KR20220007440A - Multi-layered tubular reactor for methanol synthesis from synthesis gas - Google Patents

Abstract

The present invention is a tube-type reactor with a multi-layer structure, comprising: an inner tube of different diameters disposed to have the center of a cavity in turn from the inside; a catalyst tube; and a shell, and specifically, to a reactor for methanol production from syngas, which contains a cooling medium inside the shell and has a source gas injection unit at a lower part of the inner tube, wherein the catalyst tube is filled with a catalyst for methanol synthesis, and an upper part of the inner tube and an upper part of the catalyst tube are connected to each other.

Description

Multi-layered tubular reactor for methanol synthesis from synthesis gas

The present invention is a tubular reactor of a multi-layer structure, comprising an inner tube, a catalyst tube and a shell of different diameters arranged to have a center of a cavity in turn from the inside, specifically, the inside of the shell contains a cooling medium, It relates to a reactor for producing methanol from synthesis gas (syngas), in which a source gas injection unit is provided at a lower portion, the catalyst tube is filled with a catalyst for methanol synthesis, and the upper portion of the inner tube and the upper portion of the catalyst tube are connected to each other.

The catalytic conversion of syngas (a mixture of CO and H ₂ ) to hydrocarbons and alcohols is one of the most important processes for improving the carbon cycle due to increasing energy demand and environmental concerns. Methanol by itself is an excellent fuel material and is a versatile material that can be used directly in fuel cells for electricity production and as a gasoline substitute. Furthermore, it can be upgraded to a chemical building block for ethylene and propylene by dimethyl ether (DME) or methanol-to-olefin processes, which are alternatives to diesel. Recently, methanol has been commercially synthesized on a _{Cu/ZnO/Al 2} O ₃ catalyst under low pressure (5-10 MPa) and low temperature (473.15-57.15K) conditions. Methanol can be formed through a series of processes including highly exothermic CO hydrogenation reaction, CO ₂ hydrogenation reaction, and water-gas shift (WGS) reaction. The methanol synthesis reaction is exothermic and reversible. At low temperatures, the thermodynamic equilibrium conversion rate is high, but the reaction rate is slow and a large amount of catalyst is required. On the other hand, at a high temperature, the reaction rate is increased, but the thermodynamic equilibrium conversion is low, so that the conversion of CO and CO ₂ is reduced and the thermal stability of the catalyst is deteriorated. Therefore, it is one of the most important issues in the methanol synthesis reactor configuration to form a high temperature at the reactor inlet for a high reaction rate (<270° C.) and then gradually decrease the temperature toward the outlet to increase the thermodynamic equilibrium shift.

Existing methods for methanol production are based on the direct combination of _{CO, CO 2} , and H _{2 gases in a catalyst packed bed reactor.} For example, a conventional methanol reactor consists of a tubular heat exchanger having a shell and a plurality of tubes filled with catalyst pellets. Saturated water circulates through the shell side of the reactor, removing heat from the reaction. Due to the equilibrium behavior of the reaction, methanol conversion in the conventional reactor is kept low, and most of the unreacted synthesis gas is reused in the process. In order to overcome the limitations of the conventional reactor, various improved reactors have been proposed. Velardi and Barresi attempted to improve the performance of exothermic and equilibrium-limited synthesis reactions by combining three catalytic fixed bed reactors, and to evaluate the effects of driving variables such as conversion time, injection rate and inlet temperature on productivity. Rahimpour and Alizadehhesari developed a membrane dual type methanol reactor system, compared its performance with a conventional dual type methanol synthesis reactor, and demonstrated that the developed reactor induces a favorable temperature profile for long-term catalyst life. Samimi et al. applied the water selective permeation membrane to the reverse water gas shift reactor installed in front of the methanol synthesis reactor, and proved that an additional water separator was not required in the process and that the methanol production rate was significantly increased. Rahmatmand et al. introduced a new methanol synthesis process form in which an adiabatic and plate water-cooled reactor replaces the conventional two-stage reactor system, and supported the effectiveness of the process by improving catalyst durability through mathematical modeling.

In the present invention, a compact reactor, which is a modification of the bayonet reactor, is proposed, and its performance is compared with the conventional reactor in terms of temperature control and productivity. Computational fluid dynamics (CFD) simulation was applied for a more complex depiction of the proposed reactor compared to the conventional reactor, and the detailed distribution of conversion rate and temperature was searched. It is worth noting from this that when the length to diameter ratio is high, the average radial temperature profile is not significantly different from that based on the radial temperature gradient. However, as the size of the tube increases, the temperature within the tube, particularly near the center, may increase due to the limited rate of heat transfer in the radial direction due to heat transfer resistance. In addition, since the scale-up of the compact reactor introduced in the present invention is not based on a multi-tubular configuration, but rather on an increase in tube size, an excessive increase in tube radius may cause thermal instability. Therefore, it is important to check the radial gradient of the catalyst tube. Finally, the model developed according to the present invention can be used to evaluate the effect of driving conditions and design variables on the performance of the structure.

The present invention is a tubular reactor of multi-layer structure, comprising inner tubes of different diameters, catalyst tubes and shells arranged to have a center of a cavity in turn from the inside, the shell containing a cooling medium inside, directly or indirectly in the inner tube is provided with a source gas injection unit connected by A reactor for producing methanol from phosphorus, syngas is provided.

In the reactor of the present invention, in the process of producing methanol from synthesis gas using a catalytic reaction, synthesis gas, which is a raw material supplied for an efficient reaction, requires preheating, and considering that the catalytic reaction is an exothermic reaction, the catalyst By designing the structure of the reactor so that the raw material can be preheated using the thermal energy generated from the reaction, energy consumption for preheating is prevented and an economical process can be implemented.

In this case, the reactor is a tubular reactor of a triple layer structure provided at the bottom of the reactor and having a source gas injection part directly connected to the inner tube, or a diameter smaller than the inner diameter of the inner tube arranged to have the center of the cavity at the center of the inner tube. It may be a tubular reactor of a quadruple layer structure additionally provided with a source gas injection tube having The lower portion is open and located higher than the closed lower end of the inner tube, through which the source gas can be delivered to the lower end of the inner tube. In particular, in the tubular reactor having a quadruple layer structure, the raw material gas is introduced from the upper part of the reactor, and the product formed through the catalytic reaction can be discharged through the lower part of the reactor, so the composition of the raw material gas inlet manifold is easy. It may be more advantageous in manufacturing.

Hereinafter, the present invention will be described in detail.

The present invention comprehensively devises a reactor for synthesizing methanol from synthesis gas in which the temperature of the reactor is increased by an exothermic reaction. It is intended to improve the occurrence of a significant temperature gradient near the outer wall and thus a non-uniform reaction. Accordingly, the present inventors have devised a triple-layer or more multi-layered tubular reactor by additionally providing an inner tube through which the source gas can pass inside the tube, and supplying the source gas to the catalyst tube through the inner tube without preheating the catalyst tube. By preheating the raw material gas by the reaction heat transferred from the catalyst tube, the temperature gradient inside and outside the tube was alleviated to achieve the effect of uniformly reacting throughout the reactor.

The overall reaction route of methanol synthesis according to the present invention is as follows:

CO hydrogenation: CO + 2H ₂ ↔ CH ₃ OH (1)

CO ₂ hydrogenation: CO ₂ + 3H ₂ ↔ CH ₃ OH + H ₂ O (2)

Reverse water-gas shift: CO ₂ + H ₂ ↔ CO + H ₂ O (3)

Among the three reactions performed in the reactor of the present invention, the reverse water gas shift reaction is a mild endothermic reaction, but since the CO hydrogenation and CO ₂ hydrogenation reactions are exothermic reactions, some of the heat generated by the reverse water gas shift reaction is offset. However, it still generates a significant amount of heat according to the progress of the reaction.

Accordingly, in the reactor of the present invention, the raw material gas injected directly into the lower end of the inner tube or through the raw material gas injection tube may be preheated by heat generated from the reaction in the catalyst tube while moving upward. Accordingly, the reactor of the present invention does not require a separate device for preheating the raw material gas, thereby simplifying the device and reducing energy consumption.

Specifically, in the reactor of the present invention, the fuel gas preheated while passing through the inner tube as described above passes through the catalyst tube as described above for the CO hydrogenation reaction, CO ₂ hydrogenation reaction, and reverse water-gas shift (RWGS). ) reaction can be carried out. As described above, all of the above reactions except reverse water gas conversion are exothermic reactions, and since the reaction is an equilibrium reaction, the temperature of the reactor may be increased by heat generated as the reaction proceeds, which inhibits the reaction or rather induces a reverse reaction. can be a factor. Therefore, in order to improve methanol productivity, it is necessary to efficiently dissipate heat generated from the reaction. On the other hand, in the reactor of the present invention, the catalyst tube is in contact with the inner tube to which the raw material gas is supplied through the inner surface and is in contact with the cooling medium supported on the outermost shell through the outer surface. The raw material gas can be consumed to vaporize the preheating and/or cooling medium to maintain or lower the temperature of the reactor so that the equilibrium reaction can be continued.

Meanwhile, the reactor of the present invention may be a cylindrical reactor having a multi-layer structure having in one shell at least one subset consisting of an inner tube, a catalyst tube, and optionally a source gas injection tube located at the center of the inner tube. By including a plurality of catalyst tubes in one shell as described above, it is possible to prevent an excessive increase in the volume of the entire reactor while enabling mass production. Specific configuration examples are shown in FIGS. 2 and 4 .

For example, the methanol synthesis catalyst charged in the reactor for producing methanol of the present invention may be a copper/zinc/aluminum-based catalyst, a chromium oxide-based catalyst, a copper-based catalyst, or a combination thereof. For example, as the methanol synthesis catalyst, a Cu/ZnO/Al ₂ O ₃ catalyst widely used in the art may be used. Specifically, a catalyst of 60 wt% Cu/30 wt% ZnO/10 wt% Al ₂ O ₃ may be used, but is not limited thereto.

Specifically, the methanol synthesis reaction may be performed under conditions of a reaction temperature of 240 to 270° C., a reaction pressure of 40 to 100 bar, and a space velocity of 1,000 to 50,000/h, but is not limited thereto. As described above, since the reaction for producing methanol is an exothermic reaction, the heat of reaction can be controlled in order to increase the yield of methanol. Specifically, the methanol synthesis reaction is an equilibrium reaction, and as the temperature of the reactor increases due to heat generated as the reaction proceeds, the equilibrium conversion rate of the reactants may decrease, so it is necessary to control the temperature in the reactor. For example, for this purpose, the reaction heat may be controlled using latent heat due to vaporization of water by flowing water as a refrigerant to the outside of the tube in which the catalyst layer of the methanol synthesis reactor is located, but is not limited thereto.

As described above, the inside of the shell may include a cooling medium in order to control the temperature of the reactor, and the cooling medium may absorb heat generated from the reaction in the catalyst tube and vaporize it.

Specifically, the shell discharges the vaporized cooling medium through the gas outlet to the outside through a closed cooling loop system formed in connection with a water tank provided externally through a gas outlet formed at the upper portion and a liquid inlet formed at the lower end to the outside and cooled again The liquefied medium may be stored in a water tank and then supplied into the shell through the liquid inlet, but is not limited thereto.

Meanwhile, in determining the size of the reactor of the present invention, the catalyst tube may be designed to have a diameter of 1.5 to 3 times that of the inner tube. If the diameter of the catalyst tube is less than 1.5 times the diameter of the inner tube, it may be advantageous to control the temperature of the catalyst bed, but the volume of the catalyst bed is relatively small, so there is a problem in that the scale of the entire reactor increases in order to perform the reaction of the desired capacity, 3 If it exceeds twice the volume of the catalyst layer is too large, it is difficult to remove the heat of reaction, and the reaction efficiency may decrease.

At this time, it may be advantageous to adjust the diameter of the inner tube to 30 to 70 mm. When the diameter of the inner tube is less than 30 mm, the volume of the inner tube is small and the residence time of the raw material gas is shortened. . On the other hand, when the diameter of the inner tube exceeds 70 mm, the volume of the inner tube becomes excessively large and the residence time of the raw material gas becomes longer. . However, the diameter of the inner tube and the effect according to it specifically presented above exemplify the effect in the reactor of the triple layer structure, and for the reactor of the quadruple layer structure having an additional source gas injection tube in the center of the inner tube, the Absolute figures of scale may vary.

The reactor of the present invention includes a tube capable of preheating while supplying a raw material gas capable of absorbing heat generated from the reactor inside the reactor for performing an exothermic reaction for synthesizing methanol from synthesis gas to control the temperature of the reactor And it is possible to perform a uniform reaction by lowering the temperature gradient in the radial direction, thereby simplifying the apparatus, reducing energy consumption, and improving methanol productivity.

1 is a diagram schematically showing a mini-pilot scale experimental system in which a single tube (inner diameter = 0.038 m (1.5 in), length = 1.5 m) with a shell is used as a reactor for methanol 10 kg/day production. . As experimental conditions, pressure, SV, and temperature (both inlet and wall) were specified as 5 MPa, 6771 mL/(g _cat h), and 210°C, respectively, and the raw material composition was H ₂ : CO : CO ₂ : N ₂ = 58.4 : 6.0 : 9.1 : 26.5 mol%.
Figure 2 shows (a) a conventional multi-tubular fixed-bed reactor (the inner diameter and the packing depth of the tube) for the production of 50 kg/day of methanol, which is 5 times the capacity of the mini-pilot system. 0.038 m (1.5 in) and 2.2 m, respectively, and the inner diameter of the shell is 0.1 m) and (b) a compact reactor according to an embodiment of the present invention (inner, middle and outermost of the tube) ) The radius is 0.035, 0.077, and 0.085 m, respectively, and the length of the central tube is 2.2 m) schematically.
3 is a diagram showing a comparison between simulation results and experimental data for (a) conversion and (b) reactor catalyst bed temperature according to reactor length. As experimental conditions, pressure, SV, and temperature (both inlet and wall) were specified as 5 MPa, 6771 mL/(g _cat h), and 210°C, respectively, and the raw material composition was H ₂ : CO : CO ₂ : N ₂ = 58.4 : 6.0 : 9.1 : 26.5 mol%.
4 is a view showing (a) CO conversion rate and (b) temperature profile in the tube along the catalyst filling axis in a conventional multi-tubular fixed bed reactor. A cross-sectional view of the radial distribution is shown at z (packing depth) = 0.6 m. (c) is a diagram showing the radiation temperature gradient at z = 0.3 (left), 0.6 (center) and 0.9 (right).
5 is a diagram showing (a) CO conversion and (b) temperature profile in a compact reactor proposed according to an embodiment of the present invention. A cross-sectional view of the radial distribution is shown at z (packing depth) = 1.1 m.
6 is under different raw material temperature of the compact reactor and in the T _inlet = 150 ℃ conditions in the conventional reactor (a) CO conversion rate, (b) CO ₂ conversion, (c) H ₂ conversion, (d) production, ( e) is a diagram showing the effect of space velocity on inlet temperature, and (f) maximum temperature. SV _ref was assigned to 5000 mL/(g _cat ·h). The designations "Conventional" and "Comp" in the drawings indicate the conventional reactor and the compact reactor, respectively, and the numbers indicate the raw material temperature in the compact reactor. The red dotted line corresponds to the conversion rate of the conventional reactor under reference conditions.
7 shows (a) conversion and productivity, (b) productivity per CO conversion, and maximum and inlet temperature, (c) CO conversion under the conditions of T _feed = 30° C. and SV = 5000 mL/(g _{cat h).} A diagram showing the effect of diameter on axial profile, (d) axial profile of temperature, (e) radiation profile of temperature at z (packing depth) = 1.8 m, and (f) temperature versus conversion profile. The designations "Conventional" and "Comp" in the drawings indicate the conventional reactor and the compact reactor, respectively, and the number indicates the D/D _ref ratio _{when D ref = 0.035 m.}
8 is a side view and cross-sectional structure of (a) a triple-bed reactor and (b) a quadruple-bed reactor for the synthesis of methanol from synthesis gas according to an embodiment of the present invention, and the movement path of the raw material gas and the cooling medium when the reactor is driven It is a schematic diagram.

Hereinafter, the present invention will be described in more detail by way of Examples. However, the following examples are only for illustrating the present invention, and the scope of the present invention is not limited thereto.

production example 1: Preparation of catalyst

A commercial catalyst, Cu/ZnO/Al ₂ O ₃ (Dalian Reak Science, RK-05), was used in a pilot-scale methanol synthesis experiment. The methanol synthesis catalyst has a molar ratio of Cu/Zn/Al = 60/30/10, and is cylindrical pellets (5 mm diameter and 5 mm high) having a bulk density of 1.1 to 1.3 kg/L. In order to prevent sudden temperature rise, methanol synthesis catalyst (0.77 kg) was physically mixed with 1.60 kg of α-alumina balls (5 mm diameter) before being charged into the reactor tube. The mixed catalyst was charged into a tubular reactor. Prior to methanol synthesis, the charged catalyst was reduced at 250° C. for 5 hours under a flow of 5% H _{2 balanced with N 2 .}

Example 1: Experimental setup

1 schematically shows a pilot-scale methanol synthesis system according to the present invention. The reactor basically has a structure of a vertical shell and a tube heat exchanger, a methanol synthesis catalyst is filled inside the vertical tube, and saturated water is present around the vertical tube. The heat generated in the exothermic reaction is transferred to saturated water to produce steam. The heated water and steam are then cooled through a heat exchanger and transferred to a water vessel. The reaction temperature can be easily controlled by adjusting the system pressure on the shell side in a closed cooling loop system. A pilot scale methanol synthesis experiment was performed in a loop reactor equipped with product separation and internal recycling (see FIG. 1 ). Model synthesis gas (H ₂ : CO : CO ₂ : N ₂ = 58.4 : 6.0 : 9.1 : 26.5 mol%) was mixed in advance and injected into the reactor. The inlet gas was preheated and converted in a fixed bed reactor loaded with a commercial methanol synthesis catalyst. The obtained mixture product (resulting product mixture, unreacted gas, methanol and water) was cooled, the liquid product was first separated under a pressure of 50 bar, and then transferred to a methanol storage tank near atmospheric pressure. The liquid phase was recovered with crude methanol containing methanol and water. The separated gas was recycled by a recycle compressor, mixed with the reaction raw material gas, and then introduced into the reactor. The recycle gas was partially purged to prevent accumulation of traces of inert gases (such as CH ₄ and N _{2 ).}

Example 2: Kinetics modelling (kinetic modeling)

The overall reaction route of methanol synthesis according to the present invention is as follows:

CO hydrogenation: CO + 2H ₂ ↔ CH ₃ OH (1)

CO ₂ hydrogenation: CO ₂ + 3H ₂ ↔ CH ₃ OH + H ₂ O (2)

Reverse water-gas shift: CO ₂ + H ₂ ↔ CO + H ₂ O (3)

The reaction rate can be calculated as:

(4)

(4)

(5)

(5)

(6)

(6)

where f is, fugacity coefficient, Introduction to Chemical Engineering Thermodynamics, 7th ed., McGraw-Hill, New York, 2005), calculated using generalized correlations Shows fugacity, and K _p represents the reaction equilibrium constant.

As shown in Table 1 above, except for the reaction equilibrium constant, the reported kinetic parameters obtained by data optimization in the process simulator (UniSim Design Suite, Honeywell Inc.) as follows were modified without modification. Used:

(7)

(7)

(8)

(8)

(9)

(9)

In this case, T is Kelvin, K _P,A and K _P,C are all MPa ^{- 2} units, and K _P,B is dimensionless. The validity of the reaction rate equation was verified using experimental data in a mini-pilot reaction system ( FIG. 1 ).

Example 3: CFDs modelling

CFD modeling of the reactor was performed using COMSOL Multiphysics 5.3 (COMSOL, Inc.), and the configuration of the two reactors is shown in FIG. 2 . Both reactors have a capacity of 50 kg/day for product (5 times greater than the mini-pilot in FIG. 1 ). A single-tube reactor was used for the experimental study for the verification of the kinetic model (see Example 1), whereas the multi-tube reactor was considered for the simulation study because it is the most widely used form for scale-up in industrial applications. However, since the limited capacity of the cooling shell in conventional reactors can cause thermal instability, a different configuration of compact reactors was introduced and its performance was improved in the conventional reactors with catalyst bed volumes remaining the same in both reactor configurations. compared with the reactor of The existing reactor (Fig. 2a) is a multi-tube shell and tube reactor, and the total number of tubes is set to three. The reaction was carried out on the side of the tube filled with the catalyst, and the cooling water on the shell side controlled the reaction temperature. The compact reactor ( FIG. 2b ) is a variant of the bayonet-type reactor consisting of three tubes. In the outermost tube, cooling water flows to control the reaction temperature, while the catalyst is charged in the inner tube. The raw material gas flows upward after being injected into the bottom of the inner tube to absorb heat generated from the catalyst bed, and enters the catalyst bed and flows downward. In this case, the raw material gas is directly injected into the inner tube from the bottom of the reactor in the case of a three-bed reactor, or in the case of a four-bed reactor, the raw material gas injected through the raw material injection unit located at the upper part of the reactor is at the center of the inner tube connected to the raw material injection unit It is indirectly transmitted to the lower end of the inner tube through the additionally provided source gas injection tube (see FIGS. 8A and 8B ). The temperature of the catalyst bed was simultaneously controlled by preheating the cooling raw material gas in the inner tube and cooling by the cooling medium (coolant media, in the present invention, cooling water) flowing in the outermost tube.

The following built-in modules from COMSOL Multiphysics were used to account for the mass, energy, and momentum balances of the reactant flow channels, respectively: "Transport of Concentrated Species" "Heat Transfer in Fluids" (Heat Transfer in Fluids)", and "Free and Porous Media Flow". A reactor body made of stainless steel (SUS 316L) was analyzed using the "Heat Transfer in Solids" module, which assumed the presence of heat conduction, ie, excluded heat transfer by radiation. Detailed fixed equations for each module are shown in Table 2, and symbols are shown in Table 3.

To calculate the profile of CO conversion and temperature, Galerkin's method was applied (The mathematical theory of finite element methods, 2nd ed., Springer-Verlag, New York, 2002), free tetrahedral grids ) was defined. A parallel direct sparse solver interface (PARDISO) that applies parallel computing methods on shared-memory and distributed-memory multiprocessors Using (Future Gener. Comp. Sy. 20 (2004) 475-487), a balance equation was derived and used to improve the calculation speed.

Since the latent heat of saturated water is used for cooling, the cooling water shell is assumed to be in isothermal condition without consideration of balance, while the momentum and energy balance equations for the syngas feed contributing to cooling of the catalyst bed are in the compact reactor It is worth noting that it is resolved within the inner tube of In addition, it was assumed that the boiling phenomenon that may occur in the cooling water shell is negligible for the sake of simplicity.

(N: 실험적 데이터의 수)로 정의되는 절대 상대 잔차 평균(mean of absolute relative residuals; MARR), 및

로 정의되는 개별 오차의 상대 표준 편차(relative standard deviation of individual errors; RSDE)와 같은 통계를 따르는 실험적 데이터와 시뮬레이션된 결과 간의 전환율 및 온도 프로파일을 비교하여 나타낸다. 반응기 모델에서 다른 디자인 변수는 촉매 베드와 냉각수 쉘 사이의 전체 열전달 계수(overall heat transfer coefficient)였으며, 이 값은 균형잡힌 방식에서 전환율과 온도 프로파일의 오차를 최소화함으로써 70 W/(m²·K)로 결정되었다. 이는, 도 3에 나타난 바와 같이, CO 전환율과 CO₂/H₂ 전환율 사이에 균형(trade-off)이 존재함을 나타내는 것으로, 달리 말하면, CO 전환율의 오차가 증가함으로써 CO₂ 및 H₂ 전환율에 대한 오차가 감소할 수 있음을 의미한다. CO 전환율의 정도가 다른 변수에 비해 높으므로, CO 전환율이 보다 주목받고 있다. 따라서, CO 전환율의 상대 오차는 다른 변수에 비해 훨씬 더 낮게 결정되었다. 또한, 상대 오차가 목적 요소(objective elements)에 따라 상이함에도 불구하고, 실험적 데이터로부터 시뮬레이션된 값들의 절대 편차로 정의되는 절대 오차(%P, percent point)는 유사하였다: 예컨대, CO, CO₂ 및 H₂ 전환율의 절대 오차는 각각 4.2, 5.4, 및 4.9 %P임. 이는 시뮬레이션된 결과가 측정된 데이터와 일치함을 나타내며, 개발된 모델의 유효성을 입증하는 것이다.3 is

mean of absolute relative residuals (MARR), defined as (N: number of experimental data), and

The conversion rate and temperature profile between the simulated results and the experimental data following statistics such as the relative standard deviation of individual errors (RSDE), defined as Another design variable in the reactor model was the overall heat transfer coefficient between the catalyst bed and the cooling water shell, which was 70 W/(m ² K) by minimizing the error in the conversion rate and temperature profile in a balanced manner. was decided with _{This indicates that a trade-off exists between the CO conversion rate and the CO 2} /H ₂ conversion rate, as shown in FIG. 3 . In other words, the error in the CO conversion rate increases, thereby affecting the CO ₂ and H ₂ conversion rates. This means that the error can be reduced. Since the degree of CO conversion rate is higher than that of other variables, the CO conversion rate is receiving more attention. Therefore, the relative error of the CO conversion rate was determined to be much lower compared to the other variables. In addition, although the relative error differs according to the objective elements, the absolute error (%P, percent point), which is defined as the absolute deviation of the simulated values from the experimental data, was similar: for example, CO, CO ₂ and The absolute errors of H ₂ conversion are 4.2, 5.4, and 4.9 %P, respectively. This indicates that the simulated results are consistent with the measured data and proves the validity of the developed model.

4 shows the CO conversion rate and temperature profile in the tube of a conventional multi-tubular fixed bed reactor. Simulation conditions were determined such that the production rate of methanol was 50 kg/day: pressure and GHSV were specified as 5 MPa and 5000 mL/(gcat·h), respectively, and inlet and wall temperatures were 150°C and 240°C, respectively. was set to °C. The raw material composition was set to H ₂ : CO : CO ₂ : N ₂ = 66.7 : 19.6 : 9.1 : 4.6 mol%. The physical properties specified in the present invention were provided from the process simulator, and are shown in Table 4 below.

Symbol Unit Value μ Pa s 1.613×10 ^-2 ε _P - 1.397 κ _br m ² 1.000×10 ^-12 κ W/(m·K) 0.1205 C _P J/(kg·K) 2.525

Conversion increased along the reactor and reached 43.2% (average, maximum conversion of 51.3%) at the reactor outlet. As shown in the cross-section at z = 0.6 m, the conversion rate at the center was lower than near the wall due to the high linear velocity at the center.

The temperature profile along the center of the reactor axis increased from the reactor inlet, reached a maximum (305.5° C.) and then decreased to 288.1° C. at the outlet. Detailed CFD results showed different temperature gradient patterns in the radial direction at different locations along the reactor axis. Near the inlet of the tube, the source gas close to the wall was heated by heat transfer with the medium in the shell, but the heat was not sufficiently transferred to the center of the tube, resulting in a lower temperature at the center (as shown in the left diagram in Fig. 4c) convex down to represent a parabola). At z (packing depth) = 0.6 m, due to the increase in the conversion rate from the center to the wall in the radial direction, the reaction rate also increases despite the decrease in the gas flow rate (the rate is determined by the product of the conversion rate and the flow rate), from the center induced an increase in temperature. Thereafter, the temperature decreased due to heat transfer with the shell (showing a binodal curve as shown in the central plot in Fig. 4c). At the second half of the reactor axis, the heat generated accumulated in the tube and the temperature at the center reached a maximum, while the temperature near the wall approached the wall temperature (showing a convex up parabola as shown in the plot on the right in Figure 4c). .

As described above, the conventional tubular reactor has a problem in that heat is removed around the tube center due to limited heat transfer capacity. In order to increase the heat transfer capacity, a compact reactor, a variant of the bayonet reactor, has been proposed, in which cooling is performed by a cooling medium in the shell as well as a cooling source gas in an inner tube ( 2b). As shown in FIG. 5, CFD results are shown in the catalyst packing tube (middle tube) under the same simulation conditions as the conventional reactor except for the raw material gas temperature of 30° C. (cf. The inlet temperature is the catalyst packing tube represents the temperature at the inlet, while the raw gas temperature represents the temperature at the inlet of the inner tube).

Regarding the conversion, the average value increased along the axis of the catalyst bed (the average conversion at the outlet was 48.4%, and the maximum value of the local conversion was 49.8%). Radial gradient (radial gradient) was lower than that near the outermost layer and due to the temperature difference between the inner tube r = r _i, the outermost tube of the conversion rate in the _{middle (outermost tube, r o,} middle). This characteristic was maintained throughout the entire catalyst bed. It is worth noting that the radial gradient in the compact reactor is more predominantly influenced by the temperature of the inner and outermost tubes than the difference in the linear velocity in the radial direction, whereas in conventional tubes it is affected by the linear velocity distribution. .

The temperature of the feed gas increased from 30° C. to about 130° C. (close to the inlet temperature in conventional reactors) due to heat transfer from the catalyst bed. At the beginning of the bed, the pattern of the radial temperature gradient is similar to that in the conversion gradient, whereas at the center of the bed (z = 1.1 m), the accumulated heat is removed by both the coolant in the outermost tube and the source gas in the inner tube. , leading to the maximum temperature between r _i,middle and ro _,middle. With the introduction of cooling feed gas, the higher heat transfer capacity of the compact reactor compared to the conventional reactor is attributed to the larger heat transfer area of the compact reactor (13.1 m ² vs. 7.72 m ² ). As a result, the maximum temperature was reduced to 284 °C compared to 305.5 °C in the conventional reactor. Collectively, since the reaction is in the thermodynamic regime rather than the kinetic regime, the lower maximum temperature of the compact reactor results in a higher conversion and greater than conventional reactors for the same space velocity. Methanol production rate was shown.

Since the heat transfer capacity of the reactor according to the present invention is larger than that of the conventional reactor, the space velocity was increased to evaluate the effect of increased heat production on conversion and the temperature profile. As shown in FIG. 6a , the CO conversion rate of the compact reactor decreased more significantly than that of the conventional reactor (gray bars labeled conventional) with increasing SV (sky blue bars labeled Comp-30). On the other hand, the CO ₂ conversion rate of the compact reactor increased with increasing SV, whereas the conventional reactor did not show a significant change in the _{CO 2 conversion rate (FIG. 6b).} Since _{the amount of CO 2} was smaller than that of CO, _{the effect of SV on H 2} conversion was similar to the effect on CO conversion ( FIG. 6c ). This characteristic can be attributed to the increased flow rate (flow rate) of the feed gas in the compact reactor, which can lead to more absorption of the heat generated in the catalyst bed, thereby reducing the inlet temperature as a result of the thermodynamic behavior (Fig. 6e). . For the conventional reactor, the inlet temperature was fixed for all SVs, and the maximum temperature was almost the same (the position of the maximum temperature shifted from the inlet to the outlet as the SV increased). As a result, the methanol productivity of the conventional reactor increased because the rate of increase in the flow rate (flow rate) was higher than the rate of decrease in the conversion rate. Conversely, methanol production in the compact reactor decreased somewhat with increasing SV.

To increase methanol productivity in the compact reactor, the inlet temperature was increased to 60 °C (orange bar labeled Comp-60), _{and the conversion was not calculated for SV/SV ref} = 1.0 because the maximum temperature was too high. Except that, _{it was observed that T feed} = increased compared to the case of 30 ℃. Furthermore, the increase in the raw material temperature increased the methanol selectivity. Although the maximum was obtained at the highest SV and feed temperature, too low conversion under these conditions significantly increased the separation cost. Thus, the titration conditions were SV/SV _ref = 2.0 and T _{feed = 90° C., which gave similar CO and H 2} conversions as under the reference conditions of the conventional reactor (cf. red dashed lines in FIGS. 6a and 6c), The productivity is similar to that for SV/SV _ref = 3.0 and T _feed = 120°C.

In order to increase the productivity, it may be considered to increase the diameter of the catalyst bed in addition to increasing the SV. 7 shows the effect of diameter on the performance of a compact reactor. From this, it was confirmed that CO and H ₂ conversion decreased with an increase in diameter when the SV value was fixed at 5000 mL/(g _{cat ·h), whereas the CO 2} conversion rate slightly increased ( FIG. 7a ). This characteristic is due to a decrease in the inlet temperature when the flow rate increases with increasing diameter for a constant SV (Fig. 7b). However, despite the reduced conversion, productivity increased due to the increase in the source gas. For better reactor efficiency evaluation, the productivity per CO conversion was calculated and shown in Figure 7b, which clearly indicated that increasing the tube diameter improved the efficiency.

7c and 7d show the axial profiles of CO conversion and temperature for conventional and compact reactors with different diameters. In the case of a conventional reactor, the reaction started at about z = 0.5 m, as the initial part of the catalyst bed serves to preheat the fuel gas. On the other hand, although the inlet temperature of the compact reactor was lower than that of the conventional reactor, the reaction started at about z = 0.2 m, which means that the temperature near the wall of the outermost tube was kept close to the wall temperature (240 °C). Because (Fig. 7e). Due to the high heat transfer capacity of the compact reactor, the maximum temperature in both the axial and radial directions was kept lower than that of the conventional reactor. Overall, assuming similar CO conversion rates, the diameter of the catalyst bed can be increased by up to 40% for the reference case and the productivity is increased by about 1.7 times compared to the reference case (2 times for the conventional reactor) did

In conventional and compact reactors with reference diameters, the reaction reached equilibrium in the center of the catalyst bed and was thereafter thermodynamically controlled ( FIG. 7f ). This indicates that the conversion was limited by equilibrium and the bed was not fully utilized. When the diameter increased by 20%, the reaction reached equilibrium near the exit of the bed, and further increases kept the reaction in a kinetic regime across the entire bed. It can be seen that the compact reactor is also more advantageous than the conventional reactor.

Claims (12)

A tubular reactor of multi-layer structure comprising inner tubes, catalyst tubes and shells of different diameters arranged to have a center of a cavity in turn from the inside, comprising:
The inside of the shell contains a cooling medium,
A source gas injection unit directly or indirectly connected to the inner tube is provided,
The catalyst tube is filled with a catalyst for methanol synthesis,
The lower part of the inner tube is closed and the upper part is an open cylindrical shape,
The lower part of the inner tube and the catalyst tube is blocked and the upper part is connected to each other,
A reactor for the production of methanol from syngas.
According to claim 1,
wherein the catalyst tube has a diameter of 1.5 to 3 times that of the inner tube.
According to claim 1,
The reactor is a tubular reactor of a triple layer structure provided at the bottom of the reactor and having a source gas injection part directly connected to an inner tube.
According to claim 1,
The reactor is a tubular reactor of a quadruple layer structure further comprising a source gas injection tube having a diameter smaller than the inner diameter of the inner tube disposed to have a center of the cavity at the center of the inner tube, wherein the upper portion of the source gas injection tube is the reactor It is connected to the source gas injection unit provided at the upper end of the reactor, the lower portion is open and located higher than the closed lower end of the inner tube to deliver the source gas to the lower end of the inner tube through this.
5. The method according to claim 3 or 4,
The reactor is a multi-layered cylindrical reactor having in one shell at least one subset consisting of an inner tube and a catalyst tube, optionally a source gas injection tube located at the center of the inner tube.
According to claim 1,
In the inner tube, the injected raw material gas is preheated by heat generated from the reaction in the catalyst tube while moving from the bottom to the top, the reactor.
7. The method of claim 6,
The fuel gas preheated while passing through the inner tube passes through the catalyst tube while CO hydrogenation reaction, CO ₂ hydrogenation reaction and reverse water-gas shift (RWGS) reaction are performed, the reactor.
According to claim 1,
The catalyst for methanol synthesis filled in the catalyst tube is a copper/zinc/aluminum-based catalyst, a chromium oxide-based catalyst, a copper-based catalyst, or a combination thereof.
According to claim 1,
In the interior of the shell, the cooling medium will absorb and vaporize the heat generated from the reaction in the catalyst tube.
According to claim 1,
The shell discharges the vaporized cooling medium to the outside through the gas outlet through a closed cooling loop system formed in connection with a water tank provided outside through a gas outlet formed at the upper portion and a liquid inlet formed at the lower end, and is cooled again to liquefy The reactor, wherein the medium is stored in a water bath and fed into the shell through a liquid inlet.
According to claim 1,
The reactor, which does not have a separate device for preheating the raw material gas.
According to claim 1,
wherein the production of methanol is carried out in a thermodynamic regime.

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