Ind. Eng. Chem. Process Des. Dev. 1002, 21, 721-725
721
d, = particle diameter, m d, = diameter of copper tube making up the grate, m d, = aperture of wire screen made of iron, m F = force on a particle, N H = magnetic field intensity, A/m I = current passing through the copper tubes of the grate, A Io = holding current, that which will just keep the flow channel blocked, A 1 = length of copper tube conductor, m P = power loss in blocking or controlling the flow of solids, W r = center to center distance between the copper tube con-
more efficient in that the holding current is reduced. (b) Iron screens placed below the copper grate reduce the holding current; the smaller the screen aperture the more effective the screen (Figure 6). (c) To stop or control the flow of solids in a 6-in. vertical flow tube, or to serve as a distributor plate for a fluidized bed of iron particles, a reasonable design consisting of copper grate and iron screen requires about 50-100 A. Power dissipation is calculated to be about 30 W. (d) Various alternative geometries will work just about as well as the copper grate plus iron screen arrangement. (e) The attractive feature of the MVS is that it has no moving parts, no mechanical action, and has rapid response. (f) These valves can be the key to the development of a new class of countercurrent gasaolid contactom through which the solids pass in close to plug flow at a controlled rate, as shown in Figure 10. This may be the way of using rapidly deactivating friable catalysts. Acknowledgment This research was done as part of a grant from NSF, No. CPE-8026799. E.J.M. is grateful for a postdoctoral grant received from the Comite Conjunto Hispano Norteamericano para la Cooperacion Cientifica y Tecnologica (Spain). Nomenclature d, = center to center spacing of the grate, m
ductor and a particle, m R = resistance of copper tube conductor, fl S = cross sectional area of copper in the copper tube conductor, m2 V , = volume of particle, m3 Greek Letters p = electrical resistance of the copper tube conductor, fl-m p = effective permeability of packed bed, H/m Literature Cited Fitzgerald, T. J.; Levenspiel, 0. U.S. Patent Applied for, 1982. Gilbert, N. E. "Electricity and Magnetism"; Macmillan: N e w York, 1941; p 210. Jackson, J. D. "Classical Electrodynamics";Wiiey: N e w York, 1962; p 176. Kunii, D.; Levenspiel, 0. "Fluidization Engineering";Wiiey: New York, 1969; p 370.
Received for review November 17, 1981 Accepted April 27, 1982
Fluosilicic Acid Extraction by High-Molecular-Weight Amines Patrlcla A. Tlvnan' and Rlchard J. McCluskey" Department of Chemlcai Engineering, Clarkson College of Technology, Potsdam, New York 13676
The extractlon of fluoslilclc acid from dilute aqueous solutions by high-molecular-weight amines in kerosene has been studied. Extraction equlllbrium curves for various commercial amines are presented for room temperature, an amine concentration of 0.2 M, and aqueous acid concentrations up to 1 M. Formation of stable suspensions, coextraction of water, and aggregation of amine salts are discussed. Amine extraction of fluosilicic acid is compared
with that of other mineral acids.
Introduction Phosphate rock typically contains 3-4% fluorine. In the United States, phosphoric acid, a basic raw material of the fertilizer industry, is produced mostly by the wet processing of phosphate rock, wherein the crushed rock is contacted with sulfuric acid. This treatment causes most of the fluorine to convert to fluosilicic acid, H a s B . Some of this acid fluoride is recovered, but much is discarded. Fluosilicic acid and ita salta are valuable byproducts. They find use in aluminum smelting and water fluoridation. Furthermore, large potential applications exist in the production of high-purity silicon crystals for solar cells and in reducing the importation of fluorine in the form of fluorspar. There is environmental concern over the effects of discharged fluoride on soil fertility, and the concern over fugitive emissions grows as the phosphate mining industry IBM, Essex Junction, VT 05452. 0 198-43051821112 1-072 180 1.2510
expands (Liner0 and Baker, 1978). A 1977 EPA study estimated that the phosphate fertilizer industry was responsible for 14% of all domestic soluble fluoride emissions (EPA, 1977). Thus, there is incentive to find economical means of recovering fluosilicic acid from waste water streams from the viewpoints of both environmental protection and increased yield of a marketable product. Some of the methods that have been proposed for recovery of the fluoride include precipitation as calcium fluoride (Blake and Stickney, 1978; O'Neill, 1980),conversion to concentrated hydrofluoric acid (Becker and Weiss, 1972), precipitation as sodium aluminum hexafluoride (Kriihnan and Mahalingam, 1973),and collection by an anion-exchange resin (Rozycka, 1978). This paper examines the feasibility of extracting fluosilicic acid from dilute waste water streams by contact with high-molecular-weight amines dissolved in kerosene. Some advantages of acid recovery via solvent extraction rather than by contact with a solid anion-exchange resin 0 1982 American Chemical Society
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Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982
Table I. Amine Extractants name
type
composition
mean mol w t
manufacturer
Primene JMT Amberlite LA-2 Alamine 336 Aliauat 336
primary secondary tertiary quaternary
trialk ylmethy lam ines n-lauryl-n-trialkylmethylamines tri-n-(octyl + decy1)amines methyltricaprylylammonium chloride
310 374 392 475
Rohm and Haas Co. Rohm and Haas Co. General Mills Chemical General Mills Chemical
or adsorbent are the ability to use continuous operation, the ability to recover the acid in a more concentrated form, and the avoidance of solids handling (Kunin, 1973). In addition, solvent extraction requires much less energy than acid concentration by distillation or evaporation. The ability of long-chain amines to extract acids from aqueous solutions has long been known. The behavior of the amine salt so formed has been likened to that of solid anion-exchange resins, earning amine extractants the name liquid ion exchangers. Considerable research has been done on the use of these amines to extract anion complexes of radioactive metals encountered in nuclear fuel cycles (Keder and Wilson, 1963) and of other valuable metals (Cattrd and Slater, 1973). An excellent summary of early research on the extraction of organic and strong acids by high-molecular-weight amines is given by Marcus and Kertes (1969). More recent investigators extracted nitric, hydrochloric, and hydrofluoric acids using LA-2 amine dissolved in kerosene (Russo et al., 1975), examined the effects of nitric acid and uranium on the extraction of hydrofluoricacid by primary amines (Beranova and Kuca, 1975), extracted phosphoric acid with tertiary amines (Marcus et al. 1978), and extracted thiocyanic acid using various high-molecular-weight extractants including amines (DeJong and Brinkman, 1978). Process development work has been published on recovery of nitrate ions by secondary amines (Mattila and Lehto, 1977) and on deacidification of glyoxal using tertiary amines (McCain and Theiling, 1980). The chemistry of amine extraction systems is very complex. The extent of extraction depends upon the nature of the organic solvent or diluent as well as upon the type of amine (Wilson and Wogman, 1962). The higher the dielectric constant of the diluent, the better is the extraction of acid. Nevertheless, cost and safety considerations usually dictate the use of kerosene as diluent in commercial processes. It has also been shown that amines and their salts act as surfactants and are concentrated by adsorption at the aqueous-organic interface (Pizzichini et al., 1977). The electrostatic repulsion of the adsorbed ionic species often leads to very stable suspensions. There are other complications encountered in amine extraction systems. Amine salts will aggregate in the organic phase, sometimes causing formation of a third phase (Verstegen, 1965). The new third phase is an organic phase rich in dissolved amine salts and water. Its density is intermediate to that of the diluent and aqueous phases. Extraction of excess acid may also occur (Hogfeldt, 1966). This is extraction of more than one mole of acid per mole of amine. It is most prevalent with weak acids and at high acid concentrations, i.e., aqueous concentrations above 4 M. Finally, considerable amounts of water may be coextracted with the acid. The general trend is for coextraction of water to increase with increasing organic phase amine salt concentration (Mullen and Diamond, 1966). It has been proposed that water extraction occurs by formation of a hydrogen bond with an anion of the amine salt. This could lead to aggregates of amine salts and water formed into six-membered rings (Mrnka et al., 1974). Previous work on amine extraction of fluosilicic acid is limited to studies using trioctylamine (TOA) in chloroform
or xylene to remove silicon anionic complexes from hydrofluoric acid solutions. It was concluded from analysis of the fluoride and silicon content of the extracted compound that the extracted salt was (TOAH+)2SiF2-(Palshin et al., 1971). The aim of the present study was to obtain equilibrium data on extraction of fluosilicic acid from dilute, less than 1 M, aqueous solutions using kerosene as diluent and commercially available amines as extractants. Information was also gathered on conditions in which a third phase or a stable suspension is formed and on the coextraction of water. Experimental Section Shaker tube tests were performed using 250-cm3polyethylene separatory funnels. Typical charges were 100 cm3 of acid solution and an equal volume of an amine in kerosene, organic phase. The funnels were shaken vigorously by hand for a few minutes and allowed to stand overnight. Samples of the clear aqueous phase were analyzed for total fluoride concentration using the method of Jordan (1970). In this procedure, concentrated ammonia was added to convert H2SiF6to NHIF and Si02. Then Orion TISAB I11 buffer solution was added to control pH to between 5 and 5.5, to complex with interfering ions, and to provide a high, constant background ionic strength. Next, an Orion 94-09 fluoride ion specific electrode was used to measure the fluoride concentration in solution. The difference in fluoride concentration before and after contact with the organic phase was taken as the extracted fluoride. The fluoride ion specific electrode was calibrated at least weekly against a standard fluoride ion solution obtained from Orion. The water content of the organic phase was determined by a Karl Fischer titration (ASTM Standard Test Method D-1744) using a K-F electrode from Luft Instruments. Viscosities of some organic phase samples were measured using an Ostwald viscometer from J. G. Schott Co. (Tivnan, 1980). Reagent grade fluosilicic acid was obtained from Fisher Scientific Co. (A-150) or Alfa Products (69109),Danvers, MA. The commercial amines are described in Table I and were used as supplied by their manufacturers. Kerosene was obtained from a local vendor. For the Karl Fischer titrations, a Fisher Scientific Co. Karl Fischer reagent and diluent were employed. A 1:3 mixture of Baker Photorex Reagent methanol and Baker Analyzed Reagent chloroform was used as a solvent. The reagent was calibrated using a Fisher Scientific Co. water-in-methanol standard solution. All experiments were performed at room temperature (20-25 OC). Results and Discussion Equilibrium extraction data are presented in Figure 1 as a plot of the logarithm of mole ratio of acid to amine in the organic phase vs. the logarithm of the molarity of aqueous acid. The logarithmic scales were used solely to permit display of the data over a broad range of magnitude. Data are presented for all of the extractants at an amine concentration of 0.2 M (10 vol %) and a volumetric phase ratio of 1:l. The primary, secondary, and tertiary amines have equilibrium curves that are qualitatively similar,
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982 723
Table 11. Experimental Results
3.6 9.6 8.6 8.2 5.6 6.5 7.8 7.5
0.5 0.1 0.2 0.4 0.2 0.4 0.2 0.4
LA-2 Alamine Aliquat
X
n n
I
-2
Log
10 10 0.5 0.5 0.5 0.9 to 8 1 to 7 9 to 108 26 t o 275
3.1 0.15 0.9 1.2
creasing aqueous concentration of acid than found for any of the other extractants. At acid concentrations in excess of 0.01 M, Aliquat 336 does not extract as much acid as the other amines. However, it does give fair extraction from more dilute solutions. If it is assumed that the extracted species is ( L ~ I - I + ) ~ S an ~ F apparent ~~-, equilibrium constant may be defined by ((AH+)zS~F~~-)~~~ (1) Kapp = (A)2,,,(H+)2aq(SiF62-)aq
/
-3
7 17 37 51 57 4 7 13 13
(H$I%)
I
aq
Figure 1. Extraction equilibrium curves for 1:l volumetric phase ratio and 0.2 M organic phase amine concentrations: (0) Primene JMT; (0)Amberlite LA-2; (X) Alamine 336; (A)Aliquat 336.
rising steeply at low aqueous acid concentrations and reaching a limiting ratio of extracted acid to amine of approximately 0.5. The strong curvature of these plots is typical of amine extraction systems. Each of the amines is fully loaded at aqueous acid concentrations above 0.1 M. A surprising feature is the poor extraction performance of Primene relative to that of LA-2. For most mineral acids, primary amines are better extractants than secondary amines (DeJong and Brinkman, 1978; Frolov and Sergievskiy, 1975). High-molecular-weight amines function as extractants of mineral acids because the polar amine group is preferentially aligned in the aqueous phase, while the amine's long organic tail keeps it soluble only in the organic phase. Since the nitrogen atom bears an unshared electron pair, it is easily protonated in an acidic medium. The protonated amine has a positive charge and can form a salt complex with anions from the aqueous phase. The maximum extracted acid to amine ratio of 0.5 suggests that the ion associate, (AH')$iFe2- is formed, where A represents a primary, secondary, or tertiary amine. This is consistent with the earlier work of Palshin et al. (1971)with trioctylamine. In view of the above, it is all the more surprising that Primene showed poorer extraction from dilute solutions than LA-2. The primary amine, having only one alkyl side chain, should be much more mobile and hence more readily brought into a (AH+).$iFe2- complex. The poorer extraction of the tertiary amine, Alamine 336, relative to LA-2 can be rationalized as due to the hindrance of three side chains. Aliquat 336, the quaternary ammonium salt, extracts the siF6" anion by an ion-exchangemechanism. The different mechanism for extraction, coupled with the greater size of the quaternary complex, accounts for the different shape of its extraction equilibrium plot. There is a much more gradual increase in organic acid to amine ratio with in-
where ( )erg denotes organic phase concentrations and ( Iaq implies aqueous concentrations. This apparent equilibrium constant is not a true equilibrium constant since concentrations are employed rather than activities and any aggregation in the organic phase is disregarded. Kappis still a measure of how well a given amine functions as an extractant of dilute acid. Another measure of extraction capability is the distribution coefficient, D, which is the ratio of organic phase to aqueous phase acid concentration. In Table 11, log Kappand the maximum value of the distribution coefficient that was actually measured are given for a number of concentrations of the various extractants. The apparent equilibrium constants of Primene, LA-2, and Aliquat 336 decrease with increasing amine concentration,while that of Alamine 336 shows an increase. Despite the reduction in KWp,aJl of the amines gave higher valuea for the distribution coefficient as their organic phase concentration increased. Even though the extracted acid per amine may decrease with higher amine concentration, the degree of separation is enhanced. The anomalous behavior of Alamine 336 may be due to greater association in the organic phase. During the shaker tube experiments stable suspensions frequently formed. Often these would not disappear from the organic-aqueous interface after 24 h of settling. The Primene extractant had the least tendency to form these suspensions. Stable suspensions occurred using Primene only at high amine concentrations (0.5M) and in only two of seven experiments. Stable suspensions resulted with LA-2 at all amine concentrations but were most frequent at the highest amine loading. Both Alamine and Aliquat caused stable suspensions in almost every experiment. There was no correlation between aqueous acid concentration and formation of a stable suspension. The above results show that care must be taken in the design and operation of liquid-liquid contact equipment to minimize disruption of the interface. The appearance of a third phase never occurred with either Primene or LA-2. However, a distinct third phase formed between the aqueous and clear organic phases whenever Alamine or Aliquat was used in experimentswith an aqueous acid concentration above lo9 M. The results of the Karl Fischer tests are shown in Figure 2 for LA-2 extractions and in Figure 3 for extractions with Primene. Both show the water concentration in the or-
724
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982
0 IO-/
/
,.'
lH2S'Fsl or(
Figure 2. Coextraded water vs. extracted acid for Amberlite LA-2: (a)0.1 M L A 4 (0)0.2 M LA-2; (A)0.4 M LA-2.
m--c---o1
I
I
1
I
1 / 05
20
15
IO
( H StF
2
6
25
30
I orp
Figure 3. Coextracted water vs. extracted acid for Primene J M T (a)0.2 M Primene; (A) 0.5 M Primene.
ganic phase plotted vs. organic amine salt concentration. Each plot shows increased extraction of water with increasing amine salt concentration. Even with considerable scatter in the data, the relationship between organic phase water and amine salt concentrations appears linear and independent of the total amine concentration. Linear least-square fits to the data are shown in each figure. In Figure 2 the slope is one-half, implying one mole of water is coextracted with every two moles of amine salt in LA-2 extraction. In sharp contrast, the slope of the least squares fit in Figure 3 is almost 10. The size and complexity of Primenefluosilicic acidwater species in kerosene may explain why Primene is a poorer extractant than LA-2 for dilute acida, but it is unclear why so many water molecules associate with the fluosilicic acid salt of the primary amine. The large size of the SiF% anion would not favor a tight solvation shell, and the amount of water coextracted by the secondary amine is not unusual. Following the reasoning of Mrnka et al. (1974),it is possible for the salts of the primary amine to associate with water and form complexes within the organic phase, though in order to account for the large amount of coextraded water, the structure of these complexes would need to be much larger than six-membered rings. Extraction with Alamine or Aliquat gave no straightforward relationship between the concentrations of extracted acid and water. With these extractants, the water content of the third phase was typically a factor of 10
d
I
1
I
I
I
3
-2
1
0
I
L o g (Acidlaq
Figure 5. Extraction equilibrium curves for 1:2, organic:aqueous, volumetric p h ratio and 0.2 M organic phase amine concentration: ( 0 )HNO,; (0) HC1; (v)HF (all from Russo et al., 1975);and ( 0 ) HPSiFB(this study).
higher than in the less dense organic phase. The water coextraction results are summarized in Table 11, which lists for each of the extractants the ratio of extracted water to extracted acid and the maximum observed molarity of water in the organic phase. The maximum water concentrations occurred at the highest levels of acid extraction. They show that Primene had the greatest tendency to coextract water, and LA-2 had the least. The very high value of the extracted water to extracted acid ratio with Aliquat reflects the ability of the quaternary ammonium salt to extract water by itself. Viscosity measurements were made on a series of organic phase samples following extraction with LA-2. Results are plotted in Figure 4 as viscosity vs. the organic phase concentration ratio of acid to amine. Data for amine concentrations of 0.4 and 0.1 M are presented. At the lower amine concentration, the viscosity is essentially independent of the extent of acid extraction. However, at an LA-2 concentration of 0.4 M, the viscosity of the organic phase rises sharply as the amount of acid extraction increases. This is evidence that significant aggregation can occur in the organic phase without causing third-phase formation. The extraction of fluosilicic acid by secondary amines is compared with that of other mineral acids in Figure 5. The logarithm of the organic phase acid to amine con-
Ind. Eng. Chem. Process Des. Dev. 1002, 21, 725-728
centration ratio is plotted vs. the logarithm of the equilibrium aqueous phase acid concentration. All of the data are for room temperature extraction by 0.2 M LA-2 in kerosene using a 1:2 organic to aqueous phase volume ratio. The data on extraction of nitric, hydrochloric, and hydrofluoric acid were taken from Russo et al. (1975). Fluosilicic acid is more readily extracted from dilute solutions than hydrochloric acid, but less readily than nitric acid. The equilibrium plot for fluosilicic acid is qualitatively similar to those for the other strong mineral acids, although the limiting acid to amine ratio a t higher acid concentration is only one-half for fluosilicic acid due to the divalent charge of its anion. Hydrofluoric acid is a relatively weaker acid. It shows less extraction from dilute solutions and gives acid to amine ratios in excess of 1for aqueous concentrations above 0.1 M. Conclusions It has been shown that fluosilicic acid is readily extracted from dilute acid solutions by high-molecular-weight amines. Of the commercial amines examined, Amberlite LA-2, a secondary amine made by the Rohm and Haas Co., gave superior extraction. Distribution coefficients in excess of 50 were obtained. The organic phase water to acid mole ratio was about 0.5 for LA-2 extraction, and of the order of 10 for Primene extraction. Aggregation of amine-acid salts may occur. Use in kerosene of Alamine 336 or Aliquat 336 as extractants led to formation of a second organic phase. When LA-2 in kerosene was used as an extractant, the viscosity of the organic phase was not influenced by acid loading for an amine concentration of 0.1 M, but at an amine concentration of 0.4 M the viscosity significantly increased with greater acid extraction. Creation of stable suspensions can be a problem in fluoeilicic acid-amine in kerosene extraction systems. Such suspensions occurred more often at higher amine concentrations. Primene showed the least tendency to form stable suspensions. Acknowledgment The authors wish to thank the Rohm and Haas Co., Philadelphia, PA, and the General Mills Chemical Co., Minneapolis, MN, for samples of the amine extractants. Acknowledgment is made to the donors of the Petroleum Research Fund, administered by the American Chemical
725
Society, for the support of this research. Nomenclature A = amine D = distribution coefficient, organic phase acid to aqueous phase acid concentration ratio Kapp= apparent equilibrium constant defined by eq 1 ( = aqueous phase concentration, mol/L ( )erg = organic phase concentration, mol/L Literature Cited Becker, W.; Weiss, W. German Patent 2248 149, 1972. Beranova, H.; Kuca. L. Collect. Czech. Chem. Commun. 1975. 40, 3608. Blake, H. E.; Stickney, W. A. "Proceedings of the Mineral Waste Utilization Symposium, 3rd, Chicago"; Mar 1972; p 179-183. Cattrall, R. W.; Slater, S. J. E. Coord. Chem. Rev. 1973, 7 7 , 227. DeJong, 0.J.; Brinkman, U. A. Th. J. Inorg. Nucl. Chem. 1978, 40, 2055. Environmental Protection Agency, Technical Report: Inspection Manual for Enforcement of New Source Performance Standards--phosphate Fertllizer Plants, Washington, DC, 1977. Froiov. Yu. G.; Sergievskiy, V. V. Tr. M s k . Khim. Tekhnol. Inst. 1975, 89, 55-60. HGgfeldt, E. in "Ion Exchange", Marinsky, J. A. Ed.; Marcel Dekker, Inc: New York, 1966 Vol. 1, pp 139-171. Jordan, D. E. J. Assoc. Off. Anal. Chem. 1970, 53, 447. Kedar. W. E.; Wilson, A. S. Nucl. Sei. Eng. 1963, 77, 289. Krishnan, R.; Mahalingam, M. S. Feffillzer News, 1973, 78, 31. Kunin, R. In "Ion Exchange and Solvent Extraction", Marinsky. J. A., Ed.; Marcel Dekker, Inc.: New York, 1973; Val. 4. Linero, A. A,; Baker, R. A. EPA Report No. EPA 600/2-78-124, June 1978. Marcus, Y.; Asher, L. E.; Berak, H. J. Inorg. Nucl. Chem. 1976, 40, 325. Marcus, Y.; Kertes. A. S. "Ion Exchange and Solvent Extraction of Metal Complexes"; Wiley-Interscience: New York, 1989; Chapter 7. Mattila, T. K.; Lehto, T. K. Ind. Eng. Chem. Process Des. Dev. 1977, 76, 469. McCain, J. H.; Theiiing, L. F. I d . Eng. Chem. Process Des. Dev. 1980, 79, 494. Mrnka, M.; Satrova. J.; Jedlnakoua, V.; Calada, J. "Proceedings of the International Solvent Extraction Conference, Lyon, 1974, published by Society of Chemical Industry, London, 1974, pp 49-62. Muller, W.; Diamond, R. M. J. Phys. Chem. 1966, 70, 3469. O'Neiii, P. S. Ind. Eng. Chem. Rod. Res. Dev. 1980, 19, 250. Paishln. E. S.; Ivanova, L. A.; Davydov, A. V. Russ. J. Anal. Chem. 1971, 26, 1406. Pizzlchinl, M.; Chlarizia, R.; Danesl, P. R. J. Inorg. Nuel. Chem. 1977, 39, 519. Rozycka, D. Przem. Chem. 1978, 57, 199; Chem. Abstr. 89, 1978, 113385. Russo, U.; Zanin, S.; Cescon, P. Ann. Chim. 1975, 65, 647. Tivnan, P. A. M.S. Thesis, Clarkson College of Technology, Potsdam, NY, 1980. Verstegen, J. M. P. J. J. Inofg. Nucl. Chem. 1965, 27, 201. Wilson, A. S.; Wogman, N. A. J. Phys. Chem. 1962, 66, 1552.
Receiued for review November 30, 1981 Accepted May 10, 1982
This work waa presented at the 2nd World Congress of Chemical Engineering, Montreal, Oct 1981.
Thermodynamics and Kinetics of the Coordination of NO to FeIINTA in Aqueous Solutions Nenhua Lin, David Littlejohn, and Shlh-Ger Chang' Lawrence Berkeley Laboratory, University of California. Berkeley, Calfornia 94720
The enthalpy, entropy, and equilibrium constant of the reversbble binding of NO to FeI'NTA in an aqueous solution have been determined to be AHo = -11.9 kcal/mol, AS = -11.1 eu, and K = 4.6 X lo5 M-' at 50 OC, respectively. The formation and dissociation rate constants of Fe"(NTAXN0) are k 1 7 X 10' M-l s-l a nd k - , 2 35 s-', respectively, at 25 OC. The rate of binding of NO to form the complex is much faster than the mass transfer rate at typical wet scrubber conditions.
Introduction One approach for simultaneous removal of NO, and SO2 from power plant flue gas is based on absorption of the 0198-4305/82/1121-0725$01.25/0
relatively insoluble NO in aqueous solutions after oxidation of NO by O3 to NO2 or by addition of metal chelates in solutions (Yaverbaum, 1979). These metal chelates have 0 1982 American Chemical Society
