Concentration of Radioactive eous Wastes ELECTROMIGRATION THROUGH ION-EXCHANGE MEMBRANE§ W. R, WALTERS, D. W. WEISER', AND L. J. MAREK Argonne National Laboratory, Lemont, Ill.
Based on experimental data an electrolytic process incorporating permselective membranes and granular ion exchangers for treatment of radioactive aqueous wastes is proposed. The first step in the process involves partial decontamination in a multicompartment membrane cell of the same general type proposed for producing potable water from saline waters, followed by final decontamination in a multicompartment permselective cell containing a mixed bed granular exchanger in the deionization compartments. For low solids wastes power cost is less than for vapor compression distillation and volume reduction is higher than for chemically regenerated ion exchangers.
S
UCCESSFUL treatment of a radioactive waste usually re-
quires that the dissolved solids be removed almost completely from the water and concentrated in as small a volume as possible for storage. There are some types of wastes in which activity can be selectively removed from the inactive solids. But, in the main, radioactive waste treatment should be considered as a concentration process producing a high quality water substantially free of activity that can be discarded or re-used, aa desired. The only completely reliable unit operation available to date for concentration of activity has been evaporation. A large portion of the wastes currently being treated by evaporation are solutions with 0.1% dissolved solids or less. Although the bulk of these low solids wastes can be successfully decontaminated by ion exchange, the volume concentration factor resulting from chemical regeneration of the resin, plus rinses, is usually inadequate. Industrial water treatment in electrolytic cells employing uncharged diaphragms has been investigated by Billiter (1) in a three-compartment cell1 and by Streicher, Bowers, and Briggs (11) in two- and three-compartment cells. Streicher, Bowers, and Briggs (11) have also developed a scheme whereby the anolyte and catholyte wastes from a three-compartment cell are utilized to regenerate ion exchange beds for further purification of the partially demineralized effluent water from the middle compartment. Such schemes although applicable t o industrial water purification do not effect the appreciable volume reduction which is necessary for radioactive waste treatment. Relatively recent commercial availability (16-19) of ion exchangr (permselective) membranes has spurred interest in elect,rolytic deionization methods for handling low solids wastes. These membranes permit high volume reduction not attainable with uncharged diaphragms. Production of potable water from sea water or brackish water by means of a multicompartment permRelective membrane cell has been discussed by Langelier (L)and Wiechrre and Van Hoek (14). The requirements lor radioactive waste purification, however, are more exacting. For the production of potable water it is only necessary to reduce the total solids to about 300 p.p.m. To decontaminate a n active waste sufficimtly it iF necessary to remove, almost completely, the inactive ions aq well. In a n electrolytic deionization process power rrquirements per 1 Present address, Natural Sciences Staff, the College of the University of Chicago, Chicago, 111.
January 1955
unit of solids removed rise markedly a s the water becomes depleted in solids. This is due t o the rapid increase in specific resistance of the solution as well as the decreasein current efficiency for salt removal. This extreme sensitivity of specific resistance to aalt content in the low solids region also imposes some severe restrictions on flow distribution in a multicompartment cell. Obvious means of attacking these problems are to resort t o extremely small spacing between membranes and t o employ very low current densities. These two methods are limited by hydraulic considerations and equipment cost, respectively. Another approach has bern taken t o this problem, namely, the introduction of a porous ion conduct,ing medium between the membranes of the deionization chambers. The ion conduction medium chosen was a mixed polyelectrolyte-in this case a mixture of strong base and strong acid ion exchange granules. The process proposed in this paper is
1. Partial deionization in a multicompartment permselective membrane cell. This step is similar t o the production of potable water from sea water or brackish water. 2. Final deionization of the partially demineralized effluent from the first step in a multicompartment permselective membrane cell employing mixed anion exchange and cation exchange granules in the deionization compartments. ELEJIESTS O F ELECTROLYTIC DEIOSIZATIOY W I l I I PEHJISELECrIYE MEJIBRASES
Membrane Function. A quantitative treatment of the hehavior of select ix-e meml)ranc~\vas proposed by Tcorc4l ( I ? ) and 1 I q w and Siel-ers ( 5 ) in 1935. Other theories or modifimtions of the initial treatment have since appwred (6,7, 13). -1 detailed cxposition on permselective nicmlmne theory is not essential. The three basic fcntrires which distinguish pcrinselective mcmbrancs are: 1. Selective ionic transfer-either cationic or anionic 2. Relatively high electrolytic conductivity-even tremely dilute solution 3. High hydraulic resistance
in ex-
The first two propertics affect the current and voltage rcquirements and thus the power required for R given trsnsfrr. The third property allows high concentration differcnccs across a incmbrane. One may visualize an ion exvhange membrane R E II polymeric sheet with positive or negative groups bound in the structure
INDUSTRIAL AND ENGINEERING CHEMISTRY
61
which are neutralized by ions of opposite charge when the membrane is immersed in electrolyte. These ions of sign opposite to the fixed charges are the exchangeable ions and arc relatively free to move, The concentration of exchangeable ions is determined by the concentration of the fixed charges plus the concentration of diffused salt, which is a function of the concentration of the external solution. I n very dilute solutio~isand with high concentration of exchange groups the membrane apploaches the ideal of an immobile cation or anion permitting passage of ions of one sign only. OXYGEN O R
DEIONIZED EFFLUENT
ONCENTRATED FEED SOLUTION
Figure 1. Multicompartment Electrolytic Cell Model for Concentration of Salts by Electromigration through Permselective Membranes Concentrated effluent flow rate governed by eleotroosmosis through membranes a = anion selective membrane c cation selective membrane
Permsclectivity as calculated from the measured potential across a membrane separating two solutions of unequal concent,ration may serve as a guide to the expected current efficiency in an electrolytic cell employing membranes. However, when appreciable voltages are applied the resulting current efficiency R-ill usually he less than indicated by membrane permselectivities as measured in concentration cells. This effect is due to the increased leakage of cations t'hrough an anion exchange membrane and anions through a cation exchange membrane; lealiage is particularly pronounced for the hydrogen and hydroxyl ions. Current efficiency may also be loTyer than indicated from permselectivity measurements due to short circuiting through solution flow lines. This is an equipment problem and the effect may be minimized by suitable cell design. In practice it may be necessary to sacrifice some current to simplicity and compactnesg of the equipment. The experimental cells desc,ribed in this paper utilized long solution Bow lines of small cross-sectional area t,o eliminate this effect. Membrane conductivity is dependent on concentration of exchange sites. Consequently membrane conductivity has a nearly constant value in very dilute solutions. As solution concentration increases, membrane conductivity also increases but a t the expense of permselectivity. Although the hydraulic resjstance of t.he membranes is high, water transfer does occur due to electroosmosis. Fev,- experimental data are available for synthetic ion exchange membranes showing the effect of concentration and current density on electroosmosis. The membranes can be considered as rather efficient barriers through which selective migration of anions or cations is allowed. For a given membrane the selectivity and conductivity will be functions of the concentrations on both sides of the membrane as well as the species of ions in t,he system, and of the cell used and its method of operation. Deionization Cell. The simplest type of deionization cell is a three-compartment cell with nonselective membranes or diaphragms. Here the diaphragms serve only t o minimize mixing among the three zones. If inert electrodes are employed, oxygen or a halogen gas will be formed a t the anode, hydrogen gas will be evolved a t the cathode, and a n acidic anolyte and a basic catholyte will result. Deionization occurs in the middle compartment. Such a cell was used for the production of high 62
qualizy water from t a p water a t a power cost less than that r(1 quired for evaporation and was described by Billiter in 1931 (1). Owing to the tendency for migration of ions from the anolyte and catholyte to the middle compartment, successful operation required that the electrode compartments be diluted continuouslx with water. The v,aaste volume amounted to 10 to 40% ol the water fed t o the cell By Pmplos ing permFelective membranes rathei tliari diaphragms containing no exchange sites several advantagcs ai e obtained. The Billiter cell depends upon the reaction H + OH--.HOH occurring in the middle compartment for its continuing operation. When an anion selective membrane is substituted for the diaphragm next to the anode and a cation selective membrane is substituted for the diaphragm adjacent to the cathode, this reaction is not eweritial and is indeed discouraged by the selectivity of the niembrancs. Thus ion transport is essentially by salt anions migrating towaid the anode and salt c-stions migrating toward the cathode with the net result that cutrent efficiency is much improved. Because of the permselectivity of the membranes, it no longer becomes necessary to maintain a n acidic anolyte and a basic catholyte. This is an important point since it makes possible the presence of a concentrated salt solution on both sides of the deionization chamhrr; thus a number of compartments effecting depletion and coiicentration of salt in alternate chamber8 may be interspored bctneen a eingle set of electroder. The advantages of a multicompai tment cell over a series of thi ee-compai tment ceIls are
+
1. Electrode cost becomes a small item. 2. The percentage of total applied voltage lost becautle of polarization a t the electrodes becomes negligible. 3. For a given cell cross section almost one third the number of compartments is eliminated. 4. Objectionable electrode reactions are minimized
I o n exchange beads
\
02 Anion exchange membrane
+l
I
I
Cation exchange membrane
i'
-
7-
I
An illustrotive example of a multicompartment cell of the type contemplated is shown in Figure 1. I n order to concentrate the salt from the incoming waste stream in as small a volume as possible, one relies on the water transfer through the membranes without further dilution of the concentrate.
INDUSTRIAL AND ENGINEERING CHEMISTRY
Vol. 47, No. 1
Ion Exchange Basic Equations. The voltage required for a multicompartment cell is obtained by summing solution and membrane voltage drops
E
a
nI
=
[(&d
+
( d ) c
+ 2(al)ml +
(1)
e
where n is the number of deionization compartments per cell, I is the cell current in amperes, E is the applied direct current potential in volts, a is the effective cell cross-section in sq. em., ‘p is the specific resistance in ohm-em., I is the length in em. of a chamber or membrane in the direction of current passage, and E is the sum of the counter electromotive force occurring a t the membranes and electrodes. The subscripts d, c, and m refer t o deionization comparhment, concent,ration compartment, and membrane, respectively. Obviously specific resistance will vary with concentration and so the above terms for specific resistance are average values for a given system. An average value for membrane resistance, both anion and cation, is used. .4 tacit assumption inherent in Equation 1 is t h a t there are a large number of compartments so t h a t the number of deionization and concentration compartment8 are almost equal. Current requirements for a multicompartment cell can be calculated from either the concentrate or dilute side of the membranes by using a combination of Faraday’s law and a material balance
where F is the faraday constant of 96,500 coulombs per equivalent, q is the average coulomb efficiency defined as (equiv. of salt removed) Fcoulombs of charge input to cell/ deionization compartment
(
HOH
+ NaR + 1/zRLS04 H R + R‘OH + l/z(Na2SO& + NaR + 1/2RiS04 NaR + ‘/zRz’SO4 + l/z(NazSO& HOH + H R + R’OH HR + R’OH + (HOH),
l/zNazS04
-f
)
equivalents per liter, V is the concentrated effluent flow rate in liters per second, X iis the influent salt concentration in equivalents per liter, Xz is the deionized effluent salt concentration, and U is the feed rate in liters per second. The over-all volume reduction expressed as per cent of the influent waste volume is (3) ELECTROLYTIC REGENERATION OF MIXED BED ION EXCHANGER BEI’WEEN PERMSELECTIVE MEMBRANES
The electrolytic regeneration of a mixed bed ion exchanger in a cell employing permselective membranes is, as far as these authors know, a new subject; therefore, the mechanism involved and the application t o deionization of dilute solutions are discussed. The electrodialysis of a mixed bed has heretofore been considcred impractical because of exhorbitant electrical cost (3). Electrolytic regeneration in a multicompartment cell employing permselective membranes exhibits promise of being very attractive from an electrical cost standpoint. The specific conductance of various forms of a cation exchanger have been measured by Heymann and O’Donnell ( 2 ) ; ion exchangers are relatively good electrolytic conductors even when measured in highly deionized water. Spiegler and Coryell (9, 10) have electrolyzed a leached granular cation exchanger in direct contact with the electrodes and have determined that the current in this case is carried exclusively by the cations. The possibilities for ion separation and resin regeneration were indicated. The regeneration of a mixed bed ion exchanger between permselective membranes is somewhat different from the model introduced by Spiegler and Coryell; here, in addition t o the simultaneous electrornigration of anions and cations, the hydrogen ion and hydroxy1 ion replacement does not depend on the electrode reactions but is a hydrolysis reaction occurring a t the
(1)
+
+
Y is the steady state salt content of the concentrated effluent in
January 1955
exchange sites. Thus, such a model is applicable to a multicompartment cell where many such beds may be generated for one set of electrodes. If a random mixture of ion exchange granules is placed between a cation selective membrane and an anion selective membrane there should be a conductive path for anions through the anion exchanger and a similar situation will exist for the cations and cation exchanger. If the discussion is limited t o beds in contact with dilute solutions, less than 0.1N, then i t is reasonable t o assume that the bulk of the conduction will occur through the exchanger. For purposes of illustration, a highly simplified model of a mixed bed in contact with a dilute sodium sulfate solution is shown in Figure 2. A three-compartment cell was chosen for simplicity. The anolyte and catholyte r e r e continuously mixed t o eliminate the possibility t h a t hydrogen and hydroxyl ions could be a prime factor in the reactions occurring in the resin compartment, and t o simulate the conditions which would be encountered in a multicompartment cell-i.e., the presence of a concentrated salt solution on either side of the resin compartment. Tbere are many ways in which the cations and anions can migrate through the exchangers, several of which are shown in Figure 2. Regardless of the sequence of exchange reactions, the end result is the same. Bnions and cations are removed simultaneously from the solution by the resin in the usual exchange mechanism and transported through the granules and menibranes under the influence of the electric field. The efficiency of salt removal will be determined by the competing reactions
(2)
(3)
where R denotes the cation exchanger, R’ is the anion exchanger, and the subscript c indicates the combined ion transfer through the membranes. Reaction 1 represents the regeneration of the exchanger due to hydrolysis and is the important reaction occurring for a batch regeneration step where no salt is added t o the bed during electrolysis. Reaction 2 is the important reaction which takes place when the bed is electrolyzed continuously during the addition of salt to the bed Competing with these two reactions is Reaction 3 where self exchange of hydrogen and hydroxyl ions results in the formation of water in the concentrate. Reaction 3 is wasteful of current and can only be minimized by operating with the bed mostly in the salt form. For a batch regeneration cycle this means that efficiency xi11 decrease with the fraction of the bed regenerated, while for a continuous process the ratio of current to flow rate must be minimized. Since the first work was reported, a continuous process has been developed. This work involved separate regeneration of a n exhausted bed follonTed by exhaustion of the regenerated bed with a salt solution. The same reasoning developed for multicompartment membrane cells not employing granular exchangers applies t o a multicompartment cell using the mixed bed. EXPERIMENTAL WORK
I n order to estimate the power requirements for a n assumed cell design by use of Equations 1 and 2 and the resulting volume reduction by Equation 3, it is necessary that the following quantities be experimentally obtained: 1. Average membrane resistance 2. Average coulomb efficiency for salt transfer 3. Average resistance of the deionization compartment 4. Water transfer through the membranes; this quantity determines the steady state salt content in the concentrated effluent according t o the relation: Y = q / O , where 0 is defined a s the liters of water transferred through B membrane pair per faraday
INDUSTRIAL AND ENGINEERING CHEMISTRY
63
The systems selected for experiment were sodium sulfate and laboratory t a p water containing about 300 p.p.m. of dissolved solids. Principal ions in the t a p water were sodium, calcium, magnesium, sulfate, chloride, carbonate and bicarbonate. Preliminary Deionization in Multicompartment Celis. Experiments of a very qualitative nature were first conducted t o find the degree of decontamination from radioactivity in simulated wastes attained with a membranc cell. The assumption wtts made t,hat the inactive constituents would stem largely from laboratory t a p water, which is the principal diluent for radieactivity. It was further assumed that mixed fission products would represent the principal contaminants. Consequently, feed solutions were composed of t a p water with mixed fission products added to about 105 counts per minute per milliliter. waste Feed
Deionized wosfe
I
t
-Ir
I'
I
I
1u 1.
A n o l y f e - COthOlyfe R e c y c l e Sump
Concentrated Waste R e c y c l e Sump
Figure 3. Scheme for Initial Series FIow Experiment of Waste and Concentrate with Small Scale Equipment The equipment used in these initial eyperiments was of simple design. Compartment frames or spacers were l/&ch thick neoprene washers; flow connections were made by inserting */loinch stsinless steel tubes through the peripheries of the neoprene spacers. A ceil of the desired number of compartments was then obtained by sandwiching membranes and spacer between platinum electrodes a t either end. Membranes used in the tests were Nepton CR-41, a cation exchange membrane (16) and Amberplex A-I, an anion exchange membrane (I?').
The results shown in Table I were ohhined in it multicompartment cell using the flow scheme of Figure 3. Although the results were not too reproducible, there is a certain amount (less than 1% of the initial activity) of a radioactive specie8 which is very difficult to trnnsport through the membranes Apparently about 997, of the initial activity is removed a t about the same rate a s the inactive ions Electromigration d o n e viill not yield high removal of activity Therefore, EI final deionization step involving passage of the partially dccontaminated cell effluent through granular ion evchangrrs mvst be considcrad essential for high decontamination. A somewhat larger model conhtructed to study a crll under more precisely controlled conditions and to allow intcrstage sampling without undue effect on operation. Pictures of Lhe components and the assemhl) are shown in Figure 4. In general, its design was similar to the small cell. C'are wap taken, however, to pi ovide a support systcm v hich ~ o u l densure reasonably uniform membrane spacing tlu oiighout the cell. Materials of construction were '/r-inch thic*k neoprene cotnpxtment frames, platinum foil electrodes, and .Imherplex A-I, C-1 permsclective membranes. The flow schenie employed was similar t o that employed for the small cell (Figuic 3) 90 radioactivity was ueecl in these runs. Feed solutions for the ccll were t a p w n k r neutralimd with srilfuiic acid, and sodium sulfatr of about the same specific re&anw For accurate dcterininations of coulomb efficiency sodium sulfate solutions of known concentration were used. Salt concentration in tap water w a p estimated by assuming that the relation between concentration and spccific resistance was similar to sodium sulfate. Justification for t h k assumption was based on similarity in current requirements €or runs made with solutions of identical conductivity for tap water and Yodium sulfate. The dilution curves were also similar. The deionized effluents were acidic and consequently it war necessary to correct for the high conductance of the hydrogen ion when calculating salt concentrations from conductivity and pII measurements. Measurements taken during a run included conductivity and pH of deionized stage effluents, flow rates, applied direct current cell voltage, compartment voltage drops, and current, Prolonged operation with t a p water feeds caused solids accumulation due to precipitation of calcium and magnesium salts in the concentrate compartments, eventually causing membrane failure resulting from decreased effective area and high localized current densities. Another unit, basically similar, was constructed which had larger flow openinga for the concentrate stream and provision for flushing out accumulated solids. An additional feature of this ccll was an electrode scheme incorporat-
T A B L11. ~ SALTTRANSFER, WATER. TRARSBER, A N D MEXIBRANE RESISTANCE AND DECOSTAMINATION IN A MULTITABLE I. DEIONIZATION
COMPARTMENT
Feed:
CELL
T a p water containing 300 p.p m. total solids with mixed fission products added
Influent WaF;te Deioniza- DecontaminaRadioactivity, tione tionb Run Type counts/min./ml. Factor Factor 1 Tap water 3 x 10; 10 20 2 Run 1 effiuent 1 . 5 X 10 6 3.5 3 2 Run 2 effluent 4 . 3 x io' 3 4 110 108 Tap water 2 . 6 x 105 8 R u n 4 effluent 3 . 5 x 10' 13 3 Plus sodium sulfate Tap water 1.1 x 105 170 100 6 7 1.1 x 106 120 100 T a p water 8 Tap water 1.1 x 105 12 39 9 R u n 8 effluent 1 . 0 x 103 120 2 Plus sodium sulfate influent waste salt concentration Deionization factor deionized effluent salt concentration counts/min./ml. in waste feed Decontamination factor = counts/min./ml. in deionized effluent ~
64
Rohm & Haas Amberplex A-1, C-l Av. Effluent Salt Resistance Unit Area Concentration, Equiv./L. of 5 Current Density, Av. Membrane Ma./Sq. ConcenCoulomb Pair, OhmCm. Deionized trated Efficiency Sq. Cm. Sodium Sulfate Feed, 0.078N 0.60 3 I x 10-3 0.3 23.7 0.57 7 0.2 5 X 10-3 27.6 Sodium Sulfate Feed, 0.005~V 2.85 3 . 6 x 10-5 0.2 0.36 150 107 0.36 2 x 10-5 0 . 2 3.03 43 0.37 1 . 6 X 10-6 3 03 0.2 T a p Water Feed 320 3 . 5 x 10-5 0.30 2.91 0.7 295 0.28 4 x 10-5 0.05 3.50 100 0.25 4 x 10-5 0 . 3 5.86 0.2 0.24 70 1 X 10-6 6.90 .. 0.6 0.30 3 X 10-4 3.75 3 X 10-4 0.6 0.24 4.38 .. 0.39 3 x 10-4 0.8 6.67
INDUSTRIAL AND ENGINEERING CHEMISTRY
Av. Water Transfer, L./Faraday
0.55 0.33 1.5 1.1 1.5
..
Vol. 47,No. 1
Ion Exchange ing a lead anode and a stainless stcel cathode with a sodium sulfate solution continuously intercirculated between the anode and cathode compartments. An anion membrane adjacent to the cathode and a cation membrane adjacent t o the anode permitted effective exclusion of the ions in the feed solution from the vicinity of the electrodes. Harmful oxidation and reduction products produced by electrode reactions are thus virtually eliminated and membrane attack should be greatly lessened. Data obtained on these two cells are shown in Table 11. Average coulomb efficiencies for salt transfer varied from 0.23 to 0.60 depending, apparently, almost entirely on the deioniLation range. The highest efficiencies were obtained for sodium sulfate feeds which were deionized from 0 078N to 0.005N and to 0.001N. The lowest efficiencies were noted for the t a p water feeds when deionized from 0.005N t o about 4 X 10-6". T h e production of very high quality effluents can only be accomplished a t the expanse of coulomb efficiency. Average unit area membrane resistance was calculated by subtracting the summation of solution voltage drops and estimated electrode polarization from the applied cell voltage The reported membrane resistance is related to the true resistance
R = r -I-e/i where R is the reported average unit area resistance of a membrane pair, r is the true resistance, e represents the counter electromotive force across the membrane, and i is the current density. It was impossible to maintain perfect membrane Ppacing in :he cell since slight head differences caused bulging; therefore, the reported values should be considercd a s approximate. However, the concentration of the dilute solution and the rurrent density are the controlling factors. Based on these results the average unit area membrane resistance for a membrane pair over a deionization range of 5 X 10-aN to 4 X 10-6N is of the order of 100 ohm-.sq. em. For a deionization range of 0.1N t o 0.001N the value would be in the neighborhood of 10 ohm-sq. cm. Water transfer across the membranes was measured by observing the concentrated effluent flow rate. The electroosmotic water transport across a Kepton CR-51 membrane was estimated by Spiegler (8) to be 0.265 liter per faraday, and has been measured by Rosenberg for very dilute solutions of sodium chloride acrosa a Wepton CR51 membrane in the sodium form and found to be about 0.3 liter per equivalent of sodium ions (8). Since each deionization chamber is composed of two membranes, the water transport for a deionization cell due t o electroosmosis would be about 600 ml. per faraday. Water transfer under a hydrostatic head for Amberplex membranes has been estimated a s (16)
Figure 4.
Larger Scale Eleatrodeionization Cell Assembled Cell Cell Components
matrix, and the r e s h is converted to the hydrogen and hydroxyl forms by hydrolysis. The eluted salt ions are concentrated in the adjacent compartments. Although the salt solution could be fed continuously t o the resin matrix during passage of current, preliminary experiments showed that higher quality effluents were produced by batch operation. Continuous and simultaneous exhaustion and regeneration presented problems of control, and so were abandoned temporarily. The work describing the efficiency of resin regeneration dewrihed here was cyclic. Batch regeneration experiments were conducted in a threecompartment cell, using the equipment shown in Figure 5. T h e results are shorn-n in Table 111. Reported data are averaged
ml. 10-25 ___(hr.)(ft.)Z(atm.) Measured water transfer rates are given in Table 11. The highest transfer rates, 1 to 2 liters per faraday, were observed when deionization t o a high quality effluent was attempted. When the effluent quality was lowered, the water transfer rate decreased markedly. Due t o difficulties encountered in accurately measuring the water transfer rates, no detailed correlation can be justified. Further, steady state conditions were not closely approached in all caBes. Although a water transfer rate a s low as 0.15 liter per faraday is reported, it is felt that 0.4 liter per faraday is a more likely figure. A conservative figure of 0.6 liter per faraday will be used in future calculations for the preliminary deionization step. Electrolytic Regeneration of Mixed Bed Ion Exchanger Using Permselective Membranes. When a mixture of granular cation and anion exchangers is equilibrated with a salt solution and then electrolyzed between a n anion and a cation selective membrane, salt is removed from the mixed exchange January 1955
TABLE111. ELECTROLYTIC REGENERATION OF GRANULAR MIXED BED IONEXCHANQER BETWEEN PERMSELECTIVE MEMBRANES AND CONCENTRATION OF ELUTED SALT 60% Naloite SAR 40% Naloite HCR by volume Rohm & Hkas Amberplex A-1,'C-1 Steady state concentrate simulated with 0.44N sodium sulfate a t start of each experiment Faradays/ Specific Equivalent Coulomb Resistance, Current of Mixed Efficiency Ohm-Cm. for Salt Density Exchange Ma./Sq. dm. Capacity Removal Efeuent Resin Resinx Ehausted with 0.0046N NaiSOa
9.8 11.5 10.5
0.12 0.21
0.38
0.52 0.40
0.37
500,000
770,000 650,000
1500 1200 1400
Resin Exhausted with Tap Water
10.5 10.0 9.2
0.13 0.18 0.33
INDUSTRIAL AND ENGINEERING CHEMISTRY
0.47 0.47 0.37
550,000 750,000 700,000
1400 1400 1800
65
from multiple-cycle operation Effluent water with a composite specific resistance of about 500,000 ohm-cm. was usually obtained during the exhaustion cycle, where exhaustion was stopped once the effluent specific resistance had fallen to 200,000 ohm-cm.
Pump
-
-'
Deionized Effluent
Figure 5.
Electroregeneration and Deionization Apparatus
Cell constants Compartment thicknesses Compartment widths Effective resin depths Effective resin volume Effective membrane area
1.75 om. 2.54 om. 6 inches 6 9 . 3 EO. 39.6 sq. cm.
Membranes used R o h m & Haas Amberplex A-1 and C-1 Method of operation Current on recycle during regeneration (159-260 00.) Current of exhaust resin with partially deionized recirculant plus exhaustant to desired resistivity
The average effluent solution reqistance was about 500,000 ohm-em. while the specific resistance of the resin was 1200 to 1800 ohm-cm. Some tests mere made in which a fixed volume of solution, about three resin volumes, was recirculated to monitor the liquid in contact with the resin during regeneration. I n all cases this solution proved to be considerably lower in salt content than the feed solution. Therefore no precipitation should occur in the resin bed during regeneration I n some experiments the regenerated resm vas transferred to a 1-inch diameter column for exhaustion. This was done because progressive Ion ering of efiluent quality sometimes occurred due t o channeling in the cell compartment. The data obtained from column exhaustion were more reproducible. Accurate data for water transfer across the membranes were not obtained. However, volume concentration factors for 0.005-Vsolutions of a t least one hundred can be attained. The reported results are based on a bed composition of 60y0 anion evchanger by volume. This choice was based on the fact that such a mixture gave effluents with pH's closer t o 7 than were obtained with a few other bed compositions, Since the cation exchanger is the better conductor it might be desirable to use even less anion exchanger. There might also be an argument from a conduction standpoint for using an equal number of beads of the two exchangers since this provides the maximum number of anion-anion and cation-cation contacts. This and other questions such as bead size or behaviour of various types of exchange groups have not a s yet been investigated. 56
h parallel flow of feed solution through the chambers of a multicompartment cell would probably be preferred in practice due to hydraulic considerat'ions. Equation 1 confirms this statement. The requirements of high volume throughput and extremely thin compartments are incompat,ible for a. multicompartment cell. with a series flow pattern. There are, however, certain restrictions to be considered in the use of a parallel flow cell. The deionization range will be limited by the maximum current density which the membranes can withstand; the rurrciit density will vary directly with the solution conductib ity, and thus a series of separate cells will be necessary to effect deionization over a large range. The attainment of nearly equal flous in :ill the deionizing compartments of a parallel flow cell is necessary t o prevent unequal ratee of deionization and consequent high specific resistance in some compartments. Such a situation will cause a disproportionately high cell resistance for a given deionization range. This effect is particularly pronounced in the estremely low solids range where the specific resistance changes tremendously with deionization. For these reasons it is not desirable to attempt complete deionization in a multicompartment electrolytic cell. Rather, a workable process would combine electrolytic deionization Tvith ion exchange for final cleanup. The reason for resorting to electrolytic regeneration of the granular exchanger \%-as to allow high volume concentration factors. Elect'rolytic regeneration of a granular mixed exchanger can be conducted with reasonably good efficiency. The highest regeneration efficiency would be achieved in a continuous process, either by means of a resin st'ream moving continuously between the regenerat'ion cell and a count,ercurrent exhausting contactor, or continuous flow of the exhaustant through the cell resin chambers; in other words the regeneration cell would be operated in precisely the same manner as the c,onventional deionization ceI1 with the resin serving the important function of providing a relatively high and nearly constant conductivity for deionization over an extremely large conductivity range. Slt,hough tBhe la.tter continuous process would seem to be the ideal one from an efficiency and equipment viewpoint, batch regeneration and exhaustion were found to produce the highest quality waters a t the time of the presentat.ion of this paper.
1
I---
L
CONSIDERATIONS IN PROCESS DESIGN FOR ELECTROLYTIC CONCENTRATION O F RADIOACTIVE WASTES
SALT CONTENT OF FEED,
eq./l.
Figure 6. Estimated Energy and Volume Concentration for Two-step Electrolytic Deionization Process I t will be assumed that the preliminary deionization step ran be taken to a specific resistance of about 2,000 ohm-em. n~itliout the parallel flow problem becoming critical. The final deionization will be accomplished by passage through a mixed exchange resin which has been batch regenerated in a multicompartment electrolytic cell. Exhaustion can take place by passing the waste effluent from the preliminary deionization cell either through the resin compartments or to a column filled with resin transported from the cell.
I N D U S T R I A L A N D E N G I N E E R I N G CHEMISTRY
VOI.
47, No. 1
- T o n ESTIiMATED POWER CONSUMPTION AND VOLUME REDUCTION FOR TWO-STEP ELECTROLYTIC DEIONIZATION PROCESS
The process considered combines preliminary deionization to a salt content of 0.005 equivalent per liter, and final demineralization in a mixed exchange resin whirh is electrolytically regenerated. Figure 6 presents the power requirements and volume reduction for a range of salt concentrations as estimated from data for sodium sulfate and t,ap water. ilssumptions upon which this figure is based are
Type process Flow rate of influent waste, gal./hr. Current density, ma./sq. cm. Number deionizing compartments Av. unit area resistance of a membrane pair, ohm-sq. om. Specific resistance of deionization chamber,
ohm-cm.
Preliminary Deionization Step Continuous 1000 10 100 10 Root mean square of influent and effluent
Memhrane spacing of deionization cham0.1 bers, cm. Membrane spacing of concentrate cham0.3 bers, cm. Electrical input/batch of refiin, faradays/ .. equiv. of mixed exchange groups Coulomb efficiency for salt transfer 0.5 Electroosmotic water transfer/deioniza0.5 tion chamber, liters/faraday a Volts required Q Amperes required Q Kv. amp. of d. c. required a Requirements dependent on salt content of influent waste
Final Deionization Step in Resin Cell Batch 1000 10 50 100
2000 0.3
0.3 0.5
0.3 0.6 405 34 14
CONCLUSION
1. The elements for a two-step electrolytic deionization
process have been proposed. The first step, bulk deionization, is generally similar to the technique of producing potable water from sea water or brackish water b y means of a multicompartment permselective membrane cell. A final treatment step was developed which consistR of electrolytically eluting salt ions from a mixed ion exchange bed through permselective membranes; such a scheme may economically produce a high quality effluent and also effect a high volume reduction factor not attainable by conventional chemical regeneration. 2. When oomparing the power costs of vapor compression evaporation with the two-step electrolytic process assumed in
Exchange
this paper, the deionization costs should be cheaper than the evaporation costs for waste feeds containing up t o about 0.4 equivalent per liter of salt. 3. It is doubtful t h a t the concentrated waste from an electrolytic deionization process will have a salt content much in excess of 2N. Therefore] in the case of radioactive waste treatment] a final evaporation step would be desirable t o further reduce storage volumes. 4. More knowledge regarding the behavior in additional chemical systems is essential. In particular, the behavior of radioactive species should be investigated thoroughly. ACKNOWLEDGMENT
The authors are particularly indebted to Alvin Glamner and
W. A. Rodger, of Argonne Kational Laboratory, for valuable criticism on the preparation of the manuscript. LITERATURE CITED
(1) Bi!liter, J., T r a n s . Electrochem. SOC.,60, 217 (1931). (2) Heymann, E., and O'Donnell, I. J., J. Colloid Sci., 4, 395 (1949). (3) Kunin, R., and Myers, R. J., I o n Exchange Resins, p. 109, John Wiley & Sons, New York, 1950. ( 4 ) Langelier, W. F., J . Am. W a t e r W o r k s Assoc., 44, 845 (1952). (5) Meyer, K., and Sievers, J., Helv. Chim. A c t a , 19, 649 (1936). (6) Nagasawa, M., and Kobatake, Y . , J. P h y s . Chem., 56, 1017 (1952). (7) Scatchard, G., J . Am. Chem. SOC.,75, 2883 (1953). (8) Spiegler, K. S., J . Electrocham. SOC.,100, 303c (1953). (9) Spiegler, K. S.,and Coryell, C. D., J . P h y s . Chem., 56, 106 (1952). (IO) Spiegler, K. S.,and Coryell, C. D., Science, 113, 546 (1951). (11) Streicher, L., Bowers, A. E., and Briggs, R. E., IND. ENG.CHEM.. 45, 2394-2401 (1953). (12) Teorell, T., Proc. SOC.Exptl. Biol. M e d . , 33, 282 (1935). (13) Teorell, T., 2.Elektrochem., 55, 460 (1951). (14) Wiechers, S.G., and Van Hoek, C., Research, 6, 192 (1953). (15) Winger, A., Bodamer, G., and Kunin, R., J . Electrochem. SOC. 100, 179 (1953). (16) Ionics, Inc., Cambridge, Mass., N e p t o n Membranes, 1952. (17) Rohm & Haas Co., Philadelphia, Pa., Amberplex A - 1 (Prelzminary Notes), 1952. (18) Ibid., Amberplex C-f (Preliminary Notes), 1952. (19) Ibid., Amberplex Ion Permeable Membranes, 1952. RECYBIVI~D for review June 10, 1954.
ACOBIPTED November 1 1964.
Electrolytic Regeneration of Spent Pickling -Solutions HENRY C. BRAMER, MeZZon I n s t i t u t e , P i t t s b u r g h , P a .
JAMES COULL, University of P i t t s b u r g h , P i t t s b u r g h , Pa. T h e results of a laboratory investigation of a proposeG process for t,,e electrolytic regeneration of spent sulfate pickling solutions are presented. The process makes use of selectively ion permeable membranes to separate the solution from the recovered acid so that iron deposition can occur in a low acid environment. The important operating variables are discussed. A pilot scale investigation of the process would be justified.
T
HE spent sulfate pickling solutions from certain steel making operations present a difficult waste disposal problem. It has been estimated that nearly a billion gallons of such solutions
are produced annually in the United States with a n average composition of 15% ferrous sulfate and 5% free sulfuric acid. Although many processes have been proposed t o recover values from this waste material, there is no generally accepted process a t the present time which can be operated at a profit or which January 1955
can be operated at no net cost. This investigation has been undertaken t o establish the potentialities of a proposed electrolytic process for the continuous regeneration of spent sulfate pickling solutions, using ion exchange membranes. The objectives of such a process would be recovery of iron from the liquor as metallic iron, production of an effluent suitable for discharge to a stream without further treatment, and production of acid of the concentration required for re-use in the baths.
INDUSTRIAL AND ENGINEERING CHEMISTRY
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