T h e eighteen papers which follow were contributed b y some of Professor Sherwood’s m a n y professional friends and colleagues, all of w h o m have benefitted greatly from his influence and his generosity.
Some Characteristics of Two-Phase Flow in Monolithic Catalyst Structures Charles N. Satterfield’ and Fahri Orel Department of Chemical Engineering, Massachusetts Institute of Technology, Cambridge, Massachusetts 02 139
Measurements were made of the pressure drop during two-phase downward flow of water-air or cyclohexaneair through vertical stacks of monolith catalyst supports. Supplemental studies and observations were made with a single vertical glass capillary tube. The results are interpreted in terms of liquid distribution, of the conditions under which slug-type flow occurs in contrast to annular flow and in terms of the relative contributions to the pressure drop of constrictions, hydrodynamic friction, and gravity head.
The large-scale manufacture of monolithic “honeycomb” catalyst supports for use in automotive catalytic converters is making available for the first time in large quantity a new form of structure which may have a numberof potential advantages for use in reactors or other contacting devices in the chemical and petroleum industries. Of particular interest here was the possible application of stacked monoliths in tricklebed reactors or gas-liquid contactors in which a gas and liquid flow cocurrently downward through a bed of solid. As an exploratory study, measurements were made of pressure drop during two-phase downward flow through a vertical 2.54 cm diameter stack of monolith sections, 122 cm high. Individual sections were either 7.6 or 15.2 cm long and contained either 200 or 360 passageways (cells) per square inch. Stacked monoliths have several potential advantages for use in trickle-bed reactors or gas-liquid contactors with cocurrent flow: (1) They operate with a low pressure drop compared to conventional packed beds. (2) Their high compressive strength may permit deep beds to be constructed without the necessity of using intermediate supports and gas-liquid redistributors, which greatly increase capital costs. (3) The “controlled channelling” that they provide may give better contacting and better liquid distribution than that obtained in conventional trickle-bed reactors, especially at low liquid flow rates. (4) The uniformity of passageways may minimize axial dispersion. (5) With liquids containing fine solids, such as those derived from liquefaction of coal, bed plugging may be minimized or avoided. (6) In the case of an active catalyst where bed pressure drop is the limiting factor on the smallest size of conventional packing that can be used to minimize diffusion limitations, higher effectiveness factors can be achieved with monoliths. (7) In a cocurrent upflow mode the maximum allowable liquid and gas flow rates through a conventional packed bed reactor are frequently limited by the desire to avoid fluidization which can, for example, cause grinding of the catalyst. The use of monoliths may permit considerably higher flow rates to be utilized. Some possible disadvantages are: (1)The restricted radial conimunication of flow elements may make the contactors more sensitive to hot spots than conventional packing. (2)
Suitable methods of obtaining good initial distribution on a large scale need to be developed. (3) If used as a catalytic reactor under conditions in which the effectiveness factor approaches unity, approximately twice the vessel volume will be needed, since the solid fraction of monolith packing is typically about one-half that of conventional packing. As a first step in evaluating their potential, this study focussed on the hydrodynamic and two-phase flow characteristics of these supports. Particular emphasis was on the effect of gas and liquid flow rate on pressure drop. This is of interest in itself but in addition pressure drop gives valuable information on type of flow and flow distribution. As used on automobiles, a monolith catalyst support comprises a web of finely porous solid containing within it an array of parallel, uniform, straight, nonconnecting channels. They are currently manufactured by several companies and by more than one process, so that a variety of cross-sectional shapes, sizes, and wall (web) thicknesses are or can be readily formed. This study utilized monoliths supplied by the Corning Glass Works which are manufactured by extrusion of a thick inorganic dough, followed by drying and firing. The monoliths had square cross-section channels, either 200 or 360 channels (cells) per square inch of cross-section and a measured wall thickness of 0.27 mm in either case. The corresponding passageway dimensions (side of the inside square) were 1.53 and 1.07 mm, respectively, and void fractions were 0.70 and 0.60. The superficial area, a , is 18.6 and 23.1 cm2/cmi3of vessel volume, respectively. For comparable values of area/volume ratio with conventional packing it is necessary to use sizes much smaller than those conventionally used in absorbers or the smallest catalyst sizes used in trickle-bed reactors. For example, for spheres i?jz in. (2.38 mm) and I/,s in. (1.59 mm) in diameter, a is 15.8 and 23.6 cm2/cm:(,respectively. The 200 cell/in.’ material was identical with that currently being used in the U.S. automobile manufacturing industry. The flow regime here is highly unusual and there is little previous work to provide guidance on what to expect. The substantial literature on two-phase vertical slug flow is concerned with tubes larger than about 0.6 cm i.d. and much relates to nuclear reactor cooling. A few papers treat gas-liquid flow in horizontal capillary tubes with concern for measuring Ind. Eng. Chem., Fundam., Vol. 16, No. 1, 1977
61
Liquid
Gas
Figure 1. Distributor no. 3.
velocity by injecting an air bubble into the flowing liquid, but air-liquid ratios were low and no pressure drop measurements appear to have been made. Here we are concerned with gas-liquid downward flow through a parallel array of square capillaries, in which the ratio of gas to liquid may be different in different capillaries, and the distribution of the ratios over the array may shift a t each passage from one block of capillaries to the next. The flow may be either annular, in which the gas passes as a continuous phase through a moving liquid film on the walls, or slug-like. Which type of flow pervades is a function not only of gasliquid ratio, various physical properties of the fluids, and channel size, but also of the mode of entrance of gas and liquid into a channel and channel length.
Experimental Apparatus Monolith blocks were carefully core-drilled into sections exactly 2.54 cm in diameter and then cut into 7.6 or 15.2-cm lengths. These sections were then stacked on top of one another so as to constitute a column 122 cm high. In one configuration the outside of the column was first sealed with a special rubber type adhesive around which glass cloth was wrapped and around this a layer of silicone rubber. This was suitable for water-air studies but could not be used with organic liquids. In a second configuration, the monolith sections were fitted into 2.54-cm i.d. Teflon tubing, after first heating the Teflon to 250 "C to cause it to expand and then allowing it to shrink so as to hold the sections in tightly. In all cases the monolith ends were cut precisely perpendicular to the channels so that adjacent sections were in contact with one another across the entire cross section. This precaution is necessary since liquid will easily move radially (horizontally) across an exposed end of a vertically arranged monolith. Sections were stacked randomly so some offset would be expected between the array of cells in one monolith section and in that immediately below it. For the structures used here the void fraction is much greater than the solid fraction so a 62
Ind. Eng. Chem., Fundam., Vol. 16, No. 1, 1977
channel would never be completely blocked off from flow by the solid web of the section immediately above it. Liquid Distributors. As liquid flows from block to block it can move radially by a distance no greater than a channel dimension. Thus it is vitally important to obtain initially as close to uniform distribution as possible, especially in laboratory work where columns will be shorter than those encountered industrially. Several types of distributors investigated were found wanting. Three of some merit were further studied and the third of these was used for most of the results reported here. The first comprised a flat distributor head with an arrangement of 37 capillaries which protruded into the gas space to distances varying from about 3 to 7 mm and dripped liquid uniformly over the cross section. However, there are about 157 cells in a 2.54-cm diameter section of 200 cell/in.' material and liquid did not distribute itself uniformly through all these cells, as shown by the fact that the pressure drop was irreproducible and changed significantly with a slight rotation of the distributor; i.e., the scale of uniformity provided by the distributor was considerably larger than the cross-sectional area of the cells. In the second configuration a layer of 4-mm diameter spheres was placed on the top of the monolith stack and the capillary distributor was retained as a predistributor. Reproducibility improved but some flooding above the layer of spheres occurred a t high liquid flow rates; some channel entrances were probably effectively plugged by the spheres. None of the data obtained with the second distributor were retained since they were of dubious value. The best reproducibility and minimum pressure fluctuations were obtained by placing a sandwich of thin monolith slices (27 slices, each about 3.2 mm thick) above the monolith array, and this is termed distributor 3. The detailed design is shown in Figure 1. A 10 cm long Plexiglas (polymethyl methacrylate) tube which had a larger inside diameter than the column outside diameter was placed over the top of the column. The annular space between the Plexiglas tube and the column was sealed a t the lower end. The pre-distributor (array of 37 capillaries) was fitted and sealed to the other end of the 10-cm tube, separated from the top of the stack of monolith slices by about 3 cm. Gas was introduced through a side arm on the Plexiglas tube and moved upward through an annulus. This design minimizes disturbances of the liquid predistribution by the gas. A 1.8-mm tube, used as a pressure tap, was placed between the stack of monolith slices and the top of the monolith blocks. The bottom of the column was open to the atmosphere, but the same pressure drop results were obtained when measurements a t this point were made with a pressure tap. All pressure drop data reported here refer to that through the stack of monolith blocks as such. For the water-air studies, tap water was filtered through a 15-w filter and a trace of copper sulfate was added to prevent algae growth. Further information about flow circuits, metering and control is given by Ozel (1976). The pressure drop and, presumably, the degree of axial dispersion and mixing is affected by the type of flow achieved in the monolith channels, specifically by whether the liquid passes downward as an annular film or as slugs. T o obtain some idea of the hydrodynamic flow pattern as a function of gas and liquid flow rates in a single channel, some observations by still photography were also made of water-air flow in a piece of precision-bore glass capillary tubing. With both the water-air and cyclohexane-air systems, studies were made at superficial liquid velocities, V I ,from 0.33 to 6.58 cm/s and a t superficial air velocities, V,, from essentially zero to 150 cm/s. For comparison, liquid velocities in trickle-bed reactors typically are in the range of about 0.008 to 0.25 cm/s in pilot plant and laboratory equipment and about 0.08 to 2.5 cm/s in commercial reactors. Gas flow rates vary greatly, generally in proportion to the liquid rate used.
With conventional catalyst packing both fluids move as continuous phases throughout most of this regime although at the highest combination of gas and liquid rates of commercial interest the liquid pulses and the mass transfer and pressure drop behavior change (Satterfield, 1975; van Eek, 1977). Experimental Results Effect of Distributor a n d Block Length: Water-Air. The marked effect of the nature of the distributor and height of individual monolith blocks is demonstrated in Figure 2, for VI = 3.95 cm/s, and Figure 3, for VI = 1.31 cm/s, which show the observed pressure drop, AP/L, (cm of water per meter length of monolith stack) as a function of superficial air flow rate (cm/s at room temperature and atmospheric pressure) for a stack of 7.6-cm monoliths as compared to 15.2-cm monoliths. As noted above, with distributor no. 1the pressure drop varied substantially with a slight degree of rotation which presumably altered the degree of uniformity of distribution. The values shown are the highest observed, which occur with the most uniform distribution (see below). Regardless of the type of distributor, a negative measured pressure drop is encountered a t low gas and liquid flow rates. Judging from the studies with a single capillary (see later), if the liquid were evenly distributed over the cells, slug-type flow should occur and the hydrostatic head would cause this phenomenon. We did not measure axial dispersion of the gas phase but it is worth noting that with uneven liquid distribution, some reversal of gas flow could presumably occur, slug-type flow of liquid causing a pumping action to return some gas upward through those channels in which liquid flow, if it occurs at all, is solely as an annular film on the wall. With either a good or poor liquid distributor, pressure drop increases with an increase in either gas or liquid rate, as expected. However, reducing the length of individual blocks by one-half (doubling their number) has some unexpected effects. Figures 2 and 3 show that going from a stack of 15.2-cm monolith blocks to 7.6-cm monolith blocks in each case brings the values of P / L with the two types of distributor much closer together, the pressure drop with the poor distributor (no. 1)being raised and that with the good distributor (no. 3) being lowered. At the lower of the two liquid flow rates, the values of AP/L became quite close together for the two types of distributor when 7.6-cm blocks were studied. The offset between adjacent blocks causes a restriction in each channel which introduces some pressure drop. Calculations of these orifice effects (see later) showed that essentially doubling the number of junctions would be predicted to increase U / L at V , = 82 cm/s by about 1cm/m at Vi = 1.31 and about 2.8 cm/m at VI = 3.95. These are small relative to measured values of AP/L and clearly other effects must predominate. We suggest that two quite different and more important phenomena exist which work in opposite directions when the monolith lengths are reduced from 15.2 to 7.6 cm, the total column length being held constant. If one starts with the no. 1distributor, uniform distribution is obtained but on a larger scale than that of the cross section of the cells. Consequently, we postulate that some cells had little or no liquid flow while adjacent cells had more than the average. Presumably the offset between adjacent monoliths allowed mixing to occur on a micro scale upon passage between adjacent blocks, thus leading to more uniform distribution of liquid among the parallel passageways with increasing depth. This would cause a higher pressure drop per unit length with 7.6-cm than 15.2-cm monoliths. However, focusing on a single channel, the flow pattern changes with length (see later). In essence, as bubbles and slugs of liquid follow one another down a tube it is observed that a liquid ring frequently appears within a bubble, grows in thickness, finally forming a liquid web across the channel which splits the bubble in two. The
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appearance of additional menisci would presumably increase the resistance to flow. It is thus postulated that considerably more bubble breakup occurs in a 15.2-cm length than in a 7.6-cm length, accompanied by a higher pressure drop per unit length. (The original data using distributor no. 3 were obtained at values of V I= 0.328,2.63, and 6.58 cm/s (Ozel, 1976) and were cross-plotted for Figures 2 and 3.) The contrast to the behavior of conventional packed beds is noteworth. There, even with excellent initial liquid distribution the liquid may tend to gather into rivulets as it flows downward, especially at low liquid rates, and it also migrates to the wall. With monoliths liquid distribution appears to improve with distance and there is no wall effect. The lowest pressure drop per unit length occurs with poor liquid distribution but apparently the system adjusts itself to a higher value of M I L with distance. The results also imply that once good liquid distribution has been achieved, the wetting of solid may be considerably more effective than occurs with conventional packing, especially at relatively low liquid flow rates. This behavior, if substantiated, combined with the high area/volume ratio should make this a highly effective contactor for cocurrent flow. Another comparison of 7.6-cm and 15.2-cm blocks, but of 360-cell materiai is shown in Figure 4. Here, the pressure drop with 7.6-cm blocks is somewhat higher than with 15.2-cm blocks at an intermediate liquid flow rate, but little difference is seen at high and low flow rates. The relative effects of bubble breakup and orifice restriction presumably weigh differently here. Both 200-cell and 360-cell monoliths had the same web thickness, but the 360-cell material had substantially smaller passageways. With the smaller passageways a steady-state degree of bubble breakup may be reached more quickly, and if so, the effect of length would be less important. The channel restrictions between adjacent blocks of 360-cell material are substantially greater than with 200-cell blocks and so a higher fraction of the total pressure drop is caused by these orifice Ind. Eng. Chem., Fundam., Vol. 16, No. 1, 1977
63
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Figure 5. Comparison of pressure drop in monoliths (200-cell,15.2-cm blocks) with that in packed beds (McIlvried, 1956). effects. Indeed if the calculated orifice pressure drop is subtracted from the measured pressure drop for 7.6-cm and 15.2-cm blocks of 360-cell material (see later) for representative flow conditions, values of the remainder are very close to one another for 15.2-cm and 7.6-cm blocks. The remainder comprises the algebraic sum of frictional and hydrostatic pressure drop and should indeed be the same since hydrostatic pressure drop is the same for 15.2-cmand 7.6-cm blocks at the same gas-liquid ratio. Comparison with Packed Beds. Pressure drop data with distributor 3 for 200-cell monoliths at representative gas and liquid rates are compared to those for packed beds in Figure 5 . The available packed bed data most nearly comparable are those published for 4 and 6-cm glass beads (McIlvried, 1956). For mass-transfer limiting conditions the basis for comparison is the superficial area per unit vessel volume which for the spheres is 9.4 and 6.25 cm-l, respectively, compared to 18.6 for the 200-~ell/in.~ monolith. If intrinsic kinetics control, U / L(1- t) is a useful basis for comparison, where t is the void fraction. The packed fraction in the vessel, (1- e), is 0.625 for spheres and 0.30 for the monolith. It is evident that on either basis of comparison, pressure drops in packed beds are higher by an order of magnitude or more than in monoliths, at the same superficial gas and liquid velocities. Cyclohexane-Air System. Figure 6 reports the pressure drop observed with cyclohexane and air in 7.6-cm and 15.2-cm lengths of 200-cell stacked monoliths. As with water-air, AP/L is greater with 15.2-cmthan with 7.6-cm blocks. Comparison with water-air shows little difference between the two systems except at the highest liquid flow rates where the pressure drop observed with the hydrocarbon is moderately less. It is not clear what balance of physical properties leads to this close 64
Ind. Eng. Chem., Fundam., Vol. 16, No. 1 , 1977
similarity between the two liquids. No foaming occurred in any of these studies. Single Tube Studies. The pressure drop and, presumably, the degree of axial dispersion and mixing is affected by the type of flow achieved in the monolith channels, specifically by whether the liquid passes downward as an annular film or as slugs. To obtain some idea of the hydrodynamic flow pattern as a function of gas and liquid flow rates in a single channel, a large number of still photographs were made under a variety of conditions of water-air flow in a piece of precision bore glass capillary tubing 0.200 cm in inside diameter and 102 cm long. This is slightly larger than the hydraulic diameter of the larger monolith channels (0.15 cm), and the wetting characteristics of glass may be somewhat different, but the observations are useful. Over a considerable range the type of flow regime is determined by the mode of introduction of the liquid into the capillary, as shown in Figure 7. If water is introduced steadily down the side of a funnel atop the capillary, annular flow is seen to be established up to relatively high liquid flow rates, about 1.7 cm/s for this size capillary, and this limiting flow rate is independent of the gas rate from essentially zero to 150 cm/s. If, however, water is dropped steadily into the funnel (here the drops were each of 0.037 cm3volume) slug flow is observed at relatively low liquid rates but the value of V Iat which annular flow changes to slug flow varies substantially with gas flow rate. Within the slug flow region, photographs showed that the behavior and appearance of successive slugs of liquid changes substantially as gas flow rate is increased at fixed constant liquid flow rate. At a V Iof 0.79 cm/s, for example, increasing gas rate gradually decreases the length of liquid slugs and the distance between liquid slugs, of course, gradually increases.
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Liquid slugs become interspersed with annular rings of liquid which flow downward, the gas phase being continuous. If the flow pattern is initially annular in the intermediate region, an annular ring typically appears and as it flows downward it grows in thickness to the point that it fills the tube cross section, thus splitting the bubble in two. Consequently the flow pattern and average bubble size will be different at the bottom of the capillary than at the top, so the average behavior of a channel is a function of its length. The quantitative characterization in Figure 7 is limited to this particular size capillary but qualitatively these types of transitions would be expected to occur in small channels in general. Figure 8 shows pressure drop as a function of V Ifor various values of V , as observed in annular flow and Figure 9 the same for slug-type flow. In annular flow, APIL is independent of V Ifor a fixed low value of V,, reflecting the region in which the liquid film insignificantly reduces the effective diameter of the tubing and in which the pressure drop is caused essentially by the gas flow. At very low values of V , the drag of liquid on the gas causes a slight drop in AP with increased liquid flow rate. At the highest gas flow rates, AP is proportional to VI. Figure 9 shows pressure drop as a function of VI for various values of V, in the slug flow regime. AP is now a nonlinear
function of V Iand the pressure drop is now negative over a considerable range in which it was positive with annular flow. Thus pressure drop information can be used to obtain some insight into the flow pattern established in a monolith. Figure 10 compares the pressure drop observed with annular as compared to slug flow as gas rate is increased with a fixed representative liquid flow rate. As expected, under any combination of liquid and gas flow rates, transition from annular to slug flow substantially increases the pressure drop, provided that this is positive. A t low rates, however, under slug flow conditions the hydrostatic head of the liquid can cause a pumping action on the gas, producing a negative pressure drop. Figure 11shows values of AP/L vs. V Ifor a monolith stack and the capillary for comparable values of V Iand V,, where both are calculated per unit of open cross-sectional area. The capillary data are for dropwise introduction. The considerably larger pressure drop in the monolith reflects the somewhat smaller passageways, but the shapes of the curves are similar, reinforcing the belief that a similar type of flow occurs in both cases.
Correlation of Results The measured pressure drop, U r n may , be taken as the algebraic sum of the frictional pressure drop, the static head, gplLfl, and orifice effects APrn =
*f
+ *or
-~PIL~I
(1)
The static head was calculated from the known liquid volume fraction, fl, and AP,,, from eq 2 below. The frictional pressure drop, Ut,was then calculated from eq 1 by difference. The orifice effects are caused by the partial obstruction of the exit of one channel by the web of the next block. No precise method seems to be available on how to predict pressure drop for two-phase flow through a small restriction as here. If one assumes that gas and liquid move as slugs through the orifice Ind. Eng. Chem., Fundam., Vol. 16, No. 1, 1977
65
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Figure 12. Contributions of individual terms to measured pressure drop, 15.2-cm blocks, ZOO-cell monoliths, VI = 6.58 cm/s; waterair.
+
with the same velocity of Vt = V , V I ,the following equation results from the Bernoulli relationship
This assumes an inviscid fluid, a friction loss factor of unity, and applies a volumetric average density of the fluid mixture. Actually the true friction loss factor is presumably moderately less than unity. The minimum ratio of orifice area to cell cross-sectional area is 0.68 for the 200-cell monoliths and 0.56 for the 360-cell monoliths. The geometry here is such that as one randomly varies the pattern of block intersections this ratio would be expected to be encountered a large fraction of the time. Only when two adjacent blocks are nearly aligned would it be higher. As calculated from eq 2, APor is generally fairly small relative to APf as calculated from eq 1. I t is concluded that the frictional pressure drop, Ut.,is the dominating term over most of the range of conditions studied. Attempts were made to correlate APf by dimensional analysis and multilinear regression methods, but none of the relationships obtained seemed useful. The relative contributions of the three terms in eq 1to the contributes a sigmeasured pressure drop is such that AP,,, nificant fraction of the total only with the 360-cell material and at the highest liquid flow rates. For a fixed liquid flow rate, the static head and frictional pressure drop nearly cancel each other out a t a low gas rate but with increasing gas rate the static head contribution decreases and frictional pressure drop becomes dominant. These generalizations are brought out by Figures 12 and 13 for V I= 6.58 cm/s which illustrate the contribution of the three terms for two sets of conditions where the total pressure drop reached relatively high values. Figure 12 is for a stack of 15.2-cm blocks of 200-cell material and Figure 13 for a stack of 7.6-cm blocks of 360-cell material. The latter shows the maximum degree to which orifice restrictions contributed to the total pressure drop. Conclusions The pressure drop in two-phase flow in monoliths is very sensitive to the nature of the liquid distribution and to the initial distribution in particular. Pressure drop measurements may thus be useful in diagnosing flow patterns in these systems. The observed pressure drop was analyzed in terms of 66
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Ind. Eng. Chem., Fundam., Vol. 16, No. 1, 1977
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contributions from static head, orifice effects, and frictional pressure drop. Over most of the range of conditions studied, the dominating term was the frictional pressure drop. Stacked monolith structures may have considerable potential for use as gas-liquid contactors or in trickle-bed reactors. Unlike the behavior of conventional packing, there is no net migration of liquid to the vessel wall and liquid distribution appears to improve with distance. Comparison with conventional catalyst supports shows that on an equivalent basis, the pressure drop for two-phase flow in monoliths is an order of magnitude or more below that for conventional packing.
A Personal Note I had the pleasure of collaborating with Tom Sherwood on a variety of projects over a long period of time a t M.I.T. The present paper is the latest of a series on trickle bed reactors, the intricacies of which I was introduced to many years ago by Tom, who in turn had become aware of their potential in the early sixties through his association with the Union Oil Co. as a consultant. His was an enormously stimulating influence. Alive and alert to industrial problems on the one hand and to new theoretical developments on the other, he strove to make theory useful and to use practical experience in guiding studies of fundamentals. Withal, his was a generous spirit and he brought zest and exuberance to the teaching and practice of engineering. A mentor to many, he will be sorely missed. Charles N. Satterfield Nomenclature f l = volume fraction of fluid consisting of liquid g = accelerations of gravity, cm/s2 L = length, m AP = pressure drop, cm of water; AP, for measured value; APf for frictional pressure drop; APor for orifice effect; AP/L, pressure drop per unit length of monolith stack, cm of water/m V = superficial velocity, if not specified otherwise, cm/s VT = velocity defined as (VI + V g ) cm/s ,
Ozel, F., Ph.D. Thesis, Massachusetts Institute of Technology, 1976. Satterfield, C. N., "Mass Transfer in Heterogeneous Catalysis", Department of Chemical Engineering, Massachusetts Institute of Technology, Cambridge, Mass., 1970. Satterfield, C. N.. A.l.Ch.€. J., 21, 209-228 (1975). van Eek, M., Ph.D. Massachusetts Institute of Technology, 1977.
Greek Letters density, g/cm;j t = bed void fraction
p =
Subscripts g = gas phase 1 = liquid phase
Received for reuieu: August 3, 1976 Accepted September 30, 1976
Literature Cited McGreary, R. K.. J. Am. Ceram. SOC.,44, 513 (1961). Mcllvried, H.G., Ph.D. Thesis, Carnegie Institute of Technology, 1956.
Supported in part by the Scientific and Technical Research Council of Turkey in the form of a N.A.T.O. Scholarship to Fahri Ozel. The monoliths were core drilled and supplied by the Corning Glass Works.
The Effect of Additives on Mass Transfer in CaC03 or CaO Slurry Scrubbing of SO2 from Waste Gases Gary T. Rochelle and C. Judson King* Department of Chemical Engineering, University of California, Berkeley, California 94720
The theory of gas-liquid mass transfer with simultaneous chemical reaction is applied to the absorption of SO2 from waste gases into CaO or CaC03 slurries. The ratio of gas to liquid phase resistances varies with SO2 gas concentration, but is accounted for by a simple integration. The effective solubility of SO2 varies with bisulfite concentration and with SO2 gas Concentration. Alkali additives such as MgO and Na2C03 accumulate as dissolved sulfate salts and generate dissolved alkalinity as sulfite species and bicarbonate species that can react with SO2 to produce bisulfite. Organic acid additives should behave quantitatively as effective dissolved alkalinity if their pKa values are in the range 4.2-5.2. The mass-transfer theory explains the experimentally observed effects of SO2 gas concentration, alkali sulfate concentration, gypsum saturation, and organic acid concentration on SO2 removal. Only 3-12 mmol/l. of dissolved alkalinity should be required to derive most of the masstransfer benefit from alkali or organic acid additives.
Introduction There is much research and commercial interest in processes to remove SOn from waste gases. The most important source of such gases is the combustion of high-sulfur coal in power plants. A typical modern coal-fired power plant may emit 2 million SCFM waste gas containing 2000 ppm of SO2 by volume. The most popular processes being commercialized use alkali aqueous scrubbing with limestone or lime to neutralize SO, and produce CaSO 3:
-
+ SO2 CaO + SO2
CaCO 3
CaSO 3
-
+ Con
Cas03
Some CaSOj is also produced because of reaction with 0 2 in the waste gas. The waste gas is cooled to about 50 "C by adiabatic saturation with water in the scrubber. The CaSO?/ Cas04 solid product is disposed of as solid waste in evaporation ponds or as landfill. Hence these processes are classified as "throwaway" scrubbing. There have been three basic types of throwaway scrubbing processes-simple slurry, double-alkali, and slurry with soluble additive (Rochelle, 1977; Stern, 1976). In the simple slurry process (Figure la; Lessing, 1938; Epstein, 1975a, 197513; Head, 1976),the waste gas is scrubbed with a slurry of product solids and unreacted limestone. To avoid CaSO? and Cas04 crystallization and scaling in the scrubber, a separate crystallizer vessel is required with adequate residence time to control supersaturation of CaSOj and CaS04. Some of the CaO or CaCO 1 (depending upon which agent is used) dissolution must also occur in the crystallizer to avoid CaSO?
scaling in the scrubber. The slurry solids concentration is typically controlled a t 10-15 wt % by clarification of a bleed stream. In the double-alkali processes (Figure IC)(Kaplan, 1973), SO2 is absorbed into a clear solution of soluble alkali, usually Na2S03: SO:
