JP2023134125A - Fixed bed catalyst reactor, plant with fixed bed catalyst reactor, and reaction method with fixed bed catalyst - Google Patents

Abstract

To provide a reactor and a reaction method of methanol capable of increasing a space-time yield.SOLUTION: A reactor 1 for reacting a first gas and a second gas in the presence of a catalyst, comprises: a reaction tube 8; a feed tube 10 arranged within the reaction tube 8 and provided with a large number of distributed side holes 9; a catalyst fixed layer 11 provided with a catalyst arranged between the reaction tube 8 and the feed tube 10; a first supplying section 12 for supplying a first gas between the reaction tube 8 and the feed tube 10 from one end side; a second supply section 13 that supplies the second gas to the feed pipe 10; a heating and cooling section 14 that heats or cools the reaction tube 8; and a discharge section 15 that discharges the reaction mixture gas from the other end side, wherein with a fixed bed catalyst reactor 1 in which while flowing the first gas in one direction to the catalyst fixed bed 11, the second gas is supplied to positions distributed in one direction to the catalyst fixed layer 11 from the feed tube 10 to react carbon dioxide and/or carbon dioxide and hydrogen.SELECTED DRAWING: Figure 1

Description

The present invention relates to a fixed bed catalytic reactor, a plant using a fixed bed catalytic reactor, and a reaction method using a fixed bed catalyst.

Industrial methanol synthesis is carried out by reacting synthesis gas, whose main components are carbon monoxide (CO), carbon dioxide (CO ₂ ), and hydrogen (H ₂ ) obtained by reforming fossil fuels, on a catalyst. (Non-patent document 1). In recent years, as part of global warming countermeasures, studies have been conducted to synthesize carbon dioxide (CO ₂ ) recovered from exhaust gas and the like by reacting with H ₂ (Non-Patent Documents 2 and 3). In this case, the following three reactions occur in parallel.

If we define an equilibrium reaction as a reaction in which a non-negligible amount of starting material remains in the reaction gas even after reaching equilibrium under realistic reaction conditions, all three of these reactions are equilibrium reactions. However, unreacted CO ₂ , CO, and H ₂ remain in the reactor outlet gas. Therefore, it is industrially essential to separate the products methanol (CH ₃ OH) and water (H ₂ O) from the reactor outlet gas and recycle the remaining gas to the reactor, so that it can be newly supplied. In a steady state where the balance between various raw materials and the amount flowing out of the reaction system is maintained, the newly supplied CO ₂ and H ₂ eventually become almost all CH ₃ OH and H ₂ O. Some methane is also produced as a by-product, but the amount is negligible. Therefore, the entire reaction system consists of only the reaction expressed by (1), which is an exothermic reaction.

It is industrially preferable to recover the reaction heat as steam by exchanging heat with boiling water, and this method is used in production from CO, and this direction is also being considered in the synthesis process from _CO2 . Boiling water cooling also has the advantage of increasing the heat transfer coefficient on the cooling side. The reactor is generally a fixed bed catalytic reactor with a shell-and-tube heat exchanger type.

Since this is an exothermic reaction that uses a solid catalyst, when a fixed bed reactor is used, even if it has a cooling function, it is unavoidable that some parts of the catalyst bed will have the highest temperature. Solid catalysts have a maximum temperature at which they can be used, and their activity decreases rapidly above that temperature. The usable temperature of the Cu-ZnO-Al ₂ O ₃ catalyst, which is currently most commonly used in this reaction, is 270°C or lower (Non-Patent Document 3), and the maximum temperature must be kept below this value in order to manage the reaction. This is especially important.

The production efficiency of a reactor/reaction system is expressed as space-time yield (STY). It is the production amount per unit volume of catalyst in the reactor and per unit time. In order to increase STY, it is necessary to increase the total value of _CO2 and CO newly supplied to the reaction system (hereinafter referred to as makeup CO2/CO), but this increases the reaction heat and increases the maximum temperature. It is also about letting people do it. Maximizing STY while suppressing the maximum temperature below the allowable value is an important item in reactor/reaction system design, and various countermeasures are being considered for the entire reaction system. These include an increase in the amount of recycled gas and a decrease in the temperature of boiling water supplied for heat removal.

Among the measures, increasing the recycling ratio (the ratio of the amount of newly supplied gas to the amount of recycled gas) has a great effect, but the larger the ratio, the more energy required in the reaction system and the reaction system configuration. This increases the capacity of the essential heat exchanger, condenser, and booster, which increases costs. This measure has these major problems. Currently, in industrial practice, when the raw material is mainly CO, a recycling ratio of 4 to 6 is selected (Non-patent Document 1), and when the raw material is mainly _CO2 , it is considered to be around 4 (Non-patent Document 3). There is.
Lowering the boiling water temperature is effective in lowering the maximum temperature, but it also lowers the temperature of the entire reactor catalyst layer, resulting in a decrease in STY.

Since it is difficult to take countermeasures in a reaction system with a single reactor, various efforts have been made to combine reaction systems including multiple reactors, heat exchangers, condensers, boosters, etc. (Patent Document 5). However, although this measure has some effect in terms of energy saving, it causes a significant increase in equipment costs, and furthermore, operation management becomes complicated, and even from the examples in patent documents, no increase in STY has been obtained. No significant improvement has been achieved.
The importance of producing methanol from _CO2 will increase in the future as a measure against global warming. Catalysts used for this purpose have been actively investigated, and some catalysts with higher activity values than the catalysts described in the present invention have been announced (Patent Document 6). However, the higher the catalyst activity, the higher the maximum temperature under the same conditions. Therefore, means that can control the maximum temperature will become increasingly important in the future. This is common to all cases in which exothermic and equilibrium reactions are carried out in a fixed bed catalytic reactor.

Chemical Engineering Vol. 46, No. 9, 55-65 (1982) Journal of catalysis Vol. 107, 165-172 (1987) Mitsubishi Heavy Industries Technical Report Vol. 35, No. 6, 384-387 (1998) Catalyst Vol. 35, No. 6,485-401387 (1993)

JP2020-63193A Japanese Patent Application Publication No. 6-312138

An object of the present invention is to provide a fixed bed catalytic reactor that can increase STY while suppressing the maximum temperature to a permissible value or less when performing the gas-solid catalytic reaction in the fixed bed catalytic reactor.

The fixed bed catalytic reactor according to claim 1 is a reactor for reacting a first gas and a second gas that cause an exothermic reaction in the presence of a catalyst, and includes a reaction tube, and a reactor disposed within the reaction tube, A feed pipe provided with a large number of dispersed holes, a fixed catalyst bed provided with the catalyst disposed between the reaction pipe and the feed pipe, and between the reaction pipe and the feed pipe. a first supply section that supplies a first gas from one end side; a second supply section that supplies a second gas to the feed tube; a heating and cooling section that heats or cools the reaction tube; a discharge section for discharging the reaction mixture gas between the feed pipe and the catalyst from the other end side, while causing the first gas to flow in one direction to the catalyst fixed bed, and flowing the second gas from the feed pipe to the catalyst fixed bed. It is characterized in that it is supplied to dispersed positions with respect to the fixed layer and reacted.

A fixed bed catalytic reactor according to a second aspect of the present invention includes, in addition to the configuration according to the first aspect, a heat exchange chamber of the heating/cooling section is partitioned within the reactor shell, and the fixed bed catalytic reactor has a heat exchange chamber of the heating/cooling section defined on both sides of the heat exchange chamber. A supply chamber of the first supply section and a discharge chamber of the discharge section are partitioned, and a plurality of reaction tubes housing the feed pipe and the fixed catalyst bed are arranged in the heat exchange chamber, and a plurality of reaction tubes are arranged in the heat exchange chamber. The reactor is characterized in that both ends thereof communicate with the supply chamber and the discharge chamber in a ventilable manner.

The fixed bed catalytic reactor according to claim 3 has the configuration according to claim 2, and further, the heating/cooling section supplies pressurized boiling water to the heat exchange chamber and discharges the pressurized boiling water from the heat exchange chamber. It is characterized by a structure that includes a heat recovery section that recovers reaction heat by discharging a multiphase flow of water vapor and water vapor.

In addition to the configuration of claim 1 or 2, the plant according to claim 4 includes the fixed bed catalytic reactor, a separation section that separates a reaction product component from the reaction mixture gas, and a reaction product component that separates the reaction product component from the reaction mixture gas. and a circulation section that circulates the separated circulation gas to the first supply section.

In the plant according to claim 5, in addition to the configuration of claim 4, the first gas is a hydrogen-containing gas, the circulating gas, or a circulating gas to which hydrogen is added, and the second gas is carbon dioxide gas or hydrogen-added circulating gas. and the reaction product is methanol.

The reaction method according to claim 6 is a reaction method in which a first gas and a second gas that cause an exothermic reaction in the presence of a catalyst are supplied to a fixed catalyst bed having the catalyst while being cooled and reacted. A feed pipe in which gases are dispersedly provided is arranged along the catalyst fixed bed, and while the first raw material gas flows downstream of the catalyst fixed bed, the second gas is passed from the feed pipe to the catalyst fixed bed. It is characterized by being distributed and supplied to the downstream side of the gas flow of the fixed bed for reaction.

According to the invention described in claim 1, it is possible to provide a fixed bed catalytic reactor that can increase STY while suppressing the maximum temperature to a permissible value or less when performing the gas-solid catalytic reaction in the fixed bed catalytic reactor. can. According to claim 2, the size of the reactor of the present invention can be increased. According to claim 3, the reaction heat from the reactor of the present invention is effectively recovered. According to claim 4, the present invention can be applied to equilibrium reactions in general. According to claim 5, the present invention can be applied to a methanol synthesis plant. According to the sixth aspect, it becomes possible to efficiently carry out an exothermic reaction in the fixed catalyst bed.

1 is a flow sheet of a plant using a fixed bed catalytic reactor according to a first embodiment of the present invention, and shows a case where there is one reaction tube. FIG. 2 is a schematic cross-sectional view showing a fixed bed catalytic reactor according to a second embodiment of the present invention. FIG. 2 is a schematic cross-sectional view showing a conventional fixed bed catalytic reactor in a comparative example. It is a graph showing the result of determining the degree of agreement between the calculated value of the rate formula used in the simulation of the example and the measured value of the reaction rate. 7 is a graph showing the results of a simulation of the temperature distribution in the downstream direction of the fixed catalyst bed in Example 2 and the temperature distribution in Reference Example 2, which has the same STY.

Embodiments of the present invention will be described below with reference to the drawings. Although it can be applied to all exothermic reactions, in this embodiment, an example will be explained in which it is applied to a methanol synthesis method in which CO, CO ₂ and H ₂ are reacted on a solid catalyst.

[First embodiment]
(Plant configuration)
As shown in FIG. 1, there is a fixed bed catalytic reactor 1 in which a first gas and a second gas that cause an exothermic reaction in the presence of a catalyst are reacted, a first gas supply line 2 that newly supplies the reactor 1, and a new a second gas supply line 3 that supplies the reaction mixture gas to the reactor 1, an outlet gas line 4 that discharges the reaction mixture gas from the reactor 1, a separation section 5 that separates the reaction components from the reaction mixture gas, and a separation section 5 that separates the reaction product components from the reaction mixture gas. This plant has a circulation part 6 that circulates the recycled gas to the reactor 1 through a circulation line 6a, and an extraction line 7 that extracts reaction components from the reaction system.
When applied to methanol synthesis, the first gas is a hydrogen-containing gas, a circulating gas, or a circulating gas to which hydrogen is added, and the second gas is a carbon dioxide-containing gas or a carbon dioxide-containing gas to which hydrogen is added, and the reaction product component becomes methanol or methanol and water.

(Configuration of reactor of first embodiment)
The fixed bed catalytic reactor 1 includes a reaction tube 8, a feed tube 10 disposed within the reaction tube 8 and provided with a large number of side holes 9 at dispersed positions, and a reaction tube 8 and a feed tube 10 arranged between the reaction tube 8 and the feed tube 10. A catalyst fixed bed 11 in which a catalyst is arranged, a first supply section 12 that supplies a first gas from one end between the reaction tube 8 and the feed tube 10, and a second gas that supplies the second gas to the feed tube 10. a second supply section 13 for heating or cooling the reaction tube 8, a heating and cooling section 14 for heating or cooling the reaction tube 8, and a discharge section 15 for discharging the reaction mixture gas between the reaction tube 8 and the feed tube 10 from the other end side. This is reactor 1.

(Operation of the reactor of the first embodiment)
H ₂ (make-up H ₂ ) newly supplied to the reaction system is pressurized and temperature-controlled, merges with the circulating gas, and flows as a first gas from the first supply section 12 installed in the reactor 1 to the reaction tube 8. The catalyst is supplied to the catalyst fixed bed 11 inside and flows downstream. CO ₂ or CO ₂ /CO newly supplied to the reaction system is supplied as a second gas from the second supply section 13 to the feed pipe 10 installed in the catalyst layer. This feed pipe 10 is provided with a large number of side holes 9 at dispersed positions in the flow direction of the first gas, and the second gas is distributed and supplied into the first gas in the catalyst layer, and both are distributed to the downstream side. It flows, passes through the catalyst layer, and reacts. The gas that has passed through the catalyst layer is discharged from the discharge section 15 at the other end of the reaction tube 8 as a reaction mixture gas. The reaction heat is removed by a heating medium flowing through the heating and cooling section 14 outside the reaction tube 8. After the reaction mixture gas is separated from the reaction products methanol and H ₂ O in the separation section 5, unreacted CO ₂ , CO, and H ₂ are circulated to the reactor 1 as a circulating gas. The separated methanol and H ₂ O are sent from the reaction component extraction line 7 to the purification system of the plant.
Make-up _H2 may be supplied from the side hole together with the second gas if it is preferable due to the side hole design. The description here is unified based on the method of supplying the gas by merging it with the circulating gas.

(Operations and effects of the first embodiment)
This will be explained when applied to methanol synthesis. The sum of CO ₂ and CO is expressed as CO ₂ /CO. In the conventional reaction system, make-up CO ₂ /CO, make-up H ₂ and circulating gas are supplied together to the reactor 1 from the first supply section 12, but in the present invention, the make-up CO 2 /CO, the make-up H 2 and the circulating gas are supplied together from the first supply section 12 and the second supply section. 13 are supplied separately. The makeup CO ₂ /CO is distributed and supplied to the gas flowing through the catalyst layer through the side hole 9 on the side of the feed pipe 10 . As a result, the next CO ₂ /CO is supplied after a considerable proportion of the supplied CO ₂ /CO has reacted, and the CO ₂ /CO in the reaction gas is CO concentration is smoothed. Since the reaction heat generated increases as the CO ₂ /CO concentration increases, the temperature rise is smoothed and the temperature peak value is significantly reduced compared to the conventional method. Therefore, under the same reaction conditions, the amount of CO ₂ /CO that can be supplied can be increased compared to the conventional method.

[Second embodiment]
(Configuration of reactor of second embodiment)
The reactor is a vertical reactor as shown in FIG. A supply chamber 17 of the section 12 and a discharge chamber 18 of the discharge section 15 are partitioned.
In the heat exchange chamber 16, a plurality of reaction tubes 8 similar to those in the first embodiment are arranged in parallel, accommodating the feed tubes 10 and the fixed catalyst bed 11, to form a group of reaction tubes 8. The second supply section 13, which has both ends communicating with the supply chamber 17 and the discharge chamber 18 in a ventilated manner, is provided with a supply header pipe 13a that can supply the second gas to each feed pipe 10 in the plurality of reaction tubes 8. It is being
The heating/cooling section 14 also includes an inlet nozzle 14a that supplies pressurized boiling water as a heating medium to the lower part of the heat exchange chamber 16 surrounding the reaction tube 8, and an inlet nozzle 14a that supplies boiling water from the upper part of the heat exchange chamber 16. An outlet 14b is provided for discharging the multiphase flow of water vapor, and the reaction heat can be recovered by a heat recovery section (not shown).
The rest is the same as the first embodiment.

(Operation of second embodiment)
The first gas is supplied from the first supply section 12 to a first supply section supply chamber 17 partitioned within the reactor 1, and flows into a plurality of reaction tubes 8. The second gas is distributed and supplied into the gas passing through the catalyst layer from the second gas supply header pipe 13a connected to the second supply part 13 through the side holes installed in each feed pipe 10, and flows downstream. . The gases that have passed through each reaction tube 8 join together in the discharge chamber 18 and flow out from the discharge section 15. A heat exchanger chamber 16 is partitioned and installed on the shell side to remove the heat of reaction, and pressurized boiling water is supplied from the heat medium inlet 14a, a part of which is vaporized into steam by the heat of reaction, and the heat medium is A multiphase flow of boiling water and steam flows out from the outlet 14b.

(Operations and effects of the second embodiment)
In the conventional reaction method, the first gas, the second gas, and the circulating gas are combined before entering the reactor 1 and are supplied to the first supply section supply chamber 17, but in the present invention, the first gas, the second gas, and the circulating gas are connected to the first supply chamber 17 and the second supply chamber 17. It is separated and supplied to section 13. As in the first embodiment, the next second gas is supplied after a considerable amount of the second gas supplied in the first gas has reacted, so the concentration of the second gas in the reaction mixture gas is reduced at once. It is smoother than in the case where the catalyst layer is heated, and the maximum temperature of the catalyst layer is also suppressed to below the allowable temperature.

[Quantitative effect of invention]
The effects of this invention were calculated using the following method.

[Determination of reaction rate formula and temperature distribution calculation formula]
(Assumption of reaction mechanism for each reaction)
(Reaction mechanism assumed for CO ₂ + 3H ₂ → CH ₃ H + H ₂ O)

Here, the species enclosed in parentheses are each adsorbed species on the catalyst including estimated intermediate products, and (A) represents the vacancy of the catalyst active site. Although the intermediate products described here have not been confirmed by spectroscopic methods, they are consistent with common chemical theory and mass balance. In general, intermediate products can be envisaged within a consistent range. The validity of the assumption is judged by the degree of agreement between the calculated value from the rate equation derived from the reaction mechanism and the measured value. The matching degree of the mechanism assumed in the present invention will be shown later in FIG. 4. As will be described later, it is determined from this figure that it has the accuracy to be used in subsequent simulations.
P _i is the partial pressure of each component in the gas phase. These symbols are also common below.

(Mechanism assumed for CO ₂ + H ₂ → CO + H ₂ O reaction)

(Mechanism assumed for CO ₂ + 2H ₂ → CO + H ₂ O reaction)

(Conversion to speed formula)
Derivation of rate equations from these elementary reactions is commonly done.
Follow Watson's instructions. Each speed formula was as follows. f _i is the fugacity of each component, and in the case of high-pressure reactions, fugacity must be used instead of pressure. Conversion from P _i was performed using the Berthelot formula.
The meanings of the symbols are as follows.

[Procedure for calculating constants of speed formula]
(Items other than K ₃ , K _{4, and} c)
Among the constants in these equations, items other than K ₃ , K _{4, and} c are calculated by entering the experimental conditions described in Non-Patent Document 2 into the equations r ₁ to r ₃ above, and using the calculated reaction rate values and the measured data in the table. The value that minimizes the sum of the squares of the differences in values was chosen. Some of the data is shown below.

(Calculation procedure of K ₃ , K ₄ , c)
For K ₃ , K ₄ , and c, the values read from the graph (influence of the amount of H ₂ O or CO added on the relative value of reaction rate) described in Non-Patent Document 4 are substituted into the above _r1 equation, and the above relative The value that best matched the values was determined using the least squares method.

(values of constants)
The following values were obtained as various constants. As shown in FIG. 4, the degree of agreement between the value calculated from the formula and the result was sufficiently satisfactory, with a contribution rate of 0.9965.

(Calculation formula for catalyst layer temperature distribution)
The formula for dT/dL (T: temperature, L: reaction tube length) used to calculate the temperature distribution is shown below.
Here, V is the reaction gas flow rate, S is the reaction tube cross-sectional area, D is the reaction tube inner diameter, Cp is the average specific heat of the reaction gas, Tw is the heating medium temperature, and U is the overall heat transfer coefficient between the reaction tube and the shell side.

(Steps for calculating steady state material balance)
Using these equations, we estimated the steady-state mass balance of the reaction system for each reactor.
(1) The shape of the reactor was selected, and the overall heat transfer coefficient was estimated using the standard method of calculating the heat transfer coefficient of the packed bed and shell side. The heating medium was pressurized boiling water. The overall heat transfer coefficient value was 120.7 J/s/K under the reaction tube inlet conditions. This value was used for all sections in subsequent calculations.
(2) The target STY was determined, and the corresponding amount of makeup CO ₂ and H ₂ was determined. It was assumed that CO ₂ , CO, and H ₂ in the circulating gas were added together with makeup CO ₂ and H ₂ to obtain the inlet gas amount.
(3) Boiling water temperature (T _w ) and inlet gas temperature (T _g ) were determined.
(4) The reactor was divided into minute sections of 1 mm each, and the amount and temperature of each component at the outlet of the sections were estimated sequentially using the Runge-Kutta method.
(4) The remaining gas (CO ₂ , H ₂ , CO) after removing methanol and water from the outlet gas is recycled to the reactor and combined with the newly supplied CO ₂ and H ₂ to become the inlet gas. In steady state, the assumed composition and the calculated value must match. The value was calculated using the solver function included in the calculation software manufactured by Microsoft, which is a variable analysis tool.
(5) If the solver does not converge, it indicates that a steady state does not hold under the selected conditions. If convergence does not occur even after adjusting T _w and T _g , it is determined that the STY selected in (2) cannot be achieved.

(Comparison of calculated values of steady state mass balance and actual measured values in the literature)
It was confirmed from the simulation results of the actually measured STY value and the maximum temperature of the catalyst layer described in the same literature based on the reaction conditions described in Non-Patent Document 3 that the mass balance calculated by this method does not contradict the actually measured value.
Assuming that the inner diameter of the reaction tube is 24.2 mm and the reaction gas inlet temperature and heater temperature are 214°C, the simulation result under the above operating conditions shows that the methanol production (STY) is 0.357 MeOH-kg/ L/h, the maximum temperature of the catalyst layer was 239°C, which was in almost sufficient agreement with the measured value.

[Restrictions on maximum STY value corresponding to reactor type]
The superiority or inferiority of reactor types can be determined by the STY values that can be obtained under various constraint conditions. Actual reactions are carried out under conditions with sufficient margin to accommodate fluctuations in operating conditions, but in order to compare superiority or inferiority, conditions must be kept the same.
The conditions should be unified to the following values.
(1) The amount of recycled gas will be unified at SV4,000/hr. This was addressed by adding this condition when assuming the amounts of CO ₂ , H ₂ , and CO at the reactor inlet. If the amount of recycled gas is increased, STY will increase, but this will lead to an increase in equipment costs and energy costs. By setting the values to be the same, the effect can be clearly determined.
(2) The maximum temperature (T _m ) of the catalyst layer should be 250°C or less. Since increasing the temperature of the catalyst layer directly leads to an increase in STY, the same value must be used to determine the superiority or inferiority of the reactor structure. As mentioned above, the deterioration of the catalyst becomes significant at temperatures above 270°C, so T _m is set at below 250°C in consideration of fluctuations in operating conditions.
(3) The minimum value (M m ) of the molar ratio ( _M ) of the reaction gas is set to 4.5 or more at all locations in the catalyst layer.
The molar ratio (M) of the reaction gas is defined as (number of moles of newly supplied H ₂ + number of moles of recycled H ₂ )/(number of moles of newly supplied CO ₂ and CO + recycled CO ₂ and the number of moles of CO).
The catalyst used in this reaction must be in a reduced state in order to maintain its activity. For this purpose, M must be greater than or equal to a certain value (M _m ), but this value is not clear. The above M in the experiment in Non-Patent Document 2 was 4.2 to 10. Since it is sufficient to use the same value when comparing the superiority of reactors, M _m =4.5 was used. In a range where the equilibrium value is not a problem, the reaction rate increases as the value of M approaches the stoichiometric ratio (3 in this case), so STY increases even at the same T _m . For comparison, it is essential to keep the same value.
On the other hand, under the operating conditions, the reactor inlet gas temperature (T _g ) and the heat medium temperature outside the reaction tube (T _w ) can be adjusted, so that the range of constraints (1), (2), and (3) above can be adjusted. Select the value with the maximum STY within the range. In many cases, boiling water is used as a heating medium in order to recover the heat of reaction as thermal energy.Since the descriptions that follow will also be based on boiling water, the temperature of the heating medium is assumed to be _Tw through the reactor.

[STY maximum value calculation procedure]
(1) Make-up CO ₂ /CO, T _w , T _g are determined appropriately, steady state mass balance is determined as described above (steady state mass balance calculation procedure), T _m , M _m
Calculate. If establishment of a steady state is confirmed, STY is determined to be a value corresponding to makeup CO ₂ /CO.
(2) Make-up CO ₂ /CO is increased if the values of T _m and M _m are within the constraint value. At the same time, T _w is increased and T _m and M _m are calculated. The following steps are repeated until T _m exceeds the constraint value.
(3) Even if T _m exceeds the constraint value, if M _m still has a margin, Tw is lowered and Tg is increased.
(4) If both T _m and M _m almost reach the constraint values, that STY is set as the maximum value.

[Reactor of comparative example]
Figure 3 shows an existing reactor as a comparative example.
In the reactor 1 of this comparative example, similarly to the reactor of the second embodiment as shown in FIG. At the same time, a supply chamber 17 of the first supply section 12 and a discharge chamber 18 of the discharge section 15 are partitioned on both sides of the heat exchange chamber 16.
In addition, in the heat exchange chamber 16, a plurality of reaction tubes 8 containing the fixed catalyst bed 11 are arranged in parallel to form a reaction tube 8 group, and both ends of each reaction tube 8 are connected to a supply chamber 17 and a discharge chamber 18. Further, the heating/cooling section 14 has an inlet nozzle 14a through which pressurized boiling water is supplied as a heating medium to the lower part of the heat exchange chamber 16 surrounding the reaction tube 8, and An outlet 14b is provided for discharging a multiphase flow of boiling water and steam from the upper part of the exchange chamber 16, and the reaction heat can be recovered by a heat recovery section (not shown).
However, in the reactor 1 of this comparative example, unlike the reactor of the second embodiment shown in FIG. No supply header tube is provided for supplying the second gas to each feed tube.
The first gas and the second gas are supplied to the supply chamber 17 of the first supply unit 12 in a mixed gas state in which the first gas and the second gas are mixed in advance. The mixed gas is distributed and supplied to each reaction tube 8, and flows downstream while the reaction progresses in the fixed catalyst bed 11 in each reaction tube 8.
Although FIG. 3 shows a cross-sectional view of four reaction tubes, the number of reaction tubes is not limited and is determined by the required production amount. Since STY is the production amount per unit volume of catalyst, the comparison of STY is a comparison of the production amount per reaction tube and is unrelated to the number of tubes. Therefore, if the cross-sectional area and length of the reaction tube, the diameter of the catalyst particles used, etc. are the same between Examples and Comparative Examples, it becomes possible to compare the superiority of the reactor.
The inner diameter of the reaction tube was 24.2 mm, the length of the reaction tube was 4 m, and the catalyst particle diameter was 3 mmΦ.
The procedure for calculating the maximum STY value in the comparative example is the same as the procedure in step 3 above. In the case of the existing type, M is the minimum at the reactor inlet, so M _m was taken as the value at the inlet.

[Reactor of Example]
In the examples, a reactor of the second embodiment as shown in FIG. 2 was used.
A diagram of the reactor is shown in Figure 2. In Figure 2, the newly supplied CO ₂ (make-up CO ₂ ) is supplied in four parts from the upstream side of the feed pipe, but there are no restrictions on the number of divisions or the interval between them, and the situation depends on the calculation results of temperature distribution, etc. selected accordingly. In the illustrated example, makeup CO ₂ is supplied from positions 750 mm, 1500 mm, 2250 mm, and 3000 mm from the upstream side.
It is desirable that the amount supplied from the side holes at each position be as uniform as possible. The pressure in the catalyst layer is higher on the upstream side due to pressure loss, so if the side pore diameters are the same, there may be a difference in the supply amount between the upstream and downstream sides, and in extreme cases, the pore sizes on the upstream and downstream sides must be different. Sometimes it doesn't. This can be avoided by creating a large pressure loss in the side holes, that is, by reducing the hole diameter. In this example, the number of side holes was two per location, and the hole diameter was 0.1 mm. By adopting a small hole diameter, the calculated value of the ratio of the supply amount from the most upstream side hole to that from the most downstream side hole is approximately the same value of about 0.94. The necessity of this level of small pore size differs depending on the inner diameter of the reaction tube and catalyst particle size. Although not always necessary, this pore size is preferred in this case.
The feed tube was a SUS tube with an outer diameter of 6 mm and an inner diameter of 3.4 mm. This insertion reduced the amount of catalyst in the reaction tube, so the inner diameter of the reaction tube was increased, and the amount of catalyst and the flow rate of the reaction gas in the catalyst layer were kept the same as in the comparative example. The inner diameter of the reaction tube was 24.9 mm.

The procedure for calculating the STY maximum value is shown below.
(1) The composition of the circulating gas is assumed at the beginning, which is the same as the calculation for the existing type. At the inlet of the reactor, the circulating gas and the make-up _H2 are combined and enter the catalyst bed. CO ₂ and CO in the circulating gas react to produce CH ₃ OH, H ₂ O, and CO. CO ₂ in the reaction gas decreases. The temperature increases by a value corresponding to the difference between the heat of reaction and the heat removed by the boiling water. This aspect is calculated using the method described in 2 above.
(2) When the calculation reaches the point where the makeup CO ₂ /CO from the feed pipe is supplied, add up the CO ₂ /CO at that point and the makeup CO ₂ /CO amount and calculate the new composition of each component. Calculate. The composition changes and temperature changes are then calculated downstream to the next supply point.
(3) Perform calculations in the same way from the next supply point to the reactor outlet. The remainder after subtracting CH ₃ OH and H ₂ O produced from the outlet gas becomes the circulating gas. The amount of H ₂ , CO ₂ , CO that is circulated is calculated.
(4) Repeat the calculation until the amount of each component matches the initially assumed amount with sufficient accuracy.
(5) Find the maximum value of STY that can be obtained using the procedure described in 3 above.
(6) In the case of the example, since M is the minimum at the most downstream supply port, M _m was set to that value.

[Quantitative comparison of possible STYs of Examples and Comparative Examples]
(Example 1 and Comparative Example 1)
α in Table 1 is an index of catalytic activity, and is set to 1 when the current catalytic activity is present.
As shown in Table 3, with the catalyst activity at this time, the STY obtained in Comparative Example 1 was 0.528 kg-MeOH/L/h. On the other hand, the STY obtained in Example 1 is 0.607 kg-MeOH/L/h. That is, the STY when using this patent is 1.15 times that of the conventional method when using the current catalyst activity. In proportion to this, the energy cost required for the reaction system is reduced, and fixed costs such as equipment costs are also reduced.

(Example 2 and Comparative Example 2)
The effect of the present invention increases as the catalyst activity increases. The fact that more highly active catalysts will be put into practical use in the future is that Patent Document 6 has already proposed a catalyst with significantly higher activity than conventional products by adding new components to the same copper-based catalyst. Since it is fully expected from the above, a case where the activity is twice the current activity (ie, α=2) is also shown.
Comparative Example 2 is a case of an existing shell-and-tube heat exchanger type reactor, and Example 2 is a case of a reactor of the type of the present invention. Each is the maximum value that can be taken under the above-mentioned constraint conditions (T _m ≦250, M _m ≧4.5).
As shown in Table 3, in the future, when the catalytic activity is increased to, for example, twice the current level, this effect will further increase to 1.19 times, and the industrial usefulness will further increase.

[Conditions for obtaining the same STY as in Examples 1 and 2 using an existing reactor]
Even with existing reactors, if the reaction conditions are made harsher, a large STY can be obtained instantaneously. However, as shown in Reference Examples 1 and 2 in Table 3, under the reaction conditions to obtain the same STY as in Examples 1 and 2, the maximum temperature greatly exceeds the allowable temperature, so it only lasts for a short period of time due to catalyst deterioration, and is not commercially viable. do not. It is clear that the reactor of this type contributes to a reduction in peak temperature and thus to a substantial increase in STY. In order to further clarify the effect of this method, the calculated temperature distribution values for Reference Example 2 (dotted line) and Example 2 (solid line) are shown in FIG.

Industrial applications

Regarding the effective use of _CO2 , the main route at present is to convert it into MeOH through reaction. In addition to its conventional uses, MeOH is being used increasingly for fuels and plastics. As the recovery of CO ₂ that was conventionally emitted into the atmosphere progresses, the supply of CO ₂ will inevitably increase, and the field of application of the present invention will expand. Furthermore, the present invention has great utility not only in the production of methanol, but also in all cases where exothermic reactions are carried out by solid catalyst reactions in fixed beds. Specifically, the fields include hydrogenation reactions and oxidation reactions using catalysts. The present invention relaxes the maximum temperature restriction.
In addition to restrictions on allowable temperature in some oxidation reactions, there are also cases where it is not possible to increase the supply amount of one of the reactants in order to avoid an explosive range of reaction gas composition. The present invention is also effective in such cases, and even if the supply amount is increased, the present invention allows for distributed supply, so that the composition inside the reactor can be maintained below the lower explosive limit. These uses are also hidden.

1 Reactor 2 First gas supply line 3 Second gas supply line 4 Reaction mixed gas outlet gas line 5 Separation section 6 Circulation section 6a Circulation line 7 Reaction component extraction line 8 Reaction tube 9 Feed tube side hole (provided in the feed tube) (many holes)
10 Feed pipe 11 Catalyst fixed bed 11a Catalyst bed tray 11b Catalyst bed holder 12 First supply section 13 Second supply section 13a Second gas supply header pipe
14 Heating and cooling section 14a Heat medium inlet 14b Heat medium outlet 15 Reaction mixed gas discharge section 16 Heat exchange chamber
16a Heat exchange chamber tube sheet
17 First supply section supply chamber
18 Reaction mixed gas discharge chamber

Claims (6)

A reactor for reacting a first gas and a second gas that cause an exothermic reaction in the presence of a catalyst, the reactor comprising:
a reaction tube,
a feed tube disposed within the reaction tube and provided with a large number of dispersed holes;
a fixed catalyst bed provided with the catalyst disposed between the reaction tube and the feed tube;
a first supply section that supplies a first gas between the reaction tube and the feed tube from one end side;
a second supply section that supplies the second gas to the feed pipe;
a heating and cooling section that heats or cools the reaction tube;
a discharge part for discharging the reaction mixture gas between the reaction tube and the feed tube from the other end side,
A fixed bed catalyst, wherein the first gas is caused to flow in one direction through the catalyst fixed bed, and the second gas is supplied from the feed pipe to positions dispersed in the one direction with respect to the catalyst fixed bed for reaction. reactor. A heat exchange chamber of the heating and cooling section is defined within the reactor shell, and a supply chamber of the first supply section and a discharge chamber of the discharge section are defined on both sides of the heat exchange chamber,
A plurality of reaction tubes containing the feed tube and the fixed catalyst bed are arranged in the heat exchange chamber, and both ends of the plurality of reaction tubes are in ventilated communication with the supply chamber and the discharge chamber. , a fixed bed catalytic reactor according to claim 1. 2. The heating and cooling section includes a heat recovery section that supplies pressurized boiling water to the heat exchange chamber, discharges a multiphase flow of heated boiling water and steam from the heat exchange chamber, and recovers reaction heat. Fixed bed catalytic reactor described in . The fixed bed catalytic reactor according to any one of claims 1 to 3, a separation section that separates a reaction product component from the reaction mixture gas, and a circulating gas from which the reaction product component has been separated from the reaction mixture gas. A plant comprising: a circulation section that circulates to a first supply section or a second supply section. The first gas is a hydrogen-containing gas, the circulating gas, or the circulating gas to which hydrogen is added, the second gas is a carbon dioxide-containing gas or a carbon dioxide-containing gas to which hydrogen is added, and the reaction product component is methanol. The plant according to claim 4. A reaction method in which a first gas and a second gas that cause an exothermic reaction in the presence of a catalyst are supplied while being cooled to a fixed catalyst bed having the catalyst to react,
A feed pipe provided with a large number of dispersed holes is arranged along the catalyst fixed layer,
Fixing the catalyst, while causing the first raw material gas to flow in the fixed catalyst bed in one direction, supplying the second gas from the feed pipe to positions dispersed in the one direction with respect to the fixed catalyst bed for reaction. Reaction method using layers.