information is applicable in selecting solvents for either extractive distillation or extraction. Past correlations make the pattern usable for materials of higher or loiver molecular veight in the consideration of paraffin-olefin as \vel1 as paraffinaromatic and naphthene-aromatic separations. Sulfolane and its derivatives occupy superior selectivity positions over the entire solvency spectrum and in view of their superior stability ( 3 ) -present a n attractive solvent family for man)' applications.
(2) Deal, C. H., Derr, E. L., Papadopoulos, M. N., Ind. Eng. Chem. Fundamentals 1, 17 (1962).
(3) Deal, C. H., Evans, H. D., Oliver, E. D., Papadopoulos, M. N., Petrol. Rejner 38, 185 (1959). (4) Doring, C., Z . Chem. 1, 347 (1961). (5) Dunn, C. L., Millar, R. W., Pierotti, G. J., Shiras, R. N., Sounders. M., Trans. Am. Inst. Chem. Engrs. 41, 631 (1945). (6) Edeleanu, L., U. S. Patent 911,553 (1909). (7) Gerster, J. A , , Gorton, J. A., Eklund, R.-B., J. Chem. Eng. Data 5, 423 (1960). ( 8 ) Pierotti, G. J., Deal, C. H., Derr, E. L., Ind. Eng. Chem. 51, 95 (1959). RECEIVED for review February 1, 1963 ACCEPTED June 15, 1964
Literature Cited
(1) Colburn, A . P., Schoenborn; E. M., Trans. A m . Inst. Chcm. Engrs. 41, 421 (1945).
Division of Petroleum Chemistry, 142nd Meeting, ACS, Atlantic City, N. J., September 1962.
THERMODYNAMIC CYCLES FOR RECOVERY OF WATER BY SOLVENT EXTRACTION R I C H A R D R . DAVISON AND
DONALb W .
HOOD
Texas A Z M L.;lioersity, College Station, T e x .
It can use low quality or
Solvent extraction desalination allows versatile use of available energy sources.
waste heat, since operation temperatures are near ambient, requiring only about a 50" tween heat source and cooling water. tion of only
40 to 80 6.t.u. per pound of
F.
difference be-
Efficient use of mechanical energy is also possible, since heats of s o b water dissolved are involved.
lrreversibilities of mass transfer tend
to limit its use to lower concentrations; but for 5000-p.p.m. feed, heat pump cycles requiring 20 kw.-hr. or less per 10010 gallons of fresh water are devised. A cycle is also shown which uses only a little over 200 pounds of high pressure steam per 1000 gallons of product.
extraction is an important separation process applied to mixtures of electrolytes; yet solvent extraction for desalination has had relatively little attention. This negkct resulted initially from a misunderstanding of the problem. Extraction. while much used to separate various species of electrolytes. Lvould not appear effective in removing mosi of the salt from concentrated solutions. Should a solvent be found that can extract salt from water, the problem of removing the salt from the solvent solution would still remain. and to complicate the picture, any liquid which appreciably dissolves water-soluble salts would almost certainly be miscible \vith Xvater. Ellis (5) pointed out these probabilities some yea.rs ago and discounted solvent extraction as a method of desalination. These arguments have all failed in not realizing that. although solvents to extract salt from Lvater are unlikely, it is relatively simple to extract water from salt. There are a large number of organic compounds which are partially miscible Lvith water and which selectively- dissolve \vater from salt solutions. If in turn the solubility changes rapidly Lvith temperature, a phase consisting largely of water with a reduced salt content may be obtained by changing the temperature of the previously saturated organic phase. T h e best solvents u i t h respect to nearly all properties are a group of secondary and tertiary amines (7, 7). Figure 1 shows the solubility curves of triethylamine and methyldiethylamine and mixtures of the two. I n general, solubility has been found to be more sensitive to temperature changes for solutions having lower consolute temperatures; thus most of the solvents applicable to the process are of this type. OLVENT
S which is often
T h e extraction is carried out countercurrently, with extract reflux if necessary, a t a temperature such that the extract contains 30 to 35y0water, the concentration being limited by a decrease in selectivity with increasing water content. T h e number of stages and the reflux rate depend on the concentration of feed, product, and raffinate. After extraction, the extract is heated to separate the reflux, which is returned to the extractor, and further heating removes the product. T h e water product contains several per cent of amine which must be recovered. After cooling, the solvent is recycled to the extractor. '4 cursory evaluation of what has been accomplished by the extraction is not too encouraging. Now, instead of being mixed with salt, the product contaics a n even larger mole ratio of a relatively expensive organic compound; so that the water activity is probably lower now than a t the beginning. T h e solvent must be recovered completely; and theoretically it will take as much (or more) energy to accomplish this as to separate the salt and water in the first place, and in addition to the energy already spent separating the bulk of the solvent from' the water phase. These seemingly discouraging facts have led many reviewers of the saline water program to dismiss solvent extraction as a practical method without investigating its advantages. I t has been reiterated ( 2 , 6) that since the water would not enter the solvent phase unless its activity was thereby lowered, the process, energywise, is not very promising. However, this argument is valid only if the magnitude of the irreversibilities is large. VOL.
3
NO. 4
OCTOBER 1 9 6 4
399
I80
1-1
6
MIXTURE
H I-2MlXTURE
160
V-4
0
1 01
L
-
02
i
03
I-3MlXTURE
. 04
L - L .
05
06
in which .\- is the pounds of salt rnoving through the system per pound of product produced, and .t is the \$,eight fraction of salt. For sea \cater. Equation 1 gives a little over 1 B.t.u. per pound or about 2.6 k\c.-hr. per 1000 gallons. If a feed water containing 5000 p . p . m is concentrated tenfold. the work is found to be 0.455 B.t.u. per pound or 1.11 kLv..-hr. per 1000 gallons. In processes involving a phase changt.. the latent heat muSt be recovered for cfficient operation. For instance. sea ivater boils about 0.75' F. higher than fresh ivater; so that if the heat liberated a t condensation is pumped to a temperature 0.75" higher, it may be used to boil the sea Ivater. In an idealized process. this is done by a Carnot cycle operating reversibly betu.een the boiling and condensing temperatures. Assuming all other steps to be reversible. the \cork of the Carnot cycle will be the same as that given by- Equation 1 or 2. .4t 212" F. the latent heat is 970 B.t.u. per pound, and the Carnot cycle work for delivering this quantity of heat from the condenser to the boiler is given by
I
H TRIETHYLAMINE H METHYLDIETHYLAMINE
I
I
1
1
07
08
09
IO
WEIGHT F R A C T I O N W A T E R Figure 1. Solubility curves for methyldiethylamine, triethylamine, and mixtures of the two
A distinct advantage of solvent extraction which has come to be appreciated (3, 7) is that the operating temperatures are near ambient, making possible the use of cheap: low grade heat sources. Furthermore. Figure 1 shows that the extraction temperature may be varied from about 65' to 125' F.. as required by environmental conditions. without loss of efficiency. T h e main thermodynamic advantage of solvent extraction is that heats of solution are much lower than heats of fusion and vaporization. It may be easily shown by devising reversible cycles employing the particular phase change in question that the magnitude of the latent heat does not affect the minimum work required to separate a .water of given concentration (8); but it has been argued very convincingly ( 3 ) that the theoretical work is of little importance, since the actual work is many times larger. T h e latent heat does have a great effect on the actual work required. In order to determine to what extent a low latent heat can offset the disadvantage associated Lvith liquid-liquid mass transfer, consideration must be given to the thermodynamics of both saline water conversion and the solvent extraction processes. Triethylamine and methyldiethylamine. for ichich data are available, will be used to evaluate the magnitude of various irreversibilities involved.
1.07 B.t.u.,'lb. or 2.6 k\c.-hr. per 1000 gallons. The heat of fusion of Lvater is only 144 B.t.u. per pound. but the freezing point depression is correspondingly larger than the boiling point rise. T h e reversible \vork is slightly higher, since the cycle operates below ambient. and the work done by the Carnot cycle must be discharged to the surroundings bv a second Carnot cycle. The heat of solution of Lvater in solvent extraction varies from 40 to 80 B.t.u. per pound of water extracted. .At the loiver figure, the average temperature range through Lchich the Carnot cycle must operate is about 19" F. The theoretical reversible \cork for this cycle is the same as for distillation. In practice. hoivever. each process must pump its latent heat through larger temperature differences than theoretical, or heat exchangers of infinite size \\,odd be required, and the efficiency loss per degree of driving force is inversely proportional to the latent heat. The magnitude of the driving force will of course vary from process to process; and a low driving force may compensate for a high latent heat, as conversely a low latent heat allows a large driving force. Extraction Thermodynamics
For a solvent extraction process employing a Carnot cycle and operating above the ambient temperature. the work is given by
ideal Energy Calculations
If water is separated reversibly from a salt solution without change in salt concentration. the n o r k per pound of water is given by
I.$'
=
RT - In (p,/lp) 18
I n any practical process the concentration must change. and the reversible work per pound of product is given by
400
I&EC
in Lvhich X is the weight ratio of \cater to solvent. T , is the temperature of separation, and T , is the temperature of extraction. In terms of bveight fractions this becomes
PROCESS DESIGN A N D DEVELOPMENT
or
in which x is the weight fraction of water. T h e heat of solution is a function of concentration, but it can be assumed constant without serious error. 'The ratio ( T , - T e ) / T is , also a function of concentration for a reversible cycle. For operation below ambient temperature, a n additional quantity of work is required, as given by
7r
in which Il-is the results given by Equation 5. From Equation 3. i t may be seen that if ( T , - T , ) / T , is plotted against X , the area under the curve will be proportional to the work required. Since both l', and T , may vary, it is convenient to choose a hase temperature which is here arbitrarily chosen to be T,, the temperature of the extract leaving the extractor. LVith this modification, Equation 3 becomes
T h e work is now proportional to the area between the two curves obtained by plotting ( T , - T E ) / T ,and ( T e- T,)/Te us.
-'L, L
x.
Curves of this type are shown in Figure 2, in which T is either T , or T,. depending on whether extraction or separation temperatures are involved. Area I in this figure represents the \cork of a reversible solvent extraction cycle discussed above.
-2
01
02
03
04
I 5 0.6
X Figure 2.
Work of so veni-extraction cycles
Evaluating Energy Losses
Csing triethylamine and methyldiethylamine, calculations have been made for various cycles producing 500 p.p.m. \cater from feed concentrations of 1500 p.p.m,, which requires no reflux. and 5000 p.p.m. For the higher concentration, 0.33 and 0.39 pound of reflux per pound of product is required for triethvlamine and methyldiethylamine, respectively. Arras 1 and I in Figure 2 show the work of a cycle in which all heat is transferred reversibly, but in which the extractors oprrate isothermally with Te = T F (125' F. for methyldiethyla m i r ~ ea r d 67' F. for triethylamine). thus giving a considerable driving force for the solution of water. An infinite number of compression stages are employed. T h e results (Tables I and 11) were obtained by fitting ( T , - T e ) / T s
Work of Various Solvent Extraction Cycles Using Methyldiethylamine Heat 'YO. of Compressor Exchange Work, Com$ressor Feed F@ciency. Terminal, Km.-Hr./ Stages Concn. AT, "F. Too0 Gallons
Table 1.
so
m
2
1
1500 5000 1500 5000 1500 5000 1500 5000 1500 5000 1500 5000 1500 5000 1500 5000
100 100 80 80 100 100 80 80
80 80 100 100 80 80 80 80
0 0 0 0 0 0
0 0
10 10 0 0 0 0 10 10
1 70 2.29 2.12 2.87 3.47 4.68 4.34 5.86 7.12 9.60 5.58 7.53 6,98 9.42 9.75 13.16
with a polynominal and integrating Equation 5. the values of T , being determined from the saturation curves in 1:igure 1 and with T , constant. Results are given for heat pump efficiencies of 80 and 100O/c. The lower percentagr acts as a multiplication factor and does not change the shape of the areas in Figure 2. For triethylamine, in Table 11. the results include the work computed from Equation 6, assuming that T , is 85' F. and T , is 67' F. Since it is not economical to employ- a large number of compression stages, results are also given in Tables I and I1 for only one and two stages. Areas I, 1: and 2 in Figure 2 represent a cycle with only one compression stage in which the
Table II.
Work of Various Solvent Extraction Cycles Using Triethylamine Neat N o . of Compressor ExchanFe W707k, Compressor Feed Egiciency. Terminal. K u -Hr / Stages Concn. % AT, ' F 1000 Gallons m
2
1
1500 5000 1500 5000 1500 5000 1500 5000 1500 5000 1500 5000 1500
5000 1500 5000
VOL. 3
100 100 80 80 100 100 80 80 80 80 100 100 80 80 80 80
NO. 4
0 0 0 0
0 0 0
0 10
10 0 0 0 0 10 10
OCTOBER
1 84 2 38 2 30 2 98 3 83
4.96 4.79 6.20
40 9 2: 6 45 8 36
7
8 O7
10 44 10 613 51
1964
401
necessary for all except dilute feeds. T h e increase in reflux with salt content tends to limit the solvent extraction process to water less concentrated than sea water. The use of reflux increases the solvent rate which; as seen in Equations 3 and 7, acts as a multiplying factor on the work loads sho\vn in Figure
i
T h e preceding discussion may be summed u p in the following expression :
:TOR
60 5
1
----- - _ - _ - _ _ _ _
Figure 3. amine
FEED-1.1
-1
Solvent extraction desalination
I RAFFINATE'O
107
with triethyl-
extractor operates isothermally but in which all heat exchangers have complete temperature approach. T h e preceding examples assumed reversible heat transfer, but results are also given in Tables I and I1 in which an approach temperature of 10' F. is assumed on the liquid-liquid heat exchangers while a zero driving force is still assumed for heat transfer in the heat pump. T h e extra heat load caused by incomplete temperature approach in the liquid-liquid exchangers must in either case be pumped at least to the final separation temperature. For methyldiethylamine this is 167' F. or 32" higher than T,;for triethylamine it is 85' F. or 18' higher than T,. As a result, the penalty for incomplete approach in the heat exchanger is higher for methyldiethylamine, even though the quantity of heat thus added to the extractors is lower (see Tables I and 11). T h e additional lvork, because of incomplete temperature approach, is not easily represented on a plot such as Figure 2, since such a plot would be a function of temperature approach rather than concentration. This work is given by
for methyldiethylamine operating above ambient, in which Q is the heat lost by incomplete temperature approach in the solvent-extract exchanger. For triethylamine, it is larger, as calculated from Equation 6, and includes heat from both the feed and recycled solvent. Similar computations were made for a cycle with one compression stage but with a 10' F. driving force at both the extraction and separation temperature, as represented by the total area in Figure 2. For methyldiethylamine using a 5000p.p.m. feed. 80% efficiency of the compressor, and a product of 500 p.p.m.: the total work load is 21 kw.-hr. per 1000 gallons, or 19 times the thermodynamic minimum work for the process. A s previously noted, the product Lvater contains several per cent amine which must be removed. T h e water activity may in fact be lower than that of the feed, and the minimum energy to remove the amine is in the neighborhood of 2 kw.-hr. per 1000 gallons. Using a heat pump cycle for this purpose ivould require about 5 or 6 kw.-hr., but because of the high volatility of the amine: it is easily removed by stripping, with the expenditure of a small amount of low grade heat energy. More serious than the solvent removal is the fact that salt is not completely excluded from the solvent phase, making reflux 402
I&EC P R O C E S S DESIGN A N D DEVELOPMENT
in \vhich A T I = the temperature difference corresponding to an ideal cycle (area I in Figure 2). AT1 = the additional A T caused by isothermal operation of the extractor (area 1 in Figure 2). 17'2 = the average driving force caused by using a finite number of stages (area 2 in Figure 2). 17-3and AT* = the temperature approach a t the low and high temperatures, respectively, of the heat pump cycle (areas r . 3 and 4). I = heat pump discharge temperature. Q = the sensible heat loss through irreversible heat transfer per pound of xvater dissolved. T h e area shown in Figure 2 corresponds to the first term in brackets, Lvhile the second term corresponds to rhe additional work caused by irreversible heat transfer on the liquid-liquid exchangers, primarily the solvent-extract exchanger. T h e effect of increasing 17.is additive, except for the liquid-liquid exchangers, for which increasing A T increases Q. All major irreversibilities have been included except pumping. which like solvent recovery \.iould be an added term in Equation 9. Cnlike solvent recovery, ho\vever, it is not subject to thermodynamic analysis and is highly dependent upon mechanical design.
Real Extraction Cycles
T h e preceding calculations show in general where the principal irreversibilities lie. They do not show the results for a particular flowsheet, many varieties of which are possible. T h e following examples are more specific, and the calculations have been made in greater detail. Heat and material balances Lvere made on all major items of equipment, and actual rather than constant values of the heat of solution ivere employed. Product feed and raffinate concentrations are i00,5000,and 50,000 p.p.m.>respectively. Figure 3 is a floivsheet of a process employing triethylamine. The ordinate sho\vs the temperature in various parts of the process. \Vith triethylamine, the extraction temperature is below that of normal cooling media; so that any unused heat removed from the extractor must be pumped to some existing heat sink, which in these calculations was assumed to be a t 85' F . T h e feed and solvent at 77' F. enter the extractors, which are held at 67' F. The flow rates of each stream are shown in units of pound per pound of amine-free product water-that is, 1.1 pounds of feed, 4.3 pounds of solvent, and 0.322 pound of reflux enter the extractors per pound of product produced. T h e raffinate flow from the extractors is 0.107 pound, which is reduced to 0.10 pound by recovery of the solvent. Reflux is separated from the extract by heating with the recycled solvent. T h e temperature change of the solvent stream is
much more than that of the extract because of the heat of solution absorbed by the extract as phase separation occurs. In order to run the extractors isothermally, as shown, over 1,000.000 B.t.u. per 1000 gallons of product must be removed. This is done by operating the extractors a t a pressure such that the contents are cooled by vaporization. The vapors in this instance have a composition almost identical to that of the recycled solvent-i.e., 6.6y0 water, which is also very close to the azeotrope compclsition. These vapors are compressed from 55 to 110 mm. of mercury, and injected into the extract, which heats it to 85' F. At this temperature, 1.05 pounds of water phase are removed, which yields 1.0 pound of product. The work of compression is 17.75 kw.-hr. per 1000 gallons, assuming the compression to be 80% efficient. This is larger than the 13.51 kw.-hr. given in Table I1 for similar conditions. Table I1 was calculated using a Carnot cycle, which is inherently more efficient than a vapor compression-condensation cycle in which the condensed refrigerant is returned to the evaporator a t a temperature above that of the evaporator. In addition a 2.5' F. 'driving force was assumed, and the mixture in the extractor is not quite that of the azeotrope; thus the pressure is below that of the azeotrope. Even so, 17.75 kw.-hr. is a very low work requirement. These results were obtained using one stage, and the data in Table I1 indicate that a second stage would reduce this by about 4.25 to 13.5 kw.-hr. per 100 gallons. There still remains the work of solvent recovery. If some source of low grade heat is available, this can be done cheaply by stripping. Csing vapor compression distillation for this separation would require about 5 to 6 kw.-hr. additional energy. In Figure 4 a cycle using methyldiethylamine is shown which differs in several ways from the previous design. After being heated from 125" to 128' to separate reflux water, the extract continues to exchange heat with both water product and solvent until it reaches a temperature of 138" F. T h e heat of solution of water in niethyldiethylamine is only about 44 B.t.u. per pound of water dissolved. As a result, the heat required
to raise the temperature of the extract a given amount is less than for triethylamine. A conventional refrigeration cycle using n-pentane (4)as the refrigerant is used to pump the heat of solution from the extractor to the separator. An approach temperature of 10" F. is assumed in both the evaporator and condenser of the heat pump cycle. A steam turbine is used to furnish power to the heat pump compressor, and the exhaust steam furnishes heat to the processes. T h e condensing refrigerant heats the extract from 138' to 153' F., and the exhaust steam heats i t to 157' F. Assuming the turbine to be 25% efficient, it is sized so that the exhaust steam will be just sufficient to supply the remaining heat requirements. In this design. 208,700 B.t.u. are supplied to the process while providing 15.3 kw.-hr. of mechanical work per 1000 gallons of product. If solvent evaporation rather than a separate refrigerant were used, these figures would have been further reduced. In this case. there would be a small loss of efficiency because the vapors are lower in water content than the recycled solvent; but because the vapor pressure is about 500 instead of 55 mm., the mechanical requirements of the compressor would be much more reasonable than for triethylamine. If, in the preceding cycle. the compressor and turbine are eliminated and the heat in the extractors is rejected to cooling water. 646,000 B.t.u. of heat energy must be supplied to raise the temperature of the extract from 138' to 157' F. This represents an additional heat load of 437,000 B.t.u. per 1000 gallons. If heat is available in the form of high pressure steam, the choice of these alternatives could be determined by an economic balance between the additional heat requirements and the cost of the turbine and compressor. If low level heat is available, its use will tend to increase the investment for heat exchangers and solvent recovery equipment, but will probably give lower costs than the more efficient cycle using mechanical energy. Conclusions
Contrary to some opinions, the solvent extraction process for desalting water is not a low efficiency method if the design characteristics are made to minimize energy utilization. Alternatively, if low temperature heat sources such as low pressure exhaust steam or solar heat are used, efficiency becomes impossible but less important, and solvent extraction can be adapted to such heat sources better than any known method. In fact, if a heat pump cycle is not used, the cost of the solvent extraction process is within limits fairly independent of the absolute temperature of operation and depends only on the temperature difference between the heat and cooling sources. A difference of 50" F. is sufficient to operate the process with an extraction tempmature anywhere between 65" and 125' F.
I
'.:t 155
Nomenclature
p p,, Q
I2O
t t
'I5 Figure 4. Solvent methyldiethylamiiie
S
= = = =
W = 11" = extraction
desalination
with
X
=
x
=
X
=
vapor pressure of salt solution vapor pressure of water heat pounds of pure solvent per pound of pure water produced work work to pump the heat added by compressor to ambient temperature weight ratio weight fraction latent heat VOL. 3
NO. 4
OCTOBER 1 9 6 4
403
SUBSCRIPTS: a
e
E J r
= =
s
= =
=
S
literature Cited
ambient conditions extraction conditions final extraction condition feed raffinate separation final separation conditions
= =
(1) Daviqon. R. R.: Smith, FV. H., Jr., Hood, D. \V.; J . Chem. En?. Data 5 , 420 (1960). (2) Dodge, B. F., A m . Scientist 48, 477 (1960). (3) Ibid.,49, 54a (1961). (4) Edmister, C. W., Petrol. Rejner 37, 194 (June 1958). (5) Ellis, C. B., “Fresh Water from the Ocean,” p. 97, Ronald Press, New York, 1954. (6) Gilliland, E. R., Ind. Eng. Chem. 47, 2410 (1955). (7) Hood, D. W., Davison, R. R., Adoan. Chem. Ser., No. 27, 40 (1960). (8) Murphy, G. LV., “Minimum Energy Requirements for Sea \Yater Conversion Processes,’’ Office of Saline IYater Research and Deyelopment, Progr. Rept. 9 (1956).
Ac knowledgment
T h e support of this work by the Office of Saline Water,
U. S. Department of the Interior, by Contract No. 14-01-0001282 with the Texas A&M Research Foundation is gratefully acknowledged. T h e use of the digital computer in the Data I-’rocessing Center of the Texas Engineering Experiment Station is appreciated.
RECEIVED for review September 3 , 1963 .qCCEPTED March 20, 1964
NEW DEIONIZATION TECHNIQUES BASED UPON WEAK ELECTROLYTE ION EXCHANGE RESINS R O B E R T K U N I N AND
B A S I L V A S S l L l O U
Rohm t 3 Haas Co.? Philudelphia, Pu.
The economic deionization of water with ion exchange resins has been limited generally to waters that have less than 600 p.p.m. of total dissolved solids. In general, deionization of brackish waters b y conventional ion exchange resin techniques cannot compete with the membrane and distillation methods, primarily because of regenerant costs and the need of deionized water for the rinse operation. For low solids waters, these factors are of minor importance; above a concentration of 500 p.p.m., they become of maior importance. A new deionization technique i s based upon weak electrolyte ion exchange resins in which the weak base anion exchanger i s employed in the bicarbonate cycle to convert the mineral acidity of the water to alkalinity, and the weak acid resin, to remove the alkalinity. This system permits one to deionize brackish waters with but little leakage and a t high regeneration efficiency.
a detailed study (6) on the regeneration of carboxylic cation exchange resins with carbon dioxide, it was discovered that certain anion exchange structures of a weakly basic nature could form the bicarbonate salt with solutions of carbon dioxide and that the chloride-bicarbonate selectivity coefficient, K;&, was favorable. This discovery suggested the possibility of employing weak electrolyte ion exchange resins in a reverse deionization system (anion exchanger followed by cation exchanger) based upon the following reactions: URING
+ HtO + C O Z
[R-N]
+
H C 0 3 (carbonation)
[R-NH]
(1)
[R-NH]
HCOI
+ NaCl
+
[R-NH]
C1
+ NaHCO, (alkalization)
+ NaHC03
R--COOH [R-NH]
C1
2RCOONa
404
+
+ NHiOH
+ H2SO4
R-COONa
+
f CO2 4- H20 (dealkalization) (3)
+
+ [R-N] NHdC1 H10 (anion exchanger regeneration)
(4)
+
2RCOOH Na2SOl (cation exchanger regeneration)
+
(2)
(5)
l&EC PROCESS DESIGN A N D DEVELOPMENT
Since the equilibria for these ion exchange reactions were favorable, the basic aspects of these reactions and the deionization system were studied in detail. Experimental Procedure
Ion Exchange Resins. T h e properties of the ion exchange resins investigated are summarized in Table I . Unless specified, all resins were in the 20- to 50-mesh particle-size range. T h e columnar studies were performed in a pressurized unit previously described (6) and the experimental details. except for the carbonation reaction, were similar to those described in the same paper. T h e h e a k base anion exchange resin was carbonated by passing water saturated with C O S a t various pressures over the basic form of the resin. The weak base anion exchange resin was regenerated using 1.1- “,OH. T h e deionization studies were all conducted on a NaCl influent and the effluents of all columns were monitored for chloride, bicarbonate, pH, and conductivity. Leakage data were calculated from the conductivity a t low leakage levels and from the chloride and alkalinity data at high leakage levels Most of the studies involved the passage of NaCl over a bed of a
