Ind. Eng. Chem. Res. 1997, 36, 513-522
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KINETICS, CATALYSIS, AND REACTION ENGINEERING Technology Design for Destruction and Removal of Chlorinated Chemicals† Jo´ zsef M. Berty‡ Berty Reaction Engineers, Ltd., 1806 Bent Pine Hill, Fogelsville, Pennsylvania 18052-1501
Sodium carbonate and carbonate minerals have been used in situ to neutralize the HCl generated during oxidation of chlorocarbons. The less than complete success of these procedures resulted from the lack of application of reaction engineering principles that were not widely known at that time. Now these are available and can be used to overcome some of the past difficulties. Using reaction engineering principles, catalytic metals were chosen whose chloride salts do not decompose easily. These are brought in close contact to the alkali carbonates to make the diffusion length to the solid-solid interface short. The carbonates are consumable reactants that regenerate the catalyst during operation and also act as a support for the catalytic metals. The reactor also works as a “dry scrubber” separator because the heteroatoms, as their salts, are removed with the moving solid. Integration of the rate equation shows much better performance from countercurrent moving-bed operation compared with that of cocurrent processing. For very toxic chemicals, an oxygen-fed, closed-loop, and below atmospheric pressure operating system can achieve almost complete conversion. The cycle gas is made from the small inert content of the technical oxygen and from the air that leaks into the system. To keep the contents of inert gas and other inert ingredients of the cycle at steady state, a small stream is released from the cycle. The release constitutes 0.025% of the cycle volume. Background There exists a well-known need to destroy halogenated hydrocarbon pollutants for environmental protection purposes. Thermal incineration is one solution. The EPA requires 1600 °C as the operating temperature for incineration, but this is expensive and generates some byproducts of incomplete combustion. Chemical reaction with carbonates needs somewhat lower, but still high, temperature. Catalytic incineration can accomplish the destruction of organic wastes at even lower temperatures, but chlorine atoms poison most catalysts, or at least reduce their activity. Still some were successfully commercialized. The catalysts used until now are metal oxides or noble metals deposited on acidic carriers (e.g., alumina, zeolites, clays). Carbonate-using technologies that had a limited success had a few common characteristics. The first condition is high temperature, necessary if a catalyst is not present. The second common condition is highly mixed beds. Gay et al. at Rockwell (1984), in their molten salt oxidation (MSO) process, used a pot of molten soda and in their secondary treatment a packed bed with concurrent contactor and recycled molten soda. Dow (1973) used sodium carbonate in a fluidized bed below the melting point temperature, and General Atomics (1989) used a recycling fluidized bed of limestone at calcination temperature. In the fluid beds the reacting components are in a strongly mixed condition, †
Paper presented as a poster at ISCRE 13 in Baltimore, MD, on Sept 25-28, 1994, and an extended version at the Days of Technical Chemistry, 96 Veszpre´m, Hungary, April 23-25, 1996. ‡ Telephone: (610) 391-1676. Fax: (610) 391-9434. E-mail:
[email protected]. S0888-5885(96)00655-0 CCC: $14.00
necessary for autocatalytic, combustion-type operation. An earlier Rockwell (1981) process used a fluidized bed and an afterburner with a toxic, chromic oxide catalyst deposited on alumina, at below the melting point temperature for soda. Separate sodium carbonate granules were also added to neutralize the acid in situ in the dry state. The consequence of high temperature with high mixedness is that homogeneous gas-phase combustion is going on in addition to the reaction with the carbonates. Even in the presence of carbonates, not all HCl can be neutralized instantly and some remains free long enough to cause corrosion damage in the reactor and so require a scrubber. In all catalytic oxidations of halocarbons the first step can be viewed as metal chloride formation:
Cr2O3 + 2C2HCl3 + 3O2 ) 2CrCl3 + 4CO2 + H2O The metal chloride then hydrolyzes to HCl, which goes into the vapor phase and this regenerates the Cr2O3 catalyst:
2CrCl3 + 3H2O T Cr2O3 + 6HCl The HCl can be oxidized to chlorine, just like in the Deacon process (1870):
4HCl + O2 T 2Cl2 + 2H2O Arnold and Kobe (1952) clarified the equilibium conditions for this process. The generated chlorine, at incomplete conversion of the chlorocarbons, can further chlorinate these, making more chlorinated, more toxic, and more refractory compounds. In the earlier Rockwell process, HCl was to be neutralized by added granular © 1997 American Chemical Society
514 Ind. Eng. Chem. Res., Vol. 36, No. 3, 1997
sodium carbonate. The HCl then had to diffuse out of the supported catalyst, through the gas phase, and to the sodium carbonate granule. Once there, the HCl still had to diffuse through a salt layer to the carbonate, in order to be neutralized. Some present catalysts can also oxidize chlorinated hydrocarbons (Lester and Summers, 1988). These maintain catalytic activity by using catalytic metals and operating conditions wherein the chlorine-metal compounds are unstable. Under such conditions, while the carbon and hydrogen parts are oxidized, HCl is rejected into the gas phase. Catalytic activity is depressed, and complete conversion is practically impossible because of the presence of HCl and Cl2 (Rossini and Farris, 1993). New Catalyst System A new catalyst system was proposed (Berty, 1991) that consists of metal oxides or carbonates deposited on alkali or alkaline carbonates, with sodium carbonate (soda) preferred for halogenated compounds. Reaction over this material is visualized as a reaction of gaseous pollutant with a solid, catalyzed by another solid; therefore, it was considered that the two solid phases should be in intimate contact. The sodium carbonate is a consumable reactant that internally regenerates the catalyst-halogen compound during operation, besides being a “carrier” for the catalytically active metals. Most combinations of the usual heavy-metal catalysts work with alkali and alkaline-earth carbonates as reactants and supports. In the case of this catalytically activated sodium carbonate, active metal salt is impregnated or intimately dispersed in an alkali-metal carbonate. The concept is similar to that used by Weisz (1962) for a bifunctional catalyst in the isomerization of n-heptane. In our case, the catalyst and the regenerating solid have to be close together to minimize solid-state diffusional length. The reaction of trichloroethylene (TCE) is complete at 300 °C at 10 000 h-1 GHSV:
The combination of manganese with copper on sodium carbonate is very effective. A manganese/magnesium mixture on soda is almost as effective and is entirely nontoxic before or after use on chlorinated hydrocarbons. Reaction takes place at 300-500 °C and at 500010000 GHSV. This catalyst is supported on a solid base, in contrast to all present catalysts which are all supported on solid acids for the destruction of volatile organic compounds (VOC) and halogenated organics. New Technology Sodium carbonate can react without catalyst with the chlorinated materials at 800-960 °C temperature as in the molten salt oxidation (MSO) process of Rockwell International (Gay et al., 1984). Catalytic metals can react alone. However, they become deactivated before a quantity of poisons, equivalent to the quantity of the catalytic materials, have been fed. In the proposed solid, sodium carbonate is in a large excess over the catalytic metals and these metals are mixed and intimately distributed in the sodium carbonate. More than 90% of the soda can be converted to salt before a major drop in activity sets in. Some of the chlorinated pollutants are very toxic, and these must be destroyed to the highest level of destruction and removal efficiency (DRE). If no byproducts are generated, this is equivalent to the requirement of very high conversion. It is known from kinetics that for high conversion plug-flow reactors are needed, where longitudinal backmixing must be avoided and very narrow residence time distribution (RTD) maintained. Thermodynamics. Example of Reaction for Trichloroethylene (TCE):
Overall:
in steps of:
1.
3MgCO3 + 2C2HCl3 + 3O2 f 3MgCl2 + 7CO2 + H2O The close contact between the minute size particles of metal chloride and sodium carbonate enables the regeneration of the catalyst from the chloride salt directly, without going through HCl in the vapor phase:
MgCl2(S) + Na2CO3(S) ) 2NaCl(S) + ∆G°r ) -92.18 kJ/mol MgO(S) + CO2(G), The depicted reactions are only the gross steps, leaving out many details. This reaction starts at 50 °C and proceeds at an industrial rate above 200 °C. If soda is not present, the regeneration of the catalyst can occur through the decomposition of the metal chloride,
MgCl2(S) + H2O T MgO(S) + 2HCl,
∆G°r ) 60.92 kJ/mol
or this combined with the Deacon reaction, which is less difficult,
MgCl2(S) + 0.5O2 T MgO(S) + Cl2,
∆G°r ) 22.79 kJ/mol
Cu/Mn
2C2HCl3 + 3O2 + 3Na2CO3 ) 6NaCl + H2O + 7CO2
2C2HCl3 + 3O2 + 3CuO ) 3CuCl2 + H2O + 4CO2
at T ) 298 K, ∆Hr ) -1.99 MJ, ∆G ) -1.98 MJ, K ) exp(800) 2.
CuCl2 + Na2CO3 ) 2NaCl + CuO + CO2
at T ) 298 K, ∆Hr ) -44.4 kJ, ∆G ) -144 kJ, K ) exp(58.2) The first reaction step occurs between the gaseous contaminant and the solid catalytic metal oxide while oxygen is supplied from air to oxidize the organic parts. This is followed by internal regeneration of the catalytic metals by transferring halogen to sodium carbonate. These two steps represent only the two major phases of this reaction and should not be considered as the detailed mechanism. In the early studies, combinations of various heavy metals on various alkali and alkalineearth carbonates were used and all worked to different conversions. During the 3-mo studies, a washer filled with acidified AgNO3 operated in the discharge line. No silver chloride precipitation was ever seen during 3 mo of testing while the reactor was operated with alkali and alkaline-earth carbonates, intimately mixed with various catalytic metals, and operated at various conversion levels at various temperatures. It is therefore safe to
Ind. Eng. Chem. Res., Vol. 36, No. 3, 1997 515 Table 1. First-Order Rate Constants and Activation Energies catalyst no. 1 1 1 1 3 3 3 7 1 and 3 1% Pt 1% Hop 11% Hopcalite
run reaction metal in surface area, activation energy, temperature, kinetic constant kinetic constant kSA, 1000/cm2‚s no. time, s support, mol/mol cm2/g kcal/mol K km, 1/g‚s 11 13 15 22 12 10 14 16 17 20 23a 24
0.72 0.72 0.72 0.72 0.72 0.72 0.72 0.72 0.72 0.60 0.71 0.73
0.20 0.14 0.09 0.12 0.19 0.20 0.20 0.12 0.16 0.026 0.05 0.52
180 180 120 150 180 180 180 120 180 100 25 31
Figure 1. Ignition curves for TCE and TCB at GHSV of 10 800 and 7550 h-1.
say that no HCl or Cl2 was ever present. This is in agreement with the previously calculated free-energy changes for CuCl2 and MgCl2 hydrolysis, which are both highly positive, hence improbable. For the reaction of these with Na2CO3, the free-energy changes are around -100 kJ, hence spontaneous. Laboratory Studies. Laboratory experiments and the reactor system for catalyst development were described previously (Berty, 1994; Berty et al., 1992; Stenger et al., 1993; Tavakoli et al., 1996). Briefly, a tubular microreactor containing 3 g of catalyst was used for the semicontinuous experiments. Analysis was made by a gas chromatograph, HP5890A, equipped with a flame ionization detector (FID). Discharge gas was monitored by acidified AgNO3 for free HCl and Cl2, and no trace of these was ever indicated. During catalyst screening, conversion vs temperature curves at constant feed rate, i.e., the “ignition curves”, were measured, and on the better catalysts chlorine balances were calculated between gas and solid. Table 1 gives some results of the early catalyst screening work. Catalysts listed in the first colum are different combinations of Cu/Mn/Mg deposited on various carbonates. Catalyst No. 7 was the first Cu/Mn on sodium carbonate. The 1 and 11% Hop indicate the quantity of commercial HOPCALITE catalyst mixed in various alkaline supports. Figure 1 is taken from work of Javad Tavakoli et al. not reported in 1993. This depicts the ignition curve for trichloroethylene and trichlorobenzene. For example, in Figure 1 it can be seen that to get 50% conversion of TCE 270 °C was needed at GHSV of
13 13 13 13 14 14 14 10.4 13 3 3 3
673 673 673 623 673 673 673 713 638 673 572 594
3.5 12 23 6 3.3 3.5 2.5 8 3 225 15 20
4 9 18 5 4 4 3 8 3 6 29 33
10 800 h-1 and 57 ppm initial concentration. For TCB (trichlorobenzene) 350 °C was required to get 50% conversion at 7550 h-1 GHSV with 31 ppm initial concentration (Tavakoli et al., 1996). Chlorine balance could be closed within (6% error for the chlorine found in the solid phase as sodium chloride/(removed from the vapor phase) working with 3 g of solid (Stenger et al., 1993.) A small initial halogen content of the catalyst is needed for high activity at the start. This was observed on the regenerated catalyst that was washed chloride free, and this also matches experiences with other similar processes, suggesting that adsorption of contaminants occurs preferentially on chlorinated sites, especially for hydrocarbons, while oxygen adsorbs on metals (Ostrovskii et al., 1962). Chlorination of hydrocarbons and generation of lower molecular weight fragments occur at low temperatures only. Kinetics of the Process. For a working hypothesis, it was assumed that the kinetics of the overall process could be described by a grain-size model with the chemical rate controlling because of the slow rate. For the grains, the sharp interface model was assumed. This was based on the granulated and pelletized catalyst manufacture and on the models of Doraiswamy and Kulkarni (1987). In contrast to processes where selectivity problems govern, here the need for total destruction places the emphasis toward the tail end of the conversion. Because of this, several times the catalyst volume needed for 90% conversion will be used, and the significance of this first section diminishes, tolerating more errors than usual. It was shown experimentally that at low concentrations of pollutant, with oxygen in air in large surplus, the rate is first order to pollutant. Under such conditions the conversion of the solid hardly changes during a few hours, while the temperature effect is measured. This was exploited to estimate the energy of activation from some “ignition-curve” measurements. For Ea, Stenger et al. (1993) found 50 MJ/mol, and evaluation from the measurement of Tavakoli et al. (1996) resulted in 60 MJ/mol. These measurements were made at different laboratories and on similar yet not identical solids. These are low values for catalytic reactions, and they may indicate some diffusional influence. At high concentration of pollutant, even with an excess of soda and oxygen, a simple first-order model does not adequately describe the process. The interest here is at very high conversion, therefore, behavior at very low concentration is the most important. A measured and fitted ignition curve for TCE is shown of Figure 2. Calculations leading to this are given in Appendix A.
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Figure 2. Measured and fitted ignition curves for 31 ppm TCE at 10 800 GHSV.
Derivation of the Model. Several models are known for solid-gas reactions (Doraiswamy and Kulkarni, 1987) for various controlling steps. Here a simple model was used for the overall rate with the customary assumptions and for the chemical rate controlling. For the sharp-interface model, where the chemical rate controls, the rate expression is
r ) kC0y(S/Vr)
(1)
where k ) f(T,O2) and C0 is the initial concentration in mol/m3, y is the mole fraction of the contaminant, and S/Vr is the reacting surface to reactor volume ratio. With S/Vr expressed in terms of fractional conversion of solids X, the rate is
6(1 - )(1 - X)2/3 dp
r ) kC0y
or
or
and Fy < 0,
-K(yi - y) ) Xi - X
RXFy and RyFS K (-dy) ) dX (3)
where Fy and FS are the molar flow rates, and RY and RX are the stoichiometric coefficients for the pollutant and solid, respectively. The rate can now be expressed for the moving bed reactor as a unique function of y. By integration of eq 3, the equivalence can be expressed. In steady state, to express X ) f(y) for cocurrent downflow operation of the solids and gas, where FS < 0
Fy (-dy) (5) Ry dVr
are the equations to integrate to receive the performance of the reactor. For countercurrent operation, where FS < 0 but Fy > 0, the equations are
K[yi - y] ) Xi - X
and 1 - X ) 1 - X0 + K(yi - y) (6)
r ) k′′y(1 - Xi + K(yi - y))2/3, r ≡
(2)
K)
and 1 - X ) 1 - Xi - K(yi - y) (4)
r ) k′′y(1 - Xi - K(yi - y))2/3, r ≡
r ) k′y(1 - X)2/3
The dp is constant, because the diameter of the reacting core is shrinking only. This constant diameter is incorporated into the other constants since this is not the size of the pellet but the diameter of the powdery granules before these were compacted to the ∼3/8 in. size pellets. Details of diffusion calculations are given in Appendix B. Conversion of the pollutant and of the solid are connected by stoichiometry. This is expressed for an incremental reactor volume by
Fy (-dy) FS dX ) Ry dVr RX dVr
Figure 3. Simulation of startup-consumption of Na2CO3 by volume and time.
Fy (-dy) (7) Ry dVr
For transient operation of the semicontinuous type, where FS is the soda content in the reactor in mol/m3, FS ) 0, and Fy > 0, the equation to integrate is
FS
-∂y Fy (-∂y) + ) k′C0FSy(1 - X)2/3 and ∂t Ry ∂Vr dX ) K′ (-dy) (8)
where, at Vr ) 0, y ) yi at all values of t and, at t ) 0, X ) 0 at all values of Vr. This model satisfactorily approximates the main features of the countercurrent process, and it is acceptable for use in design. Figure 3 gives the simulation of conversion changes in volume and time for the first 24 h of operation of the first 10 L of a pilot-plant reactor. Pilot-Plant Test. The aim of the pilot-plant test was to evaluate the fraction of the carbonate in the solid that can be converted to salt, i.e., the chloride retaining capacity of the solid. In a pilot-plant experiment, lasting 101 h, a total of 14.5 kg (102.5 mol) of trichloroethylene (TCE) was evaporated into 225 m3/h (10.9 kmol/h) of air, resulting in about 100 ppm initial concentration. This stream was fed through 18.6 kg of OXITOX
Ind. Eng. Chem. Res., Vol. 36, No. 3, 1997 517
the numbers were rounded out to give an easier visualization. The experiment was semicontinuous, with the gas flowing continuously while solids were fixed. The solid soda content gradually changed to salt over a period of 4 days. In real operation, the catalyst on the first or bottom tray would have to be dumped once a day, all higher trays would move down, and the top would get a fresh catalyst. In Figure 4, hourly quantities are shown for the changing (but not flowing) solids also, to adhere to the hourly basis of the calculation. Reactor Calculation
Figure 4. Conceptual flowsheet of the pilot-plant experiment in Veszpre´m, Hungary.
(registered tradename of Berty Reaction Engineers, Ltd., for the catalytically activated sodium carbonate) containing 132.1 mol of Na2CO3, spread out on four trays in a 0.5 m φ cylindrical reactor. The activated soda used here was regenerated from one that was used in a previous pilot-plant experiment, by dissolving salt and unused sodium carbonate with water. Fresh sodium carbonate was added to the insoluble catalytic compounds, and the mixture was pelletized. Figure 4 shows a conceptual flowsheet of the pilot plant. Below the flowsheet, results are summarized on the spreadsheet, with the hourly molar flows given. The basis for calculation was an actual pilot-plant run where
Figure 5. Technology scheme.
Once-Through Operation for Air Cleaning. Technology and catalyst can be designed together for conversion of pollutant to over 99.9999%. This is based on previously mentioned kinetics and on using a countercurrent moving bed reactor and a partially chlorinated catalyst. Designing catalyst/reactant and reactor/ separation equipment is the main task. Therefore, both rate and equilibrium problems have to be solved. The equilibrium in the solid-phase reactions is entirely on the product side, and diffusion between solids could limit the attainment rate of the equilibrium. Due to the intimate contact of catalyst and soda inside the grains, this equilibrium is maintained. Between the grains and inside the pellets some diffusion limitation can exist at higher concentrations. Higher pore volume would help diffusion yet would significantly weaken the compression strength of the pellet. To ensure the smooth flow of the solid, a minimum size of 1/4 in. is recommended in the literature. To minimize pressure drop, an even larger size would be better. The optimum pellet size is between 3/8 and 1/ in. sizes. 2 Plug flow of both gas and solid is important for both reaction and separation. To approximate plug flow for gas, Rep > 100 will be maintained to maximize mixing in the radial direction and a dimensionless bed depth of L/dp > 150 will be used to minimize axial mixing.
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Figure 6. Volume of reactor needed for low conversion of mustard gas.
Figure 8. Closed loop for destruction of mustard gas and toxic liquids with oxygen.
Figure 7. Reactor volume required for conversion of mustard gas for greater than 90%.
These conditions for a given flow define the bed size and shape and usually result in larger than minimum needed reactor volume Vr. The proper size of the pellets and that of the discharge nozzles for the solids result in plug flow for the solids. The integration of the rate equations for steady-state operation was done by the Romberg method as presented in the Mathcad program of MathSoft Inc. At 100 ppm TCE in the feed, the rate is dependent both on TCE concentration and the available surface. At 1 ppm TCE, the rate starts to behave as a first-order reaction. For the partial differential equation the “double stepwise integration” method of Grossman (1946) was used by replacing the graphical method with an equivalent system on a spreadsheet. In handling toxic substances very high destructionand-removal efficiency (DRE) is needed. For the separation the hypersorption process (Berg, 1946) was used as an overall example. This process was a continuous countercurrent adsorption method for separation of light gases. It worked well except that the active charcoal, available 40 years ago, was not strong enough to withstand the continuous erosion. The cost of active charcoal losses therefore made the process uneconomical. The Thermofor catalytic cracking (TCC) process was another model to emulate for the moving solids operation (Newton et al., 1945). Extractive distillation
and the production of methyl butyl ether in a distillation column are conceptually similar processes. On the basis of the previous examples, the technology scheme considered is shown in Figure 5. Here, the contaminated air or toxic liquid is introduced somewhat above the bottom. At the bottom, hot air is fed for purging the solids before these are discharged. This part is similar to a reboiler. It is a safety measure to react any toxic material before the solids are discharged. Some toxic material could have entered the purging zone by diffusion, by imperfect plug flow, or by adsorption on the solids, and these are removed there. The solid enters at the top of the reactor and slowly, or even periodically, slides downward. At both the entrance and discharge, the solids are led through air gates. Even if leaks cannot be prevented completely, the below atmospheric pressure operation will cause air to leak in, rather than any toxic material to leak out. The solid catalyst/reactant mixture is inexpensive. Therefore, more than needed can be charged to the reactor to increase DRE. This increases the inventory of the solid, the size and cost of the sheet-metal body, and the pressure drop but not the consumption of the solid. While the energy needed for the blower also increases, this is the least of the operating expenses. Oxygen-Fed, Closed-Loop Operation for Very Toxic Substances. For destruction of very toxic substances, very high conversions are needed. Since catalyst volume used is several times that needed for 90% conversion, most of it will work with low ppm of pollutant. At this low concentration the first-order decay law will control. An example for a proposed destruction scheme, mustard gas (dichloroethyl sulfide, S(C2H4Cl)2), is used to illustrate an extreme problem. In this case a magnesium carbonate is used instead of sodium carbonate for the reactant carrier to avoid certain eutectic mixtures in the solid phase. Ten percent of the total magnesium compound fed is already in salt form to assure a smooth start; another 10% is used as excess to assure high conversion of mustard.
Ind. Eng. Chem. Res., Vol. 36, No. 3, 1997 519
Figure 9. Conceptual layout proposed for mustard gas destruction.
The overall reaction is Mn
S(C2H4Cl)2 + 7O2 + 2MgCO3 ) MgSO4 + MgCl2 + 6CO2 + 4H2O With excess magnesium carbonate, partially converted at the entrance, the design basis is
S(C2H4Cl)2 + 7O2 + 2.25MgCO3 + 0.125MgSO4 + 0.125MgCl2 ) 0.25MgCO3 + 1.125MgSO4 + 1.125MgCl2 + 6CO2 + 4H2O Since no experimental data for mustard destruction were available, the kinetics for TCE destruction were used. This helps to study the problems before any work is done. The first observation is that exact kinetics will be needed only to scale the first fraction of the reactor where large mustard concentration is, while in the rest the first-order decay law will apply again. Results will be more than qualitative; only scaling of the graphical results will be needed. The rate equations for the cocurrent and countercurrent cases were integrated for the corresponding bound-
ary condition. Results of integrations, for the given boundary conditions, are shown in Figures 6 and 7. Figure 6 shows the reactor volume needed to achieve up to 90% conversion of the 1000 ppm mustard. As can be seen below 80% conversion, the cocurrent method needs less reactor volume. At this point the lines cross and above 80% conversion the countercurrent case needs less catalyst-charged volume. Since the goal in destroying toxic chemicals is to obtain the highest achievable conversion, the high conversion case was studied, and the results are illustrated in Figure 7, with a log scale for ppm concentration of mustard. Here the integration was done between increasing orders of magnitude for conversion and, therefore, no points are shown between 1000 and 100 ppm; subsequently, the crossing effect shown in Figure 6 cannot be seen on Figure 7. Figure 7 shows that, above 90% conversion, the destruction of 90% of the leftover pollutants always takes the same additional volume. In other words, above 90% conversion the system behaves close to a first-order reaction and this is expected at very low concentrations of reactant. In this case the countercurrent method needs less than half of the catalyst volume needed for cocurrent operation.
520 Ind. Eng. Chem. Res., Vol. 36, No. 3, 1997
For very toxic materials, like chemical warfare agents, oxygen can be used in a closed loop. Here, the accumulated inerts (argon from oxygen) and the water and CO2 produced can be used as diluents and only these gases require disposal. The used solid catalyst/carbonate, as well as the water and carbon dioxide formed in the oxidation, must be removed to maintain steady state. All materials removed from the reactor, including the off-gases, have to be stored for 3 days until detailed laboratory analysis confirms the total absence of any trace of toxic materials (NRC, 1993) and before these can be disposed. The conceptual scheme for the closed-loop operation of the reactor is shown in Figure 8 for mustard gas, which is an oily liquid boiling at 217 °C. The per pass destruction efficiency in this example is designed to be 99.99%. It is possible to design for much higher per pass efficiency, but here the 99.99% was used to show the efficiency increasing effect of the almost closed loop. The DRE is significantly higher than the per pass efficiency due to the 4000:1 recycle to discharge flow ratio. In short, 99.99% of the 0.1 ppm mustard in the recycle will also be destroyed. Only the 0.1 ppm will survive in the small discharge stream that has to be vented to maintain steady state. This is about 0.25 mol/1000 mol of recycle. In other words, the recycle to discharge ratio of 4000 is close to complete and the DRE is 99.999 998%. This unit would be started up by air and oxidizing acetone. Once the operation is smooth, air can be replaced gradually by oxygen, never exceeding 25 vol % O2. As the operation is stabilized again, half of the acetone gradually can be replaced by mustard. In case of operational trouble, the mustard feed can be cut off immediately. The inventory of mustard in the system is very low. It consists of 0.1 ppm of the small cycle volume and the log mean concentration of 10 ppm (from 1000 and 0.1 ppm) in the very small volume between the reacting solids. The cycle volume contains about 30 mg of mustard gas only. The requirement to store every discharged material, including the off-gases, for 3 days requires the use of industrial oxygen with 0.5 vol % total inert content. This then sets the minimum gas discharge rate, and this can be stored in medium-pressure gas cylinders. The plant layout for the full assembly can be seen in Figure 9, where the mid-level of the structure is made gas-tight, with instruments and motors located behind a protective wall. In Figure 9, the center square is a reduced form of Figure 8, where details can be seen better. To achieve this high DRE, the total inerts that leak into the destruction system must be kept very low. Here it is assumed that leaks amount to not more than 0.5% of the volume of O2 fed. That is about again as much as the quantity of inerts getting in with the technical oxygen. This then puts stringent requirements on construction quality and requires many additional measures to prevent leaks. These are beyond the scope of this paper. It was recently realized that low-level radioactive waste that is stored in chlorinated solvents could be reduced to solids in a scheme similar to that proposed for mustard gas. Conclusion It was shown that alkali- and alkaline-earth-supported catalysts can totally destroy halogen-containing organic compounds without generating corrosive acidic byproducts, and therefore no acid scrubber and no corrosion-resistant construction is needed. Countercurrent moving beds assure the highest destruction ef-
ficiency. A closed-loop, recycle system fed with oxygen can approach the total destruction and removal of very toxic chemicals. By putting the solid catalyst in close contact with the solid reactant in this gas-solid reaction system, the solid-solid exchange reaction was greatly enhanced. This solid-solid reaction, which is thermodynamically favored over the decomposition reaction of the catalystchloride complex, thus eliminated any HCl formation. This, in turn, made the operation possible at low temperature, eliminating the need for acid- and hightemperature-resistant construction. The countercurrent moving bed removes the salt made from the sodium carbonate and organic chlorides and acts as a dry scrubber inside the reactor. Therefore, no separate scrubber is needed, saving the investment and operating costs for additional corrosion-resistant equipment. Investment can be as low as 30-50% of the best competing technology. The cost of the catalytically activated sodium carbonate can be as low as $0.50/lb in bulk quantities; therefore, operating costs will also be low, especially where low-concentration pollutants have to be destroyed to high conversion. Acknowledgment The contributions of Dr. H. G. Stenger, Jr., Dr. G. E. Buzan, and Mrs. K. Hu at the Chemical Engineering Department of Lehigh University in Bethlehem, PA, and Dr. J. Tavakoli and Ms. B. Kline at the Chemical Engineering Department of Lafayette College in Easton, PA, are recognized. Thanks are due to Dr. A. Balogh, Mr. D. O. Erdene, and Mrs. M. K. Hima´n of the former Hungarian Oil and Gas Research Institute in Veszpre´m, Hungary, for performing the pilot-plant experiments. The contribution of Ms. Gail B. C. Marsella to edit this paper and make it printer ready is greatly appreciated. Appreciation is also expressed to Ben Franklin Foundation and to EPA for grants that made some part of this work possible. Notation C ) concentration of pollutant (mol/m3) d ) diameter of reacting shrinking core (m) dp ) catalyst (solid) particle diameter (m) F ) total molar flow of gas or air (mol/s) FS ) molar flow of solid (as Na2CO3) (mol/s) Fy ) molar flow of pollutant (mol/s) f ) fraction of NaCl formed GHSV ) gaseous hourly space velocity (h-1) K ) constant defined in text k ) kinetic constant (s-1 or m/s) k′ ) kinetic constant defined in text (mol/(m3 s)) r ) reaction velocity (mol/(m3 s)) S ) surface of the shrinking core (m2) V ) catalyst particle volume (m3) Vr ) reactor volume (m3) X ) fractional conversion of solid y ) mole fraction of pollutant Greek Symbols Ry ) stoichiometric coefficient of pollutant RX ) stoichiometric coefficient of solid ) void fraction of catalyst bed FS ) mole of soda per reactor volume (mol/m3)
Ind. Eng. Chem. Res., Vol. 36, No. 3, 1997 521
Appendix A. Calculation Kinetic Constants
Appendix B. Estimation of Thiele Modulus
Basis: Lafayette experiments of 7/15/93, Figure 2, ignition curve, reactor 2, MMG catalyst 3, 57.1 ppm TCE, GHSV of 10 800 h-1.
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Received for review October 15, 1996 Revised manuscript received December 10, 1996 Accepted December 12, 1996X IE9606552 X Abstract published in Advance ACS Abstracts, February 1, 1997.
