In chemical reactions occurring within porous solid catalyst, the rates are affected not only by intrinsic catalytic activities, but also by the resistance of internal diffusion due to pore structures. Therefore, in the analysis of true kinetics, the decision of effective particle diameter or the estimation of catalytic activity, effectiveness factor for each catalyst should be predicted.
From this point of view, kinetics for low-temperature oxidation of ammonia with air in an iso-thermal flow system (inside diameter of reactor: 1.0cm) involving fixed bed of several coprecipitated porous Fe₂O₃-Bi₂O₃-MnO₂ catalysts (calcined at 550°C) were investigated at atmospheric pressure in steady state under the following experimental conditions: reaction temperature 180 to 340°, mass of catalyst 0.2 to 0.6g, average particle diameter 3.56×10^-2 to 14.10×10^-2cm, partial pressure of ammnia 0.25atm. and feed velocity 1.6 to 12.4cm/sec.
The experimental reaction rates under the resistance of negligible external diffusion were determined, and on the basis of Wheeler's theory the calculation methods of effectiveness factor were examined.
On the other hand, the physical properties of internal surface area, pore volume, average pore diameter, etc. as shown in Table 1 were measured by means of the B. E. T. and mercury-water displacement methods, and then the effects of pore structures on reaction rates were examined.
In this way, the effectiveness factor for each catalyst was estimated with considerable precision as mentioned below.
Since, the experimental results generally gave a first-order irreversible reaction with respect to ammonia (it was confirmed that under conditions described above, principal reaction was a conversion to nitrous oxide), effectiveness factor can be expressed by a function of a dimensionless modulus φas given by Eq.(1) for a single spherical particle or by Eq.(3) for a single cylindrical pore.
When the catalysts with different particle diameters were employed, the ratio of moduli φ became equal to the ratio of particle diameters as given by Eq.(5), for in Eq.(2) or (4), rate constant, pore radius and effective diffusion coefficient could be assumed to be approximately constant in the same catalysts, which was experimentally confirmed from the physical properties as shown in Table 1. The ratio of experimental rates (N) uninfluenced by external diffusion was equal to the ratio of effectiveness factor as given by Eq.(6).
Accordingly, by introducing η₃/η₁=N₃/N₁=α and φ₃/φ₁=D_p3/D_p1=β, the correlations given by Eqs.(8) and (9) were derived from Eqs.(1) and (3), respectively.
Now that the relationships between different particle diameters of catalysts employed were ap-proximately D_p3/D_pi=β=0.25, 0.35, 0.50 and 0.70, effectiveness factor could be easily evaluated from either the curves for Eqs.(1) and (8) or those for Eqs.(3) and (9), as given in Fig. 2.
Both effectiveness factors obtained from Eqs.(1) and (3) lay very close to each other, though the one obtained from the former equation was a little lower than that given by the latter, as shown in Table 3.
Comparison of effectiveness factors obtained from η₃/η_i=N₃/N_i=α and φ₃/φ_i=D_pi/D_pi=β proved that they held good for each particle. For example, each effectiveness factor η₁ showed good agreement within ±0.05 as shown in Table 3.

抄録全体を表示

There are two quite different types in chemical reactors, -piston flow type and complete mixing type, -so classified according to the absence or presence of back mixing. When the two cases are compared from kinematic point of view, it is found that in the former, the concentration driving force is larger, because there is no mixing in axial direction at all, while in the latter, the reaction velocity constant is observed to increase, because the temperature of the system is increased by mixing. Therefore we can not easily determine which type is more profitable. So far, many people have maintained that reaction velocity decreases in a tank type reactor. But the above discussion shows this is not always true. The only instance where it holds is found in an isothermal operation.
The auther studied this problem, using recirculation operation and found out important characteristics of the operation field. He defined this as “Adiabatic Characteristics of Reaction Operation Field” which can be employed to show which is the more profitable of the two types. This is the characteristic of operation field itself to be determined from its gas composition and its catalyst activity. He also called complete mixing type superior, piston flow type superior and intermediate type by the names of C, P and R type areas, respectively. It is very interesting to note that the mechanism of the decrease in volume of catalyst when the piston flow operation is turned into complete mixing operation is different from that in ordinary cooling operation. Economical evaluation of it will be made in the near future.

抄録全体を表示

A new simplified design method is presented for the design of fixed bed catalytic reactors.
The basic partial differential equations, (1)′ and (2)′ were reduced into the corresponding simpler forms, (15) and (16), or (17) and (18), by means of introducing the following assumptions besides the ordinary assumptions: 5) The radial profile of temperature (t) at any axial position is geometrically similar to that in the hypothetical model reactor, although the temperature level itself may be different from the one in the other. 6) The radial mass transport at the central axis of the catalytic reactor tube is negligibly small.
The calculation is to be conducted as follows: 1) In order to obtain the temperature and conversion profiles at the central axis, Eqs.(17) and (18) ought to be integrated simultaneously by stepwise numerical method. It is to be noted, however, that the values of the factor φ are not constant, but vary with the axial distance from the inlet. 2) The temperature at any radial position and the mean temperature at any cross section of the reactor tube can be calculated by Eqs.(21) and (22), respectively, using the value of t_cobtained in the first step. 3) The mean conversion curve is given by Eq.(29) or (29)′, using the value of t_m, obtained in step2.
Diagrams for φ and ψ as functions of the axial distance should be prepared in order to carry out the procedureed mentioned above. In this work the authors obtained the formulae representing these values, φ and ψ, for three kinds of reaction rate models and showed them in Figs. 2, 3, and 4.
Temperature and conversion distribution curves computed by the proposed method are in good agreement with the experimental data reported by Smith et al. and Wilson.(Figs. 5, 6, and 7)
It is also shown in Fig. 8 that the key assumption 5) is reasonable.

抄録全体を表示

The average degree of polymerization and the molecular weight distribution in the outflowing fluid from the homogeneous continuous reactors were theoretically estimated, assuming that the fluid mixing was expressed by mean, mixing diffusivity.
The effects of the fluid mixing on the degree of polymerization and the molecular weight distribution were found to depend on the polymerization mechanism. Provided that the polymerization mechanism consisted of the initiation reaction of zero order for monomer, of propagation reaction and of termination reaction by recombination, and that the concentration of active polymer was much smaller than that of dead polymer, the degree of conversion of monomer and the average degree of polymerization was increased with the decrease in fluid mixing. On the other hand, the spread of molecular weight distribution was decreased with the increase in fluid mixing at a certain mean residence time.
The effects of the fluid mixing was great in the range of M=0-5 and in higher degree of reaction yield.
The polymer properties obtained by means of diffusion model was approximated by means of the stirred tanks in series.

抄録全体を表示

To design equipments for industrial production of ethyl dioxymalonate (II), the process of the oxidation with N₂O₃ of ethyl malonate (I) was studied. The process may be considered as to be consisting of two steps: oxidation of (I) into a complex of (II) and oxides of nitrogen, and decom-position of the complex into (II) and oxides of nitrogen.
The first step proceeded under the low temperature of about 2° for 50 hours. The reaction in this step was vigorously exothermic. The conversion rate of the reactant and the yields of the products are shown in Fig. 1. Complexity of the reaction prevented us from deriving the rate equation. However, the mechanism of the reaction could be conjectured from the rate data above and some other observations, as shown in Table 1. Heat of reaction was also surmised as shown in Fig. 2. Data on a 30L stirred reactor coincided with those on a small vessel. A peak of reaction heat, about 12Kcal/hr/mol of (I) was observed to occur in the process of the oxidation.
Besides the above-mentioned experiment, some other reaction methods were tried: for example, the reaction by leading the N₂O₃ gas into the liquid of reactant (I) at high temperature (cf. Table 2). These results convinced us that only the low temperature method was suitable for industrial production.
The second step was that of the decomposition of complex. Methods employed for carrying out the second step were as follows:(a) batch wise method at low temperature (b) batch wise method at high temperature (c) continuous flow type method using a pipe at high temperature (d) continuous flow type method using a wetted-wall column at high temperature.
From the view point of safety and economy, a wetted-wall column was found to be the best equipment for the second step. The column used for experiments was of a stainless steel pipe of 20mm I. D. with a jacket of 65mm I. D. The length of the heated part was 1000mm. For heating was used the steam of 130°C flowing parallel to the flow of the solution.
Experimental results are shown in Table 5. The adequate capacity of the column determined by the experiments was 50-60g·Esoln/min, while the one obtained as the result of calculation in which Eqs.(2) and (7) were used on several assumptions was 47.5 and 67.5g/min, when λ=136.5 and 58.5 Kcal/kg (maximum and minimum values of enthalpy change by decomposition), respectively. These values coincided well with each other.
The design equation for the wetted-wall column was obtained as given by (8), in whose deriving a larger value of λ, 136.5Kcal/kg, was used for safety's sake.

抄録全体を表示