Pulsed packed column reactor

Fig. 26. Mechanically agitated industrial contactors, (a) mixer-settler (b) rotating-disk column (c) mixco column (d) asymmetric rotating-disk column (e) pulsed packed column (f) Podbielniak centrifugal extractor. (Reprinted from Doraiswamy, L. K and Sharma, M. M., Heterogeneous Reactions Analysis, Examples and Reactor Design, Vols. I and 2, 1984, John Wiley and Sons.)...

There are three wo well known types of column that can be used for continuous countercurrent liquid/liquid contacting rotating disc contactors (RDC) and pulsed packed columns (PPC) (these are not to be confused with three-phase packed coplumn reactors, operated under pulsed flow conditions, see section 4.7.2.3). 

When a narrow residence time distribution is desired, good heat transfer will usually require high flow rates, which in combination with a given residence time would result in long or tall reactors. Occasionally very long coaxial tubes are used for single liquid phase processes. For processes wiUi a liquid as the continuous phase, a cascade of stirred tanks (sections 3.3.3 and 7,2,1,3) or a pulsed pack column (section 45,15) may then be interesting alternatives. 

Although packed columns are simple and have no moving parts, their large space requirements have resulted in the replacement of packed columns by pulsed columns, or by other more compact contactors, in more recent installations for reprocessing irradiated reactor fuel. 

Pulsation seldom poses a problem in FI separation and/or preconcentration systems in which packed columns or long reactors are arranged on-line, or where gas-liquid separators are used. On-line separators often function as an effective pulse-damping system, so that pulsations are noticeable only when pump rotation speeds are extremely low, a condition which in any case should be avoided in FI operations. 

In spite of the widespread use of fixed-bed reactors, much remains to be done to define the dynamics of the reactor [30], Most of these reactors are operated in the concurrent mode at which the gas and the liquid both flow from the top to the bottom. A number of flow regimes have been distinguished for packed columns in downflow operation [31, 32], Based on the Reynolds number for liquid and gas flows, the flow regimes include (i) trickle flow, (ii) pulsed flow, (iii) dispersed bubble flow, (iv) wavy flow, and (v) spray flow (Figure 12.7) [17]. In general, countercurrent flows lead to much larger pressure drops across the bed, and this would be the case for FTS. Thus, the countercurrent flow mode is not used in today s plants. 

For readable discussions of plug flow and the concept of dispersion the reader is referred to the literature [1, 2, 17]. Here we shall confine ourselves to a brief illustration of axial and radial dispersion. Let us consider the behavior of a tracer introduced into a stream flowing through a packed bed reactor, as sketched in Fig. 7.12. Here (a) shows the axial dispersion or spread by the time it reaches the exit of a tracer pulse introduced uniformly over the cross-sectional area of the column at a given axial position. Figure 7.12b depicts the radial dispersion of a tracer introduced continuously at a point source. 

The catalytic activity was examined in the reaction of CO oxidation. Reaction was performed by a pulsed microcatalytic technique. Impulses of (2% CO + 1%C>2) in He were passed through the catalysts. Products were analysed at the output of reactor on 1 m packed column (Porapak Q) coupled with thermal conductivity detector. Impulses time interval was 5 min, because of reaction mixture analysis. 

Ross (R2) measured liquid-phase holdup and residence-time distribution by a tracer-pulse technique. Experiments were carried out for cocurrent flow in model columns of 2- and 4-in. diameter with air and water as fluid media, as well as in pilot-scale and industrial-scale reactors of 2-in. and 6.5-ft diameters used for the catalytic hydrogenation of petroleum fractions. The columns were packed with commercial cylindrical catalyst pellets of -in. diameter and length. The liquid holdup was from 40 to 50% of total bed volume for nominal liquid velocities from 8 to 200 ft/hr in the model reactors, from 26 to 32% of volume for nominal liquid velocities from 6 to 10.5 ft/hr in the pilot unit, and from 20 to 27 % for nominal liquid velocities from 27.9 to 68.6 ft/hr in the industrial unit. In that work, a few sets of results of residence-time distribution experiments are reported in graphical form, as tracer-response curves. 

The experimental data were largely obtained in bubble-flow and pulsed-flow regimes. A typical radial variation in the liquid holdup obtained under pulsed-flow regime is shown in Fig. 7-11. Runs nos. 1 and 2 in this figure are duplicate runs. Although the manner in which the column was packed may have had some effect on the holdup profile, it is clear from this figure that the liquid holdup profile was relatively flat in the center of the tube and was very sharp near the wall. It should be noted that the liquid holdup in this study was defined in terms of fraction of open reactor volume occupied by the liquid. 

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