P / E ratio fluctuated by a maximum of 0.10. This example shows that even when the zeolite-containing catalyst contains silver, by carrying out a catalyst regeneration process, the catalytic performance of the zeolite-containing catalyst, whose activity has been reduced by being subjected to a reaction process, can be restored to a state similar to that before being subjected to the reaction process.

[0164] [Example 8] This example was carried out in the same manner as in Example 1, except that the reaction time was changed from 48 hours to 72 hours and the switching frequency was changed from 0.5 times / day to 0.7 times / day. Over 144 hours, the propylene fluctuation rate was 13.0 mass%, and the P / E ratio fluctuated by a maximum of 0.21. The operating status of each reactor in this example is shown in Figure 17, and the production rates of ethylene, propylene, and aromatic compounds obtained from the reaction process are shown in Figure 18.

[0165] [Example 9] This example was carried out in the same manner as in Example 1, except that the conditions for the catalyst regeneration step were changed as follows: Over 144 hours, the propylene fluctuation rate was 8.9 mass%, and the P / E ratio fluctuated by a maximum of 0.13. (Catalyst regeneration process) The reactor was filled with the zeolite-containing catalyst used in the reaction process, and the reactor was purged with nitrogen gas until the hydrocarbon concentration in the reactor was 5.0% by volume or less. Coke combustion was then initiated at a temperature of 440°C and an oxygen concentration of 5%. The regeneration gas at the reactor outlet was analyzed every four hours, and it took 32 hours for carbon dioxide associated with coke combustion to be detected. The reactor was then purged with nitrogen gas until the oxygen concentration in the reactor was 2.5% by volume or less. The reactor was then further purged with raw material. In this example, coke combustion was carried out at a lower temperature than in the catalyst regeneration step of Example 1, but it took 32 hours to complete the regeneration. A longer time required for the catalyst regeneration step is undesirable because it impairs the ease of process operation. Furthermore, although coke combustion at low temperatures can be expected to cause catalyst degradation during the catalyst regeneration step, since the target compound production rates in this example and Example 1, in which high-temperature treatment was performed, are similar, it is preferable to perform the catalyst regeneration step at 450°C or higher in the ethanol conversion method according to this embodiment. The total fluctuation in the effective raw material supply mass flow rate during the switching operation was 0%.

[0166] [Example 10] This example was carried out in the same manner as in Example 1, except that in the reaction termination step and reaction restart step, the feed rate to the first reactor was reduced to 9.50 kg / hr, and then the feed rate to the second reactor was increased to switch the reactor. At this time, the feed rate of the raw material was reduced by 5% from the maximum flow rate. This reactor switching operation was completed within 1 minute. Over 144 hours, the propylene fluctuation rate was 14.6 mass%, and the P / E ratio fluctuated by a maximum of 0.14. The production rates of ethylene, propylene, and aromatic compounds obtained from the reaction steps of this example are shown in Figure 19.

[0167] [Example 11] In this example, a reaction process was carried out continuously for 144 hours by switching between two reactors consisting of reactor 1 and reactor 2 as follows. The operating status of each reactor in this example is shown in Figure 20.

[0168] 1)0~48 hours (Reaction step) The raw materials were supplied to reactor 1, and a 48-hour reaction process was initiated in the same manner as in Reference Example 1. At this time, the effective raw material supply mass flow rate was adjusted to 10.0 kg / hr. From the 47.5-hour point, the inside of reactor 2 was replaced with the raw materials. (Switching between reaction stop and reaction restart processes) The feed rate of raw material to reactor 2 was increased, and at the same time, the feed rate of raw material to reactor 1 was decreased, so that the total effective raw material feed mass flow rate supplied to the two reactors was adjusted to 10.0 kg / hr. After 48 hours, the feed rate of raw material to reactor 1 was stopped, and the feed rate of raw material to reactor 2 was set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0169] 2)48~96 hours (Reaction step) At the 48-hour point, a 48-hour reaction step was started in the reactor 2 in the same manner as in Reference Example 1. At this time, the effective raw material supply mass flow rate was adjusted to 10.0 kg / hr. (Catalyst regeneration process) The zeolite-containing catalyst used in the reaction step was packed in the reactor 1, and the hydrocarbon concentration in the reactor 1 was replaced with nitrogen gas until it reached 5.0% by volume or less. Then, the catalyst regeneration step was carried out under the following conditions 1 and 2. Procedure 1) Temperature: 700°C, oxygen concentration: 1%, 1 hour Step 2) Temperature: 700°C, oxygen concentration: 5%, 4 hours In step 2, since carbon dioxide due to coke combustion was no longer detected, it was determined that catalyst regeneration was complete. At this point, the catalyst regeneration process had taken 5 hours. Furthermore, in step 1, a sudden heat generation due to coke combustion was observed. After step 2 was completed, reactor 1 was purged with nitrogen gas until the oxygen concentration in reactor 1 was 2.5% by volume or less. From the 95.5-hour mark, the interior of the nitrogen-purged reactor 1 was purged with raw materials.

[0170] (switching) The feed rate of raw material to reactor 1 was increased, and at the same time, the feed rate of raw material to reactor 2 was decreased, so that the total effective raw material feed mass flow rate supplied to the two reactors was adjusted to 10.0 kg / hr. At the 96-hour point, the feed of raw material to reactor 2 was stopped, and the effective raw material feed mass flow rate to reactor 1 was set to 10.0 kg / hr. This reactor switching operation was completed within 1 minute. The fluctuation in the total effective raw material feed mass flow rate during the switching operation was 0%.

[0171] 3)96~129 hours (Reaction step) At the 96th hour, the reaction process was initiated in reactor 1 in the same manner as in Reference Example 1. At this time, the effective raw material supply mass flow rate was adjusted to 10.0 kg / hr. However, with the catalyst that had undergone the catalyst regeneration process in this comparative example, the ethylene production rate increased and the propylene production rate decreased from the early stage of the reaction, compared with the start of the reaction process from 0 to 48 hours. At the 129th hour, the ethylene production rate became equivalent to that at 48 hours, so the reaction process in reactor 1 was terminated and switched to reactor 2. (Catalyst regeneration process) With the zeolite-containing catalyst used in the reaction step still packed in reactor 2, the reactor 2 was purged with nitrogen gas until the hydrocarbon concentration therein was 5.0% by volume or less, and then the catalyst regeneration step was carried out. After the coke combustion removal was completed, the reactor 2 was purged with nitrogen gas until the oxygen concentration therein was 2.5% by volume or less. After 128.5 hours, the inside of the nitrogen-purged reactor 2 was purged with raw materials.

[0172] 4)129~144 hours (Reaction step) From the 129th hour, the reaction process was started in reactor 2 in the same manner as in Reference Example 1. At this time, the effective raw material supply mass flow rate was adjusted to 10.0 kg / hr. In reactor 2, as in the reaction process from 96 to 129 hours, the ethylene production rate increased and the propylene production rate decreased. The reaction process was terminated when 144 hours had elapsed. (Catalyst regeneration process) With the zeolite-containing catalyst used in the reaction step still packed in reactor 1, the reactor 2 was purged with nitrogen gas until the hydrocarbon concentration therein was 5.0% by volume or less, and then the catalyst regeneration step was carried out. After the coke combustion removal was completed, the reactor 2 was purged with nitrogen gas until the oxygen concentration therein was 2.5% by volume or less. The production rates of ethylene, propylene, and aromatic compounds obtained from the reaction steps that were continuously carried out are shown in Figure 21. As described above, by setting the treatment temperature in the catalyst regeneration step to 650°C or higher, catalyst degradation became significant, and it was necessary to increase the switching frequency in order to maintain the ethylene production rate at a certain level or lower. Furthermore, the reaction time shown in Table 1 is 33 hours, which is the shortest cycle in this example, out of 96 to 129 hours. Thus, the operation of increasing the switching frequency during operation places a very high operational load on the process.

[0173] [Table 1]

[0174] [Example 12] This example is carried out using the purification equipment shown in Fig. 22. When ethylene and propylene are purified from a reaction gas produced under the same conditions and with the same switching operation as in Example 1, the processing capacity of the ethylene purification tower is constant, while the ethylene production rate increases from the start to the stop of the reaction process. Therefore, it is necessary to reduce the feed rate of the raw material in accordance with the fluctuations in the ethylene production rate, or to provide a buffer zone in the processing capacity of the ethylene purification tower. In Example 1, the ethylene production rate varied from 2.50 kg / hr to 2.91 kg / hr. Using the minimum ethylene production rate of 2.50 kg / hr as a reference, a 5% buffer zone was set in the throughput of the ethylene purification column, yielding 2.50 × 1.05 = 2.63 kg / hr. The feed rate of the raw material must be reduced once the ethylene production rate reaches 2.63 kg / hr. When the reaction process is stopped, at which the ethylene production rate is at its maximum, the feed rate of the raw material must be set to 9.04 kg / hr, resulting in a propylene fluctuation rate of 20.4%. Similarly, when considering the buffer zone of the ethylene purification column at 10%, 15%, and 20%, a 20% buffer zone must be set in the throughput of the ethylene purification column to minimize the propylene fluctuation rate. The buffer zone and propylene fluctuation rate values for the ethylene purification column are shown in Table 2. Since it is difficult to vary the feed rate of the raw materials during the reaction steps, it is preferable to provide a buffer zone in the processing capacity of the ethylene purification column. However, the larger the buffer zone, the greater the amount of energy required to operate the ethylene purification column, and therefore the lower the energy efficiency in ethylene production.

[0175] [Example 13] This example was carried out in the same manner as in Example 12, except that the reaction gas produced in the same manner as in Example 8 was used. The buffer zone and propylene fluctuation rate of the ethylene purification column are shown in Table 2. A comparison between this example and Example 12 reveals that an increase in the ethylene fluctuation rate increases the buffer zone required in the ethylene purification column, resulting in a decrease in the energy efficiency of the process.

[0176] Comparative Example 2 The effect of the ethylene fluctuation rate was evaluated in the same manner as in Example 8, except that the reaction gas produced in the same manner as in Comparative Example 1 was used. In Comparative Example 1, the ethylene production rate fluctuated from 2.50 kg / hr to 3.62 kg / hr. This fluctuation in the ethylene production rate could not be absorbed even when the buffer zone of the ethylene purification column was set to 20%, and the propylene fluctuation rate in this case was 48.9%. Although it is possible to absorb the fluctuation in the ethylene production rate by further increasing the buffer zone of the ethylene purification column, the amount of energy required to operate the ethylene purification column would increase, and the energy efficiency in ethylene production would decrease. Comparison of Example 8 with this Comparative Example shows that by suppressing fluctuations in the production rate of the target compound, the buffer region in the purification system can be reduced, enabling the target compound to be produced efficiently.

[0177] [Table 2]

[0178] From the results of the Examples and Comparative Examples, it can be seen that the ethanol conversion method according to the present embodiment has low fluctuations in the production rate of the product, and can be operated without significantly changing the conditions of purification equipment such as distillation columns based on fluctuations in the production rate, thereby reducing the operational load on the purification system.

Claims

1. A method for continuously converting a raw material containing ethanol into a reaction gas containing olefins having 3 or more carbon atoms, using two or more fixed-bed reactors each equipped with a zeolite catalyst, comprising: a reaction step of supplying the raw material to at least one of the fixed-bed reactors, bringing the raw material into contact with the zeolite-containing catalyst, and obtaining the reaction gas; a fixed-bed reactor for supplying the raw material being switched within a time period in which the propylene production rate in the reaction step varies by 15% or less from the initial production rate;

2. 2. The method for producing ethanol according to claim 1, wherein, in switching the fixed-bed reactor, a fluctuation in the total effective raw material supply mass flow rate of the raw material to the fixed-bed reactor related to the switching is kept within 5%.

3. the reaction gas contains ethylene, 2. The method for converting ethanol according to claim 1, wherein a fluctuation in the rate of ethylene production from the fixed-bed reactor due to switching of the fixed-bed reactor is kept within 20%.

4. 2. The method for converting ethanol according to claim 1, wherein, in switching the fixed-bed reactor, supply of the raw material to the fixed-bed reactor that is currently in operation is stopped after supply of the raw material to the fixed-bed reactor that is the switching destination is started.

5. Three or more fixed bed reactors are used, 2. The method for converting ethanol according to claim 1, wherein in the reaction step, the raw material is introduced into at least two of the fixed-bed reactors.

6. 2. The method for producing ethanol according to claim 1, further comprising a catalyst regeneration step of introducing an oxygen-containing regeneration gas into a reactor different from the fixed-bed reactor used in the reaction step, and combusting coke that has adhered to the catalyst in the reaction step at a temperature in the reactor of 650°C or lower.

7. 7. The method for converting ethanol according to claim 6, further comprising, prior to the catalyst regeneration step, a hydrocarbon purging step of flowing an inert gas through the fixed-bed reactor to reduce the hydrocarbon concentration in the fixed-bed reactor to 5.0% by volume or less.

8. 7. The method for converting ethanol according to claim 6, further comprising, after the catalyst regeneration step, an oxygen purging step of supplying an inert gas to the fixed-bed reactor to reduce the oxygen concentration in the fixed-bed reactor to 2.5% by volume or less.

9. 7. The method for converting ethanol according to claim 6, further comprising an inert gas purging step of supplying a purge gas mainly containing hydrocarbons to the fixed-bed reactor after the oxygen purging step, thereby reducing the concentration of inert gas in the reactor to 10.0 vol % or less.

10. The time T R and the time T required for the catalyst regeneration step in one fixed bed reactor. C But, T R ≧T C 7. The method for converting ethanol according to claim 6, wherein the above formula (1) is satisfied.

11. 7. The method for converting ethanol according to claim 6, wherein in the catalyst regeneration step, the coke is combusted at a reactor temperature of 450°C or higher.

12. 7. The method for converting ethanol according to claim 6, wherein the oxygen content of the regeneration gas is 20% by volume or less.

13. 7. The method for converting ethanol according to claim 6, wherein the fixed-bed reactor has an operating rate for the reaction step of 50% or more.

14. 2. The method for converting ethanol according to claim 1, wherein the silica / alumina molar ratio of the zeolite contained in the zeolite-containing catalyst is 20 to 3,000.

15. 2. The method for converting ethanol according to claim 1, wherein the zeolite-containing catalyst comprises elemental silver or elemental phosphorus.

16. 2. The method for converting ethanol according to claim 1, wherein the amount of the raw material supplied per one fixed-bed reactor is 15,000 kg / hr or less.

17. C further comprises a purification step of purifying ethylene and propylene from the reaction gas obtained by the ethanol conversion method according to claim 1. 2-3 A method for producing olefins.

18. The method includes a combining step of introducing hydrocarbons having 4 or more carbon atoms into a cracking furnace and combining the cracked components or a fraction derived therefrom with the reaction gas or a fraction derived therefrom to obtain a combined fraction, 18. The process according to claim 17, wherein the combined fraction is fed to the purification step. 2-3 A method for producing olefins.

19. 1. A method for regenerating a catalyst in a process for continuously converting an ethanol-containing feedstock into a reaction gas containing olefins having 3 or more carbon atoms using two or more fixed-bed reactors each equipped with a zeolite catalyst, comprising: A method for regenerating a catalyst, comprising a catalyst regeneration step of introducing an oxygen-containing regeneration gas into a fixed-bed reactor containing a catalyst to which coke has adhered, and combusting the coke at a temperature in the reactor of 650°C or less.

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