ENGINEERING, DESIGN, AND PROCESS DEVELOPMENT
a
*
As an alternative to evaporation at 100” C., the solution may be cooled to 0 ” C. where sodium sulfite heptahydrate is the equilibrium phase. The anhydrous salt will be converted to the heptahydrate, with the removal of some water from the solution. By using this method, 76.2y0 of the sodium sulfite can be recovered from the Vorce cell liquor and 83.7% from the ColumbiaHooker cell liquor. T o bring the composition of the sulfited liquor closer to the invariant point, the solution can be saturated with sodium chloride at 100” C. by the addition of 13.6 pounds of sodium chloride to the Vorce cell liquor and 15.7 pounds to the ColumbiaHooker cell liquor. This will cause the precipitation of sodium sulfite, so the recovery from Vorce cell liquor will be 83.0% and t h a t from Columbia-Hooker cell liquor will be 86.3%. Some sodium sulfite iemains in solution after each of these processes and it inust be removed or reduced considerably before the sodium chloride brine can be re-used in an electrolytic cell. One possibility is t o oxidize the sulfite to sulfate using aeration. The solution can then be cooled to 0 ” C. to remove sodium sulfate decahydrate and leave a final solution containing 1.86 pounds of sodium sulfate and 36.4 pounds of sodium chloride per 100 pounds of water ( 1 3 ) . The raw brine from salt vi-ells or dissolvers must be processed to remove calcium and magnesium ions before the brine can be used in the electrolytic cell. The reaction of sodium sulfite with calcium salts to form calcium sulfite is of interest. Lime slurry tould be added to the solution containing the residual sodium sulfite. For calcium sulfite in water, the solubility is 0.011 gram calcium sulfite per 1000 grams of solution at 100” C. ( 1 1 ) and a t 18” C. is 0.043 gram per liter of solution. Thus the solution from the evaporation steps could be limed t o yield a
Potential for Steel
final solution containing only about 0.0005 pound calcium sulfite per 100 pounds of initial cell liquor. References
Domask, W. G., and Kobe, K. A., Anal. Chem., 24, 989-91 (1952). Foerster, F., Brosch, A., and Norberg-Schulz, C., 2. physik. Chem., 110, 443-55 (1924). Handbook of Chemistry, N. A. Lange, editor, 6th ed., p. 871, Handbook Publishers, Sandusky, Ohio, 1946. Hartley, H., and Barrett, W. H., J. Chem. Soc., 95, 1178-85 (1909). Heinemann, Gustave, Columbia-Southern Chemical Corp., Corpus Christi, Tex., personal communication, March 23, 1954. International Critical Tables, IV, 235, 1928. Kirkpatrick, S. D., Chem. & Met. Eng., 35, 158-61 (1928). Kobe, K. A., “Inorganic Process Industries,” p. 133, Macmillan, h’ew York, 1948. Lewis, N. B., and David, A. C., J. Chem. Soc., 125, 1156-62 (1924). Illantell, C. L., “Industrial Electrochemistry,” p. 420, McGrawHill Book Co., New York, 1950. Seidell, A., “Solubilities of Inorganic and Metal Organic Compounds,” 3rd ed., Vol. I, p. 327, Van Sostrand, New York, 1940. Ibid., pp. 1217-18. I b i d . , p. 1234. Ibid., pp. 1296-7. Seidell, rl., and Linke, W. F., ”Solubilities of Inorganic and Metal Organic Compounds,” 3rd ed. suppl., pp. 450-3, Van Nostrand, Kew York, 1952. Thomoson. E. T.. Chem. Ena., 61, No. 6. 105 (1954). Wendrow, B., and Kobe, K: A., IND. ENG.CHEM.,44, 143947 (1952). RECEIVED for review- July 19, 1954.
Mill Waste Disposal
ACCEPTED January 3, 1955.
...
Electrolytic Treatment of Waste Sulfate Pickle liquor.Using Anion Exchange Membranes C. HORNER,
A. G. WINGER, G. W. BODAMER, AND R. KUNlN
Rohrn B Haas Co., Philadelphia, Pa.
PEST pickle liquor is produced by the process in which iron oxide scale is removed from semifinished steel by immersion in a dilute (15 t o 25%) sulfuric acid solution. The acid reacts with the oxide and some of the base metal to form ferrous sulfate, and when the liquor is finally discarded it may contain from 12 t o 22% ferrous sulfate and from 0.5 to 10% unused free acid. It has been estimated that nearly 1 billion gallons of spent liquor is produced annually. Since the latent values of the components of spent pickle liquor allow little margin for turning the waste into a profit, the problem has come to be treated almost exclusively as a problem of waste disposal. Years ago this presented little difficulty, for frequently the waste liquor could be disposed of b y discharging i t untreated into nearby rivers, lakes, or abandoned mines. More recently, however, in view of the growth of population, the establishment of suburban communities, and the enormous increase in the volume of liquor to b e disposed of, civic lune 1955
groups and legislative bodies have cooperated to prohibit pollution of the rivers, lakes, and underground water tables in this way. It thus has become necessary to employ other means of disposal. Probably the least costly and most often employed disposal method is that of neutralization with lime, followed by “lagooning” the resulting slurry. This method introduces the additional cost of neutralizing agent and its handling and involves the cost of lagooning areas. With the advent of Amberplex (a trademark of the Rohm & Haas Co.) ion exchange membranes, however, another avenue is opened whereby the recovery of electrolytic iron and the regeneration of sulfuric acid may be effected, while at the same time the disposal problem is eliminated. Such a process would take on added significance in time of sulfuric acid shortages, as recently experienced. An ion exchange membrane is simply an ion exchange resin in sheet form rather than in the conventional bead or granular
INDUSTRIAL AND ENGINEERING CHEMISTRY
1121
ENGINEERING, DESIGN, AND PROCESS DEVELOPMENT form; and whereas the granular resin is customarily employed in direct chemical processes and is regenerated b y chemical means, the membrane finds particular merit in electrochemical processes, where i t requires no regeneration. The most important property of an ion exchange membrane is its permeability to ions of one charge and impermeability to ions of the opposite charge. Thus, since ion-permeable membranes may be made in either cation-permeable or anion-permeable form, control of ionic movement of either species may be exercised.
+\
ANODE
AMBERPLEX ANION EXCHANGE M E M B R A N E 7
REACTION:
4 OH--
Figure 1 .
0,t
+ 2H20+4e
CATHODE
/-
Effects of operating variables on batch treatment in single compartment cell are investigated
Figure 2 shows the assembly of the actual cell used in the Rohm & Haas laboratories for this work. As in most design problems, compromises had to be made. T h e electrodes had to be far enough apart to allow room for the build-up of electrodeposited iron and the ready escape of gases, but they had to be close enough together to reduce the voltage to a minimum. The cell shown is 12 X 1 8 l / 2 X 2 3 / 4 inches, comprising simply a catholyte chamber wall of Plexiglas, next to which is located a Type 316 stainless plate cathode, followed in turn by a $ / l e inch catholyte chamber, the Amberplex membrane, a coarse mesh saran screen, an antimonial lead anode, a glass heat exchanger, and finally the Plexiglas wall of the anolyte chamber.
REACTIONS:
2 H + + 2 e -HZ Fe++ + 2 e - F e 0
t
Mechanism of electrolytic treatment of waste sulfate pickle liquor
Since cationic iron may be electro-deposited only from a solution relatively low in acidity, and since deposition must take place at the cathode, i t follows that the sulfuric acid must be concentrated in the region of the anode. Furthermore, since hydrogen ions from the area of the anode must be prevented from flowing to the cathode, but sulfate ions from the cathode region must be allowed to flow to the region of the anode, an anionpermeable membrane should be selected. The use of a cationpermeable membrane would permit the free flow of hydrogen ions as well as ferrous ions to the cathode, and since hydrogen ions are the more mobile, they would be transported in far greater numbers, resulting in a hydrogen ion concentration where little or no deposition of iron takes place. Figure 1 shows the schematic plan of an electrolytic cell for the study of batch treatment of spent pickle liquor. It consists of two compartments separated by an Amberplex anion exchange membrane, a suitable anode and a suitable cathode. The spent liquor to be treated is placed in the cathode chamber, becoming the catholyte, while a dilute sulfuric acid solution may be used as the anolyte. If the membrane is perfectly “permselective,” for each faraday of current one equivalent of Soh-- ions from the catholyte migrates across the anion-permeable membrane into the anolyte, while one equivalent of OH- ions is oxidized and released as gaseous oxygen a t the anode and a total of one equivalent of H * ions and F e + + ions is reduced at the cathode. low pH-in If the hydrogen ion concentration is high-Le., the catholyte, only H + ions will be reduced a t the cathode; however, when the p H rises to a value of about 1.8, as the result of 1122
hydrogen liberation, F e + + ions will codeposit. At the same time, migration of Soh-- ions through t h e membrane together with the release of oxygen a t the anode result in the concentration:of sulfuric acid in the anolyte. A t a p H of about 1.8 the rates of hydrogen and iron reduction appear to be in equilibrium. If the p H should be artificially raised above this value, iron will be deposited faster than the equilibrium rate, tending to lower the pH. Similarly, if the p H should be reduced below this value, hydrogen will be discharged a t a faster rate, tending to bring the p H up again.
118’ T Y P E 316 SS. CATHODE MEMBRANE RAN SCREEN
8’ Sb-Pb ANODE PLEXIGLAS
WALL
f
P L A N VIEW
Lhdl
Figure
2.
Details of electrolytic cell assembly
For treatment in this batch cell, a representative synthetic waste pickle liquor was made up, containing 15.7% ferrous sulfate and 4.0y0sulfuric acid. This was to be the catholyte solution. The anolyte was made u p t o 3.0% sulfuric acid so as to be a good carrier of electricity. Just prior to the run, the exact volumes of electrolytes were measured and the p H of the catholyte was determined. Then the solutions were introduced into their respective chambers. The cathode was weighed, inserted into the catholyte chamber, and connected to the negative lead from a selenium rectifier of ample capacity. Then current was passed
INDUSTRIAL AND ENGINEERING CHEMISTRY
Vol. 47, No. 6
ENGINEERING, DESIGN, AND PROCESS DEVELOPMENT through the cell a t a predetermined level. During the run the voltage was adjusted as necessary to keep the current constant; readings were taken a t 15-minute intervals. T h e voltage, after an initial drop, rose slowly, increasing its rate of climb as the catholyte concentration decreased, until with the attainment of a certain high level the run was terminated.
4
:":v 0.4
reach the particular concentration, the average temperature for that time period, and the average voltage. Costs were calculated for each run in terms of kilowatt hours per liter of pickle liquor treated, as well as in terms of kilowatt hours per pound of sulfuric acid regenerated. Figure 6 is a plot similar to that shown in Figure 4, illustrating the interpolation of data for comparative cost calculations. The number of SO4-- ion equivalents available in the 970 ml. of original pickle liquor is shown by the level of the upper horizontal dotted line. The number of F e + + ion equivalents available is shown by the lower horizontal line. If data are to be extracted from this run that would apply if the catholyte mere depleted to 2.25% ferrous sulfate, that point on the lower curve must be found which corresponds to residual catholyte of 2.25%. From a knowledge of the volume and specific gravity of the catholyte a t the concentration in question it is now possible to calculate the number of F e + + equivalents a t that point. 0.0225 (%) X 1.024 grams/ml. X 755 ml. = o,23 equivalent 75.96 grams/equiv.
1
2
3
FARADAYS
Figure 3.
4
5
OF
6
7
8
9
ELECTRICITY
Variation of pH with faradays passed in cell
Figure 3 shows the value of the p H of the catholyte with respect to the quantity of current passed. The p H beginning a t about 0.5 rises steadily until a value of about 1.8 is reached, whereupon it levels off and remains essentially constant for the remainder of the run. The upper curve in Figure 4 shows the progressive transport of Sod-- ions through the ion exchange membrane as a function of faradays, or time, if current is maintained a t a constant value for the duration of the run. The lower curve shows the progressive deposition of iron simultaneous with the curve above. The distance between the curves and the limiting values (known from analysis of the initial catholyte solution) of iron and Sod-shows the concentration of the catholyte a t any time. Consideration of the voltage rise (with the consequent increase in power consumption) as the catholyte concentration is reduced leads to the establishment of economic limits of depletion. Figure 5 shows how the voltage changes as the ferrous sulfate concentration of the catholyte is reduced. The voltage increases slowly until lower concentrations of around 3 to 2% ferrous sulfate are reached, after which the voltage slope begins to rise much more steeply. It would be economically advisable, then, to deplete the ferrous sulfate concentration only to, say, 2.25%, or some other reasonable concentration, without encountering severe resistance; and since the depleted catholyte could be reintroduced as the anolgte for the next run, the residual sulfuric acid in the catholyte could be recovered and there would be no waste disposal problem. Converted into process mechanics, this would mean the spent pickle liquor could be treated and returned to the pickling tanks as a solution containing the original sulfuric acid less the sulfuric acid equivalent of 3% ferric sulfate (the ferrous sulfate introduced into the anolyte oxidizes to ferric sulfate). Costs are calculated on basis of specific concentration of ferrous sulfate in depleted liquor
For purposes of comparison, costs were calculated from data interpolated to yield specific concentrations of ferrous sulfate in the depleted liquor. More or less arbitrarily, these costs were computed to 4.5, 2.25, and 0.75% ferrous sulfate concentrations. Runs that were interrupted t o obtain spot data were followed by uninterrupted runs designed to check the slope of the Sod-transport and Fe + + deposit curves. Thus, reliable interpolated data were available from which could be ascertained the time to
June 1955
Figure 4.
Transport of Sod-- and Fe++ ions in cell Run 26
With 0.23 equivalent of F e + + remaining in the catholyte, 2.39 0.23, or 2.16 equivalents must have been removed. This point is designated by the intersection of the center vertical line with the F e + + deposit curve. The duration of this hypothetical run may be found graphically by measuring the distance from the point to the ordinate. Figure 6 indicates 6.7 faradays, or 2.99 hours. The voltage readings may be averaged and the cost of treatment determined. I n this case the voltage averages 4.17 volts, and since the current drawn is steady a t 60 amperes, 60 amperes X 4.17 volts X 2.99 hours = o.771 kk57v.-hr, per liter 1000 watts/kw. X 0.970 liters or
60 amperes X 4.17 volts X 2.99 hours 2.83 equiv. X 49.04 grams/equiv. 1000 watts/kw. X 453.6 grams/pound 2.45 kw.-hr. per pound HzS04
INDUSTRIAL AND ENGINEERING CHEMISTRY
1123
ENGINEERING, DESIGN, AND PROCESS DEVELOPMENT Similarly, the costs of treatment to other levels of depletion are determined. Figure 7 shows the dependence of the cost of treatment on the degree of depletion. Since the efficiency of the process decreases as the concentration of iron in the catholyte is reduced, the power cost for treatment to different levels of depletion is greater than proportional to the voltage.
52 grams of water, while electrokinetic transport through the membrane accounted for 176 grams more of water. The SO4-transport across the membrane amounted to 139 grams, or 2.89 equivalents, resulting in an efficiency of 41.7%. This run was terminated a t 1.8% ferrous sulfate. At 2.25% ferrous sulfate the efficiency is slightly higher (42.4%). Continuous operation of cell is advantageous, but n e w diffusion problems are encountered
12
-
IO
-
Y
4 W
5 > 6 -
4t 21
)8
I 16
I
I
14
12
Figure
5.
t 6
8
IO
% FeS04
IN
I
4
2
1 0
CATHOLYTE
It is conceivable that an automatically operated process could be designed to accommodate a batch operated electrolytic cell such as has been described. Pumps and solenoid valves, for instance, activated b y cycle controllers might be used for filling, draining, and transferring the treated catholyte to the anolyte chamber. Self-acting temperature regulators and current controllers would complete the automation. I n general, however, continuous operations are to be preferred to batch operations wherever it is possible, for, although the cell operates all the time at low concentration in continuous treatment as against only part of the time in batch treatment, this disadvantage might be more than offset by the elimination of the time lost in emptying and refilling between batches.
Voltage change with ferrous sulfate concentration in catholyte
/I /
Figure 8 is another voltage curve showing, in addition, the relationship between the current density and the voltage. The upper three curves are obtained from the same cell assembly and may be compared directly with each other. The lowest curve represents the same cell reassembled in a different way; the principal difference was that the electrodes were located closer together. The greatly reduced voltage over that of the upper 60-ampe-e run is a direct result of the closer proximity.
i 5
06'
4l5
410
315
io
F I N A L Fe SO4
215
2'0
115
I1O
015
CONCENTRATION (%)
Figure 7 . Dependence of cost of treatment on degree of depletion of liquor
I
2
3
4
5
FARADAYS
6
7
8
9
10
CONSUMED
( 2 2 4 FARADAYSIHOUR)
Figure 6. Interpretation of data for comparative cost calculations
Figure 9 is a diagrammatic weight balance showing the disposition of materials introduced into the cell. Of the original 1167 grams of catholyte only 765 remained after 3 hours, 6 minutes. Somewhat less than 5 grams of hydrogen were released while 62 grams of iron were deposited at the cathode. More than 55 grams of oxygen were released a t the anode. Evaporation took 1124
To permit continuous treatment of pickle liquor in the electrolytic cell shown in Figure 2, it was necessary only to connect the two drain outlets from each chamber by means of a small section of rubber tubing and to add a discharge spout at a n appropriate level to the anode chamber. Spent liquor was then metered into the catholyte chamber and was depleted of its free acid and ferrous sulfate components as it flowed down past the face of the cathode. From the bottom of the catholyte chamber the depleted liquor flowed through the connecting tubing and into the bottom of the anolyte compartment. As it flowed upward through this chamber it evolved oxygen and received the incoming sulfate ions to become continually more concentrated in sulfuric acid. Discharge was effected by the overflow spout, with the discharge level automatically regulating the level of the catholyte. As in batch treatment, economic limits of depletion were established to take advantage of the more favorable costs of operation. The limit chosen was such that it would also allow direct comparison between batch and continuous operations-namely, 2.25% ferrous sulfate in the catholyte. But, since ferrous sulfate upon introduction into the anolyte oxidizes to ferric sulfate and the corresponding amount of ferric sulfate is 3.0%) the continuous runs were designed to produce a 3.001, ferric sulfate effluent. Operation of a continuous cell afforded a better opportunity t o study the influence of the several variables encountered in the electrolytic treatment of pickle liquor. Accordingly, the influ-
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
Vol. 41, No. 6
.
ENGINEERING, DESIGN, AND PROCESS DEVELOPMENT
e
ence of current density, feed flow rate, temperature, and pickle liquor concentrations were studied. Higher current densities were found to have the net effect of reducing the current efficiency and of increasing the power cost. Also, higher current densities were found to increase the tendency to grow “trees” on the cathode. With regard to capital costs, higher current densities would serve to reduce electrolytic cell costs b y obtaining more work per cell, but would also increase the cost of rectification equipment by lowering the efficiency of the process. The flow rate provided the principal means of coiitrol over the catholyte concentration, which in turn influenced to a varying degree the rate of iron deposit. The total sulfate concentration of the feed liquor, too, exercised its influence upon the flow ratenot only by virtue of the number of sulfate ions to be removed, but also by virtue of the greater concentration of sulfuric acid in the anolyte and, therefore, the greater rate of diffusion from the anolyte to the catholyte. Temperature was found to be one of the more important variables. Higher temperatures not only reduced the impressed voltage required to pass a specific current, but also increased the rate of iron deposit and so increased the permissible flow rate, reducing costs all around. The effect of higher temperatures on the quality of the iron deposit is probably beneficial, though this is not conclusive. Figure 10 showe the relationship between the rate of deposit of iron and the concentration of ferrous sulfate in the catholyte. As the concentration of the catholyte is lowered, the rate of deposit of iron is reduced, a condition which contributes toward the higher cost of greater depletion. Figure 11 shows how the rate of iron deposit varies with the temperature. I n this case the advantages of higher temperatures of operation are obvious. Particularly in the case of the two-chambered continuous cell opeiated a t high temperature, diffusion of sulfuric acid from the high concentration anolyte through the membrane to the lo\\Concentration catholyte becomes significant. This diffusion, v-liich is quite apart from any electrochemical activity, tends to reduce the over-all efficiency of the process. Measurements of the rate o? diffusion mwe made under operating conditions a t
+
15.6 Yo FeS04 3.9 % H z S 0 4
3.0% HZS04
17.7% HzSO4
1.81 % FeSO, 0.35 Yo H z S 0 4
9. Weight balance for batch operation of
Figure
single compartment cell Run
39
h FeS04
CONCENTRATION
I N C A T H O L Y T E 1%)
Figure 10. Relation between rate of iron deposit and ferrous sulfate depletion in continuous operation
-
.
20
-
a
L
19-
5
18-
t ln
j 7 -
P W
16
z
15
-
‘4-
2
iL I1
60
CATHOLYTE,
Figure
8.
C O N C E N T R A T I O N ( S . Fe SO4 1
Relation between density and voltage
current
L 70
CATHOLYTE
80
TEMPERATURE,
90 OC
Figure 11. Variation in rate of iron deposit with temperature in continuous operation
INDUSTRIAL AND ENGINEERING CHEMISTRY
1125
ENGINEERING, DESIGN, AND PROCESS DEVELOPMENT various levels of temperature. The results of these measurements appear in Figure 12, where the sulfuric acid diffusion coefficient is shown as a function of temperature. As an example of the magnitude of diffusion taking place, it was determined in run 85, made a t 89" C., that no less than 20.1 grams of sulfuric acid diffused from the anolgte to the catholyte each hour.
",-
5 0 -
'0 n I
40-
c W z
-
L!
30-
LL
W
0
u
2.0
-
2
: 3
1.0-
Y
80
30
40
SO
Figure 12.
BO
70
60
T E M P E R A T U R E,
90
Multicompartment cell results in lower costs, reduced back diffusion, optimum electrode spacing
In an effort to reduce the effects of back diffusion, a threechambered cell was assembled and operated. Two barriers were thus provided through which diffusion would have t o take place between the high concentration anolyte and the low concentration catholyte. The two-chambered cell was in effect split through the membrane and a center chamber was inserted. This cell was designed so that spent liquor could be introduced into the catholyte chamber, passed from there into the central chamber, and finally passed through the anolyte chamber to discharge. Figure 14 shows the assembly of this cell. An over-all weight balance showed that evaporation from this cell amounted to 30 grams per hour, comparing well with the two-compartment cell. The comparable iron deposit rate was 30.3 grams per hour, compared to 23.83 grams per hour for the two-chambered cell. This iron deposit rate corresponds to a flow rate of 517 ml. per hour versus 397 ml. per hour. The current efficiency is 30'3 - 23'83 X 100 = 27.1% (higher) 23.83
'C
Rate of sulfuric acid diffusion in continuous operation
A schematic weight and material balance showing the quantities involved and courses taken by the several components of the incoming feed liquor is presented in Figure 13. The unit of time is the hour, and a t equilibrium all figures are constants. I n continuous operation the determination of sulfate efficiency is less direct than with batch operation, since the sulfate path is split in its approach toward the anolyte, but it may be determined by means of a series of weight and material balances based on iron. I n this example the sulfate current efficiency is 54.4 grams 48.03 grams/equiv. X 2.48 faradays
x
100 = 45.7yo
But, as the electrodes were moved farther apart to accommodate the central chamber, the voltage increased appreciably. I n this run the potential employed was determined to be 6.5 volts, yielding a power cost of 6.5 volts X 66.5 amperes = 0.836 kw.-hr. pef liter 1000 watts/kw. X 0.517 liters/hour
or, again, in kilowatt-hours per pound of sulfuric acid recovered, 6.5 volts X 66.5 amperes 541 X 0 . 1 4 gram 'Oo0 watts/kn" 453.6 grams/pound 2.59 kw.-hr. per pound H2SO
Had no diffusion taken place t o counter electrolytic transport of sulfate ions, a net of 74.1 grams per hour would have been transported, raising the efficiency from an actual value of 45.7% t o a calculated 62.4%. That this actual efficiency is higher than that illustrated for the batch cell does not in itself imply that continuous operation is more efficient from a current consumption standpoint. Indeed, had the batch cell been run a t a temperature equal to that of the continuous run, its efficiency would have been considerably higher and its power cost lower. The current employed in operating the flowing cell shown was 66.5 amperes a t a voltage of 5.2. The power cost in terms of kilowatt hours per liter of pickle liquor treated was
The return per kilowatt-hour of electricity is similarly found to be 0.155 pound electrolytic iron plus 0.386 pound sulfuric acid. The lower cost figures of the three-compartment cell, considered in their proper light, only begin t o indicate the advantages to be gained from multiple-compartment cell operation. Multicompartment cells promise to combine the advantages of reduced back diffusion of sulfuric acid with the advantages of stepwise depletion of the catholyte feed liquor. Whereas depletion in one step requires a compromise spacing of electrodes, depletion in several steps would allow optimum spacing for each step in the depletion phase. Stepwise reduction of the catholyte would also serve to increase the average catholyte concentration which would in t u r n reduce the applied potential and lower back diffusion.
and, in terms of kilowatt hours per pound of sulfuric acid recovered,
Progress i s made in studies of stripping of cathodes and control of tree growth
'Oo0
66.5 amperes X 5.2 volts = 2.69 kw.-hr./pound H2S04 58.25 grams watts/kw' 453.6 grams/pound
Perhaps power cost is best illustrated by similar calculations which show t h a t one kilowatt-hour of electricity produced 0.152 pound of electrolytic iron and recovers 0.372 pound of sulfuric acid.
-
66.5 amperes X 5.2 volts X 1 hour 1000 volt-amperes/kw. 23.83 grams 58.25 grams H2S04 453.6 grams/pound Fe 4-453.6 grams/pound
1126
The cathodes used in this experimental work were for the most part relieved of their deposits by dissolving off the iron in dilute sulfuric or nitric acid. Some work was done on stripping off t>heiron, however. Deposits have been stripped effectively from aluminum cathodes that have been given a sodium sulfide dip. They have also been stripped from untreated aluminum and stainless steel cathodes, though less effectively. Graphite-coated stainless cathodes were found ineffective for obtaining easily removable deposits, principally because of the difficulty of obtaining an adherent film of graphite on the stainless steel. A thin, adherent coating of graphite might well be expected to be practical, however. The use of a cathode material, such as tita-
Vol. 47,No. 6
INDUSTRIAL AND ENGINEERING CHEMISTRY
_ _
--
~
ENGINEERING, DESIGN, AND PROCESS DEVELOPMENT
) L
477g,/hr. 15.5 % F e S 0 4 4.0% H2S04
+
19 9.
3 0 % Fez(S041, 14.4%H2SO4
02
_-
-
4 0 4 5 g , / hr.
set it is obvious that the two perform somewhat different functions. Electrolytic treatment regenerates sulfuric acid and recovers electrolytic iron, whereas the lime neutralization process discards potentially valuable materials by expending additional valuable materials. Lime neutralization also involves periodic purchasing costs for lime and sulfuric acid and writ,e-offs for the loss of iron. I n addition, makeup tanks are required for both lime and sulfuric acid, reaction tanks are needed for neutralization, valuable land is required for lagooning, licenseF are needed for dumping, and trucks or other vehicles are required for handling. I n contrast, an electrolytic treatment installat ion mould be a self-contained unit eniploying supply tanks and receiving tanks. The only recurring comparable expense would be for electrical energy. T o determine the dollar comparison, consider a moderate size mill-one that produces 1000 gallons of waste liquor per hour of an average composition (approximately 15.6% ferrous sulfate and 3.9% sulfuric acid). Assume that this volume of liquor is depleted to 3.0% ferric sulfate (2.25% ferrous sulfate). Further, assume a value of 1 cent per pound for recovered sulfuric acid, 2 l / 2 cents per pound for electrolytic iron suitable for alloying into special steels in the electric furnace, or 1 3 / 8 cents per pound as scrap, and a cost of 7/10 cent per kilowatt hour for electrical energy. Then, for batchwise operation, as in run 39, Figure 9, the cost is
A N
54
'
c _
0 D
E
58.1 c _
2 . 4 8 E/hr.
Uf 90 A M R / F T 2
Figure 13.
a
\\
2.85 % FeS04 1.70 lhr. */e H 2 S 0 4
Weight and material balance operation of cell
for continuous
0.771 kw.-hr./liter X 0.7 cent/kw.-hr. = 0.54 cent per liter
Run 8 5
From this may be deducted the value of the sulfuric acid recovered
nium, to which deposits do not readily adhere might provide an even better answer to the stripping problem. Aluminum cathodes are considered unsuitable for commercial use, since there is some tendency for the aluminum to dissolve. Stainless steel cathodes, on the other hand, were determined to be usable repeatedly with no detectable loss of weight. Anodes of both chemical lead and antimonial lead were found to be satisfactory chemically. Chemical lead anodes, however, had the disadvantage of being softer and therefore less easily handled. Anodes in all cases were perforated. The elertro-deposition of iron from spent pickle liquor was sometimes hindered by the tendency to grow trees on the cathode. Grid or open mesh cathodes were found to promote the growth of these trees and were therefore discontinued in favor of smooth-faced sheet or plate cathodes. Other factors favoring tree groq-th were found to be low concentrations of ferrous sulfate in the catholyte, high current densities, and low temperatures, although the influence of the last factor is far from conclusive. Considerable progress has been made in the control of tree growth, but complete elimination of these growths has not yet been achieved. The use of multiple-compartment electrolytic cells offers further opportunity for study and control of undesirable tree growth. Costs of electrolytic treatment compare favorably with lime neutralization of pickle liquor
A comprehensive comparison of the electrolytic treatment of pickle liquor with the n-idely used lime neutralization treatment would be complex and very difficult. However, from the out-
134.4 grams X 98'08 X 1 cent/pound 96.06 = 0.312 cent per liter 0.970 liter X 454 grams/pound ~
b B'
E X P L O D E D V I E W S H O W I N G FLOW
m
I
I
I
I I
I /2" B A K E L I T E
4-1/31. NEOP. GASKLTS
-
I/I%'NCOPRENE BISKET
u Figure 14.
Three-chambered cell assembly
INDUSTRIAL AND E N G I N E E R I N G CHEMISTRY
June 1955
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1127
ENGINEERING, DESIGN, AND PROCESS DEVELOPMENT and the value of iron recovered
showing an operating profit of
60.4 grams X 2.5 cents/pound = 0.343 cent per liter 0.970 liter X 454 grams/pound
+ 0.323 - 0.585 = 0.061 cent per liter
or, a comparative advantage of
resulting in a net operating profit of 0.312
0.323
+ 0.343 - 0.54 = 0.115 cent per liter
One liter of 3.9% sulfuric acid and 15.6% ferrous sulfate spent liquor contains
0.178 - (-0.061)
=
0.239 cent per liter
For the 1000-gallon-per-hour mill the operating saving is 3785 X 24 X 0.239 = $217 per day
1000 X 1.204 (specific gravity) X 0.039 = 47.0 grams I&s04 1000 X 1.204 X 0.156 = 187.9 grams FeS04
and
Then, according to the two basic neutralixa-
to be neutralized. tion reactions,
+ CaO + HzO FeS04 + CaO + HzO HzSOc
-+
+
+ H20 CaS04 + Fe(OH)2 Cas04
it follows stoichiometrically that 96.3 grams of calcium oxide would be required, or a t an average of 5% excess, 101 grams. And, if lime having a basicity factor of 0.94 costs cent per pound delivered, lo’
0’75 0.94 X 454
=
0.178 cent per liter
the difference between electrolytic treatment and lime neutralization is 0.115 - (-0.178) = 0.293 cent per liter For the 1000-gallon-per-hour plant the figures show an advantage in favor of electrolytic treatment equal to 1000 gallons/hour X 3.785 liters/gallon X 24 hours/day X 0.293 cent/liter = $266 per day
If the Amberplex process is credited with iron a t only the scrap value of la/, cents per pound, the operating advantages would still be $126, $67, and @5 per day, respectively. Capital investment for continuous two-chambered cell i s comparable to lime neutralization plant
The capital cost of a cell or cells to handle 1000 gallons per hour of liquor depends to a large extent on the current density employed. T h e higher the current density the greater is the amount of liquor treated by one unit cell. Since the power requirements also rise as the current density is increased a balance must be made between power and rectification equipment on the one hand and cell cost on the other. It is convenient to base cell cost on square feet of membrane area; therefore, the membrane area requirements are first determined. I n the batch cell illustrated in this discussion the membrane area was 0.672 square foot. Conversion to the area required to treat 1000 gallons per hour results in 0.672 square foot X 2.99 hours X 1000 gallon/hour X 3.785 liters/gallon 0.970 liter ‘78 $0 square feet
0.871 kw.-hr./liter X 0.7 cent/kn.-hr. = 0.61 cent per liter
At a cell cost of $35 per square foot of membrane area, which is probably a good estimate, this would mean an investment of $275,000. Power rectification equipment a t $85 per kilowatt of capacity would add t o this cost
less, for recovered sulfuric acid,
0.771 kw.-hr./liter X 3785 liters/hour X $85/kw.-hr. = $248,000
Similarly, the cost of continuous treatment in accordance m-ith run 85 in Figure 13 gives
58.25 grams H2S04 X 1 cent/pound = 0.324 cent per liter 0.397 liter X 454 grams/pound and for electrolytic iron, 23.8 grams F e X 2.5 cents/pound = 0.331 cent per liter 0.397 liter X 454 grams/pound
If a batch cell down time of 8% for draining and refilling is allowed, the total investment in equipment would be $545,000. The continuous two-chambered cell, by comparison, employed a membrane area of 0.738 square foot. For 1000 gallons per hour, cell investment would be 0.738 X 1 X 1000 X 3.785 0.397
The operating profit is 0.324
+ 0.331 - 0.61 = 0.045 cent per liter
Neutralizing this liquor with lime would cost the same as in the former illustration-namely, 0.178 cent per liter. Thus 0.178
- (-0.045)
=
0.223 cent per liter
For a 1000-gallon-per-hour plant the calculations show an operating saving of 1000 X 3.785 X 24 X 0.223 = $202 per day The cost of treatment using the three-component cell is 0.836 kw.-hr./liter
X 0.7 cent/kn-.-hr. = 0.585 cent per liter
less, for recovered iron 30.3 grams F e X 2.5 cents/pound = 0.323 cent per liter 0.517 liter X 454 grams/poundand, for recovered sulfuric acid 541 grams X 0.14 X 1 cent/pound = o.323 cent per liter 0.517 liter X 454 grams/pound
1128
=
7050 square feet
and 7050 square feet X $35/square foot = $247,000 Again, power rectification a t $85 per kw.-hr. would be expected to add
0.871 kw.-hr./liter X 3785 liters/hour X $85/kw.-hr. = $280,000 And since no down time need be allowed for draining and filling in this case these figures may be combined to give a total equipment cost of $527,000. T h e estimated cell cost figure of 835 per square foot of membrane area is reduced to $22.50 per square foot with the introduction of an additional chamber and membrane. Thus, the investment using three-chambered cells may be calculated to be 0.738 X 1 X 1000 X 3.785 X 2 0.517 10,800 square feet
x
=
square feet
$22.50/square foot = $243,000
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
Vol. 47, No. 6
ENGINEERING, DESIGN, AND PROCESS DEVELOPMENT plus power rectification investment 0.836 kw.-hr./liter X 3785 liters/hour X $85/kw.-hr. = $269,000 results in a total investment in equipment of $512,000. With an allowance of an estimated $100,000 for instrumentation, tanks, and piping, and another $50,000 for building, including lighting and heating, the total capital outlay for a continuous two-chambered cell installation would be $677,000, a figure which, together with the figures reported previously, may be resolved into a cost-per-gallon figure. Thus, an hourly output of 1000 gallons for a 350-day working year, amounting to 8,400,000 gallons, gives the following breakdown of costs:
Electrolytic cells Rectification equipment Building Tanks, piping, etc. Interest on investment
Depreciation/ Year, % 10 6 3 20 6
Less operating profit (0.045 cent/liter)
cost, Cents/Gallon 0.294 0,200 0 0179 0,238 0.491 1 ,2409 0.1704 1.071
Labor and maintenance would bring the cost to approximately 1.4 cents per gallon total. The figure of $677,000 is comparable to the capital investment of a lime neutralization installation as given in a recent private communication. This particular installation includes a sufficient number of holding tanks to permit satisfactory flexibility of operation for a contract treating organization and is reported to
Aid to Blast Furnace Studies
have cost $2,000,000. The plant produces 60,000 to 75,000 gallons of waste liquor per day. In addition to this investment the company pays 1.8 cents per gallon to the contract treating organization to haul and dispose of their waste liquor. Other plants are reported by the same authority to contract their disposal a t 2 cents per gallon. A figure of 1.4 cents per gallon, then, should be attractive to a company that is paying 2 cents per gallon for waste disposal plus the write-off on what investment it has in holding tanks. Electrolytic regeneration of waste liquors should also be considered where ground for expansion of productive capacity is limited, for expansion of facilities could be made over ground formerly reserved for lagooning. Significant reductions in operating power requirements are foreseen through modifications in cell design and methods of operation. These modifications, which would promote higher efficiency, would also mean lower capital investment for cells and rectification equipment. Even without these certain reductions i t may be said that electrolytic treatment of waste sulfate pickle liquor using Amberplex membranes holds promise, not only because it completely eliminates the need for disposal, but because it competes economically with present processes. Acknowledgment
The authors wish to acknowledge the assistance of J. Lirio of the Rohm & Haas Co. during the experimental program and the mechanical department of the Rohm & Haas Go., Research Laboratory, for their assistance in developing and fabricating the electrolytic cells. RECEIVED for review September 15, 1954.
A 4 c c E ~ TFebruary s~ 5, 1955.
.. .
Air Flow in Beds of Granular Solids J. B. WAGSTAFF
AND
E. A. NlRMAlER
Fundamental Research laboratory, United Stater Steel Corp., Kearny,
I T T L E is known about the flow pattern of the gas in the blast furnace except that it is extremely complicated. The air is blown in through a number of tuyeres around the base with sufficient velocity to form a turbulent recirculation zone (6, 14). Both the solid and gas are in motion and probably neither move uniformly past any horizon. This article is a report of a study of the flow of air through beds of granular solids, in order to determine the correct criteria for an investigation of the blast furnace process by means of models. Bennett and Brown (1) found that when there was relative movement in a bed of granular solids, a region of loose packing developed. It is also known that gas tends to flow,preferentially along a region of loose packing. Therefore, it seems highly probable that the gas and solid flows interact. As a result, resewch on flow in a blast furnace must be done either on an actual furnace, as by Kinney and coworkers (11), or on models in which both solid and gas are in motion. Work on an actual furnace is extremely difficult and the use of models would seem preferable, provided that models are so constructed that the flow pattern is the same as in the full scale unit, a t least in its principal features. It is probable that in an actual furnace the formation of liquids presents additional complications, but in the initial investigations reported in this article such effects are ignored. The literature on gas flow through granular beds is extensive. The greater part of the work reported has been done by measuring June 1955
N. 1.
the pressure drop through columns of presumably homogeneous packings of granular materials. I n most cases, the columns were less than 6 inches in diameter and the granular materials were stationary. I n spite of these restrictions, there are remarkable areas of disagreement. Therefore, the work of five different authors was compared with measurements made in the laboratory in order to determine which of the recommended formulas best indicated the scaling factors necessary in the model experiments. The comparison has been made from the point of view of both the ease and accuracy with which a pressure drop may be calculated if the properties of the packing are known, and, conversely, the accuracy with which the properties of the packing may be determined if the pressure drop is known. The second approach is important when rough and porous materials, such as coke, are used, because such simple properties as particle surface and particle density are not easy to determine. The investigations of Chilton and Colburn ( 5 ) , Carman ( 4 ) , Brown (6),Leva and coworkers ( l a ) , and Ergun (9) were considered in this study. Chilton and Colburn in 1931 correlated the results of earlier workers on the basis of a modified friction factor and a modified Reynolds number. I n this way they deduced equations for the viscous and turbulent regions, respectively. They obtained for the viscous region
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
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