In the previous papers it was shown that with many systems, the rates of extraction from drops during their free rising were greatly decreased by the contamination of the systems with a trace amount of poisonous contaminating substances. In this paper, results of further studies on the effects of contamination are reported to make clear the total amount of mass transfer through one plate with one nozzle (D_N=0.101cm) (Cf. Fig. 1), in which the acetic acid is transferred from the dispersed phase of benzene to the continuous phase of water. Said total amount is affected not only by the amount of transfer during the free rising of drops which was already reported in the previous papers but also by the end effects of both coalescence and drop-forming.
Experimental results of l=5.0cm are shown in Fig. 2 for contaminated and non-contaminated systems, where it is shown that for both the systems, the rates are approximately the same, although in the case of drops reported previously, it was very small. This may be due to the large end effects of contaminated systems to compensate the decrease in extraction during free rising of the drops. The end effect at the time of drop forming ought to be considerably large as shown in Fig. 2, but it actually tells little, because the time required for drop-forming from liquid jet is too short to allow the drop to absorb the contaminating substance on to its surface.
Coalescence effect is also large in contaminated system, because, while in non-contaminated system the drops are broken down as soon as they reach the interface, in contaminated system, the drops are not broken but are accumulated at the interface and the solute is transferred during this accumulation.
Consequently, it is no use in the case of contaminated system to increase the plate distance so as to increase the total amount of extraction for one stage, unlike in the case of non-contaminated system. (Cf. Fig. 3.)
The effects of concentration of contaminating substance and column height on the extracted acid percent were investigated into and one example of the results of these experiments is given in Fig. 5.

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A theoretical treatment has been developed to determine the concentration of CaSO₄ in sea-water evaporators seeded with anhydrous CaSO₄ crystals to control the formation of hard scale of CaSO₄. Eq. (10) gives the transient behavior of the degree of supersaturation of the scaling material, in the evaporator liquor and Eq. (11) and Fig. 3, the behavior under the steady state. In the former the degree of supersaturation is found variable with c₀/c_s, c_F/c_sK, Kγ, and τ, and in the latter with two dimensionless parameters of c_F/c_sK and Kγ which cover entire operating variables of the seeded evaporators.
Published data on concentrated-sea-water evaporators seeded with anhydrous CaSO₄ crystals are analysed by using the correlation developed here, and summarized in Table 2 and Fig. 5, where the previous measurement of the growth rate of anhydrous CaSO₄ crystal is utilized to carry out the calculation. The correlation proves reasonable, since a good agreement is found between the calculated and measured concentration values as shown in Fig. 5 (a)5 (d). From this conclusion an evaporator design to give higher liquid retention is indicated equally effective with the increase of seed concentration in evaporator liquor to suppress the degree of supersaturation of scaling material, when the seed concentration is limited to a certain value to avoid operational difficulties due to denser suspension.

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It is believed to be very important to know the behaviours of particles in a moving bed, when designing an apparatus in which a moving bed is used; particulary so when a severe requirement is made of the apparatus as to the particle holding time.
The authors carried out observations and measurements with regard to the mixing degree of particles in a moving bed, by using tracer particles.
The apparatus used in the experiments is shown in Figs. 1 and 2. Solid particles used were steel balls, silica alumina balls and Soma sand, whose properties are listed in Tables 1 and 2.
Experimental methods employed were as follows. Tracer particles were added on to the top of the bed, and then the concentration of the tracer particles was measured at the bottom of the bed. Concentration curves of tracer particles at the bottom were as shown in Figs. 36, and coefficients of axial mixing of particles were calculated from these curves and Eq. (7).
As it was observed that there was a difference between the velocity of the particles lying in monolayer on the tube wall and that of the particles in the bed core (particles in bed except for the particles in monolayer on the tube wall), the experimental results shown in Table 3 may be summarized as follows:-
1. The concept of the axial mixing coefficient of solid particles, E_sx, is introduced on the assumption that the solid particles in bed may be regarded as to be in a homogenous phase. The axial Peclet number of particles in the bed core, (N_Pe=D_pu_c/E_sx), in a moving bed approaches 150 or so for Froude number, (N_Fr=u_c²/gD_p), 10^-310^-5. The mixing degree of particles calculated by N_Pe=150 agrees well with the experimentally observed data, as shown in Fig. 11.
2. The values of u_c/u_m and u_w/u_m are 1.1 and 0.85, respectively, and are independent of u_m for D_p/D_T =0.110.06, as shown in Fig. 9, where u_w is the velocity of monolayer particles on the tube wall, u_c is that of particles in the bed core and u_m is average velocity of particles in bed.

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In the previous report, we gave an account of the analysis of the mechanism of jet-pulverization observed in a small-scale experimental jet-pulverizing apparatus, and derived some assumptions, of which the proposed relation between mill-power (H) and mill-capacity (G) had the most important meaning. We proposed there the following equation:
G∝H^σ
in which σ_??_2.5}(1)
This equation shows that the jet pulverizer becomes more efficient as the mill-power increases. With most mills of other types, value of σ, which shall be called a coefficient of the size effect of pulverizer, is about 1.21.5.
In order to confirm this relation and other assumptions, we conducted some experiments with a laboratory-scale jet-pulverizer of practical use (Jet-O-Mizer 202 with 30HP compressor), and the results are given here to make clear the effects of variable factors, such as nozzle pressure, feeding rate of raw materials and diam. of nozzle, etc. on the rate of grinding.
This type of jet-pulverizer has a ring-shaped circular tube through which the mixture of air and particle circulates continuously, while some part of the mixture is taken out and sent to the collector successively.
The number of circulation, which was inversely proportional to the feeding rate was found to be one of the most important factors in this pulverization. The grinding action was strongest in the jet-stream, and it was negligibly weak on the surface of the mill wall. (Cf. Fig. 7.)
From these and other assumptions previously proposed, we derived Eq. 5.
X=100KA_N1P_N1+100(1-KA_N1P_N1)[1-(1-KA_N2P_N2)^2N] (5)
as a fundamental equation for calculating the capacity of the circulating type jet-pulverizer.
By comparing the capacity of this type of mill with that of the above-mentioned small-scale experimental apparatus (2HP) and of Jet-O-Mizer 808 (250HP), we substantiated the results of our experiment.
Data on Jet-O-Mizer shown in the information by Fluid-Energy Equipment and Processing Co. also verify Eq. 5. (Cf. Fig. 10.)
We had assumed, in the previous report that when mill-power became larger, the mechanism of breakage would change from surface grinding to bulk crushing and that the constant for the degree of surface grinding (n) would become small. The results of this experiment shown in Table 4 also verify this assumption.

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