ENGINEERING AND PROCESS DEVELOPMENT
Recovery of 2,3=Butanediol from Fermented Beet Molasses Mashes J. A. WHEAT Division o f Applied Biology, National Research laboratories, Offawa, Ont., Canada
- B:
u
a
CAUSE of the relatively high boiling point of 2,3-butanediol approximately 180' C.), the principal problem i n its recovery from fermented mashes is the separation of the diol from nonvolatile mash solids. The initial steps of all the recovery processes that have been reported are similar (3, 18, 33) and consist of removal of ethanol, filtration or screening if insoluble solids are present, and concentration by evaporation. Various methods have been suggested (3, 33) for the recovery of butanediol from the resulting evaporator sirup. Othmer et al. (18) presented pertinent solubility data and suggested several processes for solvent extraction of butanediol from corn mash sirups. Freeman and Morrison (10)gave laboratory data for the recovery of diol from molasses sirups by precipitation of the solids with ethanol and lime. A process based on this method has been patented by Walmesley and Davis (28). Blom et al. ( 3 ) investigated continuous steam stripping of butanediol from wheat mash sirups on a pilot plant scale. This process has been used with wheat and barley mash sirups in the National Research Laboratories (25, 33). This paper describes the steam stripping of diol from molasses sirups. Molasses mashes fermented by Aerobacter aerogenes and Pseudomonas hydrophila (31) contained butanediol, ethanol, acetoin, water, impurities, and nonvolatile solids. The mash was stripped of ethanol, evaporated t o a sirup containing 2070 diol, and the diol was removed from the solids by steam stripping a t 40 pounds per square inch in a packed column. The diol-containing vapor from the stripping column was scrubbed with water fn a similar column to yield a n aqueous solution containing about 2401, diol. Details of the equipment and procedures used to concentrate and purify the diol have been given in a n earlier paper (30). However, impurities analyses and balances were not included but are presented in this paper. Ethanol purification was not investigated, although methods for its purification when recovered from whole-wheat mashes have been reported ( 2 4 ) . Recovery of acetoin from the mash and recovery of by-products from the solids were not investigated. The latter problem has been extensively studied by others-e.g., Owen (19)and Reich (21). The aims of the pilot plant investigations reported i n this and an earlier paper (31) were to obtain sufficient data for the design of a commercial or semicommercial plant and to make an estimate of the initial cost of a commercial plant and the cost of producing 2,3-butanediol. Based on the experimental results, a quantitative flow sheet for a hypothetical plant with a capacity of 60,000 pounds of molasses per day has been prepared. A cost estimate ( 3 2 )for such a plant has been corrected t o 1952 prices and is summarized in this paper. 4
Pilot Plant Scale Equipment Is Used for Process Study
Most of the experimental data reported in this paper were obtained by the recovery of diol from mashes of the first six runs described in a previous paper (91). Mashes for these runs were
November 1953
prepared from 1949 Chatham, Ont., molasses and were fermented b y A . aerogenes. The amounts of materials used in preparing and fermenting the mash and the average composition of the fermented mash are given in Table I. Some results of earlier runs, in which mashes were prepared from 1948 molasses and fermented by A . aerogenes, were also used, principally for calculating average distributions of impurities. Although diol was recovered from mashes fermented by A . aerogenes only, the methods are applicable t o mashes fermented by 2.' hydrophila, since the composition of fermented mashes was essentially the same (31).
Table 1.
Fermentation Conditions and Results Experi-
Values Quantitative Used fof
Basis 1000 Ib. molasses
Material mentala Flow Sheet Acetic acid (SO%), 1. 6.8 7.0 3.33 Triple superphosphate, lb. 3.33 0.67 0.67 Corn oil 1. 11.4 10.2 Ammonium hydroxide (26'j, 1. 1000 lb. molasses, Diol Ib. 176.7 174.8 7.4 6.4 50% invert Aceioin, lb. 40.9 42.7 sugar Ethanol lb. 4.62 4.4 100 Ib. diol Acid ( a i acetic acid), Ib. Total' neutralizinz eauiv. (as 18.94 acetic acid), 1b.18.9 2.05 Ammonia, Ib. 1.8 3.6 3.4 1000 gal. mash Aeration, cu. ft./min., 0-24 hr. 94-36 hr. 1.8 2.0 -. Sugar fermented, % 98.5 98.6 Efficiency, yo 90.0 89.6 4.50 4.63 Ratio, (diol plus acetoin)/ethanol 6.2 6.2 Initial p H 5.82 Final pH 6.01 1.042 1.045 Initial sp. gr. 1.016 1,016 F/nal sp. 3.76 3.71 Final eoli%L: % a Avesage of runs 1-6 of Table I1 (32). ~
Analytical Methods. The analytical methods reported by Neish (16) were used for the determination of diol and ethanol. For materials containing solids (evaporator sirup, stripper residue, and distillation residues), samples were weighed out and after suitable dilution diol was extracted with 1-butanol. The acetoin content of recovery samples was neglected. Solids were determined by evaporating samples to dryness a t 110' C. or to a thick sirup and determining the diol content of the residue. Free acid content was determined by titrating R sample t o p H 7.5. Total neutralizing equivalent was determined by adding an excess of sodium hydroxide to the sample, refluxing for 2 t o 3 hours, and back titrating to a pH of 7.5 with standard acid. Concentrations of acid and total neutralizing equivalent were calculated as weight per cent of acetic acid. Ammonia was determined by adding an excess of sodium hydroxide to a sample, distilling the ammonia into an excess of standard acid, and back titrating. Specific gravities were measured with narrow range hydrometers. Flow Sheet. Figure 1 is a flow sheet for the recovery process, except for the purification steps, This drawing does not show all the piping nor all the storage and measuring tanks. Only a few of the indicating instruments have been included.
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
2387
ENGINEERING AND PROCESS DEVELOPMENT
Figure 1 .
Flow Sheet for Butanediol Manufacture
P = Pressure PD = Pressure difference I = Indicator I I = Liquid level R = Recorder C = Controller F
= Flow
Figure 2 is a general view of the operational floor of the pilot plant. At the extreme left is the ethanol stripping column. The stripping and scrubbing columns are in the middle, behind the steel column and in front of the 1000-gallon slops tank. The vessel with the top-entering agitator to the right of the instrument panel is the fermentor. On the upper floor, the continuous cooker coil and its instrument panel can be seen. Ethanol Stripping. The ethanol stripping column and method of operation have been described (33). Briefly, the feed rate to the column, which was 2 feet i n diameter with 22 plates on 17inch spacings, was measured with a n orifice meter and automatically controlled. Steam t o the coil was regulated by the pressure drop across the column and the take-off was adjusted manually according to the top temperature of the column. Slops discharge was controlled by the liquid level a t the base of the column. Evaporating. Mashes were concentrated in a stainless steel, standard vertical-tube evaporator with 160 tubes, 1'/z inches in outside diameter by 30 inches long, giving a liquid side heat transfer area of 137 square feet. The evaporator head was equipped with a manhole, pipe nozzles, and two window t o observe foam. A liquid level recorder controller was used to control the discharge of sirup, which was sent t o a steel storage tank not shown in Figure 1. Vapor from the evaporator was scrubbed of diol in a 1 foot in diameter, copper rectifying column containing 33 plates. Steam to the evaporator was controlled by the pressure drop across the column, using the same instrument that controlled steam t o the ethanol stripping column. Distillate was pumped through reflux and take-off rotameters; take-off was
2388
discarded after measuring and sampling in an open tank not shown in Figure 1. I n evaporating a batch of mash or slops, the feed and distillate take-off rates were adjusted to maintain a constant level in the evaporator. When the desired diol concentration of about 20y0 was reached, generally about halfway through a n 800-gallon batch, sirup was removed continuously. rlfter all the feed had been added, the remaining sirup was transferred with air pressure and the evaporator washed with hot water. Some mashes were also evaporated in the smaller unit described previously (SO). Stripping-Scrubbing. Tn o columns, 1 foot in diameter and 20 feet high, packed with about 18 feet of 1-inch Raschig rings, served as stripping and scrubbing rolumnq. Sirup was pumped by a chemical proportioning pump from m t o 30-gallon measuring tanks through a heater to the top of the stripping column. +4 variable-speed rotary pump ' i ~ as used t o add water t o the scrubbing column. Liquid discharge fiom the bottom of each column was regulated by a liquid level controller and the discharge was passed through a water cooler. Residue was weighed in tubs, sampled, and discarded, uhile scrubbing column product was collected in a 200-gallon tank not shown in Figure 1. Steam flow to the stripping column was automatically controlled. Vapor from the scrubbing column, whieh was controlled a t 40 pounds per square inch gage, \%ascondensed and collected in an open tank where it was measured, sampled, and discarded. T o start the stripping-scrubbing unit, the steam rate and pressures were established before the liquid feeds were introduced. After an hour or more, after the s? stem had come to equilibrium,
INDUSTRIAL A N D E N G I N E E R I N G CHEMISTRY
Vol. 45, No. 11
ENGINEERING AND PROCESS DEVELOPMENT readings and samples were taken over two successive 1-hour periods. One or more rates were changed and, after another stabilizing period, two more 1-hour runs were made. At the end of each day’s operation the stripping column was washed by running warm water through i t for about 20 minutes and then filling it with water and sparging steam a t the bottom for 5 t o 10 minutes after the water reached the boiling point, The column was then drained and rinsed.
The impurities in the distillate varied over a wide range, primarily because of differences in the p H of the feed. The extent of fractionation appeared t o have little effect on the distribution of impurities since nine additional runs gave essentially the same average distributions as shown in Table 11, even though the reflux ratio was such that the diol content of the distillate was increased by a factor of 10. The p H of 14 evaporator feeds varied from 4.59 to 6.30 and the respective distillate and sirup pH’s were as follows:
Diol Is Steam Stripped from Molasses, Scrubbed, and Recovered by Distillation
*
*
Ethanol Stripping. About twelve batches of mash were put through the ethanol stripping column (all were earlier runs than those for which fermentation data have been presented) and readings and samples nere taken over 3-hour periods of continuous operation during two of the runs. All other batches of mash were evaporated without recovering ethanol. Although the ethanol stripping column performed satisfactorily when wheat mashes were processed (SS), calculations showed that with molasses, the fifteen stripping and seven enriching plates gave 0.3 and 2.6 theoretical plates, respectively. It is believed that the performance was due t o the fact that the plates were not level and t o misalignment of plates and liquid downcomers. Because of the time required to dismantle the column and reassemble it, correction of this difficulty was not warranted
The average distillate p H for the five runs in Table I1 was 8.31 and the average sirup p H was 6.75. Sirup was usually concentrated t o about 20% diol with a corresponding solids concentration of 28 t o 30%. More concentrated sirups were prepared and little difficulty was experienced in piping them, but a standard concentration of 20% diol was selected. No difficulties were encountered in the use of proportioning, gear, and centrifugal pumps. Heat transfer to boiling sirup and fouling of evaporator surfaces have been discussed (30). The viscosity of one of the early sirups, which had a lower diol to solids ratio than shown in Table I, was measured at various solids concentrations and temperatures. The results indicated that the viscosity of sirups containing 20% diol and 28 t o 30% solids would be about ten times the viscosity of water at the same Table II. Material Distribution for Evaporating temperature. (Per cent of material in feed) Stripping-Scrubbing. Total material balances were obtained Total Total Neutralizing for the stripping-scrubbing unit by measuring the amounts of Material Diol Acid Equiv. Ammonia sirup, residue, water, condensate, and product over the 1-hour 17.9 44.0 0.68 0.0 Distillate 85.92 periods of continuous operation. These rates were used to calSirup 11.82 92.48 107.9 5 2 . 7 45.2 93 16 107.9 70.6 89 2 Totalout 97.74 culate vapor and steam rates in the stripping column. The flow rates a t the bottom of the columns were about 32 pounds per (hour)(square foot) greater than at the top because of vapor Evaporating. Average material balances for five evaporations condensed by heat loss and cold feed. A feed rate of 280 pounds are given in Table 11. The 6.84% loss of diol was caused per (hour)(square foot) a t 100’ C., which was about the maxiprincipally by a n incomplete transfer of sirup from the evapomum temperature t o which sirup and water could be heated, rator. No decomposition of diol was detected in laboratory tests, would condense about 22 pounds per (hour)(square foot) of in which samples of sirup were refluxed at alkaline, neutral, and vapor. Hence, heat loss amounted t o 10 pounds vapor per (hour) acid pH’s for 72 t o 96 hours. Further evidence that the diol (square foot). loss was not due to decomposition is the fact that diol recoveries Over both columns the diol nearly always balanced t o within for six runs in the purifying unit ranged from 96.3 t o 105.2%, 5%. If i t did not, the results were discarded. When the diol with an average of 99.5%. Total material and diol balances for balance was within 5%, i t was corrected by adjusting the diol single-stage and three-stage evaporation of mash in the purifying concentration of the sirup, since repeated analyses of several sirup unit have already been reported (90). samples showed that the precision of the analysis was about 5%. The diol lost t o the residue varied from 1.1 t o 11.0% of the diol in the sirup, with an average loss of 4.56%. Sirup rates were varied from 172 to 290 pounds per (hour)(square foot) and steam to sirup ratios from 2.46 t o 3.65. It was not possible t o calculate mass transfer coefficients or mass transfer unit heights because the effect of solids on the vaporliquid equilibria was not known. When the solids were assumed t o be inert and the diol concentration of the sirup calculated on a solids-free basis, the actual vapor concentration a t the top of the stripping column was higher than the concentration in equilibrium ( 2 ) with the sirup in 35 out of 49 sets of data. This indicates that the solids combined with a portion of the water t o give a n effective diol concentration higher t h a n t h e solids-free concentraFigure 2. Operational Floor o f Pilot Plant for Butanediol Production tion.
November 1953
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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ENGINEERING AND PROCESS DEVELOPMENT
I
' 7
rl
II
I!
t
I \
/FLASH
Y30.
COOLER
SEVEN
FERMENTORS
D. 18. U - 18. 30,000 go1 C d - 2 4 6 5 f t I I/2'00,968,qfI r \ p ~ l v l o r - 3 O ' D , 2 O h p 152 r p m
I
E.
Figure
To test this point, vnpoi-liquid equilibria deterininirtions weie made with sirup. An attempt to use the apparatus described by Tomkins et al. ( 3 6 )a t 40 pounds per square inch was not successful because the foam formed by the boiling sirup could be neither observed nor controlled. Instead, a glass 0thmr.i still, in which sirup was heated in a circulating leg, was used a t atmospheric pressure. Diol and solids concentrations were varied independently from 5 to 44qc diol and from 14 to 35y0 solids. The effective liquid concentration that would be in equilibrium with the vapor was determined fioin vapor-liquid equilibria data obtained in the same apparatus with diol-water solutions. The data were best correlated (correlation coefficient of 0.937) by the equation
D* _ - 0.03733 S D
+ 0.6631
(1)
where D * / D is the ratio of the effective to the actual liquid concentration, in weight per cent, and S is the per rent solids. From the relationship
where DS is the weight per cent diol on a solids-free basis, the ratio, D*/Ds, can be computed. Over the range of solids concentrations tested, this ratio was greater than unity, showing t h a t the solids were combined with water. W7hen Equation 1 was applied to stripping column data, the vapor concentration was still above the equilibrium concentration of the sirup in 25 out of
2390
I I
I I I I
3. Flow Sheet for Plant Processing
the 49 sets of data. Apparently, under the conditions a t the top of the stripping column, the ratio, D * l D , was greater than in the Othmer still. Solids in the sirup also introduced the problem of possible accumulation of solids in the packing and subsequent plugging and flooding of the column. With a wheat sirup containing 23.69h solids it was found ( 5 3 ) that at a sirup rate of 160 pounds per (hour)(square foot) and a steam t,o sirup rat,io of 3.81, the stripping column flooded after 6.2 hours of operation. With molasses sirups the stripping column never flooded. Throughout each day's operation, the pressure drop changed a small amount but showed no upward trend. It varied from day to day from a minimum of 0.3 inch of water per foot of packing t o a maximum of 1.7, with an average of 1.0. Although the pressure drop indicated that solids were not accumulating in the column, solids balances showed almost consistent losses ranging from 0.7 to 14.67,, with an average of 8.070. Early in the investigation, a direct test was made t o determine how long the stripping column could be operated. Sirup containing 16.9y0 diol and 52.3% solids was stripped in a 6-inch stripping column (33) at a rate of 205 pounds per (hour)(square foot) for a period of 80 hours. The solids loss over each 8-hour period was uniform, averaging 9.4y0, but' the pressure drop, although it varied throughout the run, showed no general trend. Hence, it was concluded that' the stripping column probably could be operated without cleaning for a week, and possibly for a much longer period. The consistent loss of solids might best be
INDUSTRIAL AND ENGINEERING CHEMISTRY
VOl. 45, No. 11
ENGINEERING AND PROCESS DEVELOPMENT
PURlFlOATlOH
f'
UNIT
I
~ D i l i i l l o t .
temperatures are A H = Height, ft. D = Diameter, ft. A P = Number of plates U
60,000 Pounds of Molasses per Day
4
ascribed to errors in the analytical method for the determination of solids. From the scrubbing column material balance data, the height of a transfer unit, H O G , based on over-all gas-phase concentrations was calculated (20)for each 1-hour test. Table I11 gives the averages of the tests made during each run and the averages of all tests. The heights required for the individual film resistances were calculated (20,34) to be 0.04 foot for the liquid film and 2.51 feet for the gas film, showing t h a t the resistance of the gas film is 98.5% of the total. Table IV gives the average impurities distribution for the stripping-scrubbing operation. The variations from run t o run were not large, considering the analytical methods and the differences in the sirups and operating conditions. Average pH levels are given in Table VI.
Table 111. No. of Tests 5 7 7 5 3 4 4 8 6 Av. of 49 tests
Vapor Lb./(Hr.) ( S q . Ft.)
November 1953
Scrubbing Column Data
F.; = = =
pipe sizes indicated near Row lines Heat transfer area, sq. ft. Temperature difference, Fa Heat transfer coefficient, B.t.u./(hr.) (sq. ft.)
(OF.)
Purifying. The following steps were used t o concentrate and purify the butanediol in the scrubber product (SO): 1. Sodium hydroxide was added t o the scrubber product and a portion of the material was distilled off and fractionated at atmospheric pressure into a water fraction and a crude diol containing 50 to 60% diol. A residue was removed from the evaporator. 2. The crude diol was fractionated under vacuum into water and im ure diol. 3. $odium hydroxide was added to the impure diol and the product was vacuum distilled, leaving a second residue. The amount of sodium hydroxide added was 5y0in excess of the amount equivalent t o the total neutralizing equivalent of the material treated. Except for impurities analyses, which are given in Table V, all data for these three steps have been presented elsewhere (SO). The results for step 1 in Table V are the means of two runs, for step 2, the means of five runs, and for step 3, the means of seven runs. Impurities analyses are given rather than distributions, as was done for other operations, because the
1
Water/ Vapor
Diol Concn. in Vapor,
%
Diol Lost i o
HOG, Condenaate, Ft.
%
3.25 3.67 3.77 3.57 3.50 3.60 3.97 4.72 3.87 3.84
4.57 6.03 6.71 6.82 7.53 6.84 9.32 4.40 7.59 6.41
Table IV.
Impurities Distribution for Stripping-Scrubbing
Residue Product Condensate Total out
(Per cent of material in sirup) Total Neutralizing Acid Equiv. 0.0 31.1 64.1 61.2 46.5 41.7
110.6
INDUSTRIAL AND ENGINEERING CHEMISTRY
134.0
Ammonia 35.1 62.3 20.8
118.2 239 1
ENGINEERING AND PROCESS DEVELOPMENT Table V. Step
Impurities Analyses for Purification
Material Scrubber product Distillate Crude diol Residue Crude diol Distillate Impure diol Impure diol Product Residue
PH 5.08 9.57 5.57 8.65 8.31 8.38 6.10 6.00 8.53 10.79
Lb./100 Lb. of Diol Total neutraJising dcid equlv. Ammonia 9.66 1.21 5.14 .... 0.0 ... 0.046 0.17 0.77 0,016 0.35 0.0 0.30 1.24 0.0 10.0 36.6 0.0 0.01 0.33 0.177 0.34 0.01 0.171 0.13 0 0 0.0 0.23 0 04 0.0
analyses showed closer agreement from run to run. The composition of the feed for one step is not the same as the product of the preceding step because the data are taken from different runs. The impurities content of scrubber product is not representative because impurities increased on storage. In commercial production, the residues would be recycled t o the stripping column; all solids would then be removed from the process in the stripping column residue. Equipment and Floor Space Are Estimated for Processing 60,000 Pounds of Molasses per Day
column because of the foaming characteristics of the sirup. For the purifying steps, experimental values were modified, where necessary, to make them consistent. Impurities analyses listed in Table VI are lower than the experimental values in Table V because it was assumed there would be negligible storage time for impurities to increase. A complete set of calculations was made for the process and the calculation was then repeated, taking the recycle into consideration. The second recycle was taken to be the same as the first so t h a t one repetition of the calculations would suffice. It was assumed that the two molasses tanks would be outdoors. Other raw materials, final product, and large spare pieces of equipment could be stored in a building 64 X 34 X 18 feet. A space 126 X 54 X 36 feet M-ould be needed t o house the fermentors, beer well, and slops tank. All remaining equipment, offices, laboratory, locker room, and control room would require ahout 90,000 cubic feet. A space 120 X 220 feet would be ample for the buildings, molasses tanks, and a steam plant if required. Initial Plant Cost Is $2,780,000, Based on June 1952 Prices
To reduce corrosion losses and contamination of the product, liberal use of stainless steel and stainless-clad steel was used in the cost estimates. Of the main items of equipment, fabrication of stainless steel was assumed for the acetic acid feed tank, sluiry tank, propagators, ethanol stripping column, evaporator tubes and columns, scrubbing column, and most of the purifying unit. Stainless-clad steel was assumed for the flash cooler, fermentors, evaporator bodies, parts of the purifying unit, and final product storage tanks. Most of the pumps and piping Iyere assumed t o be stainless steel. The Engineering Sews-Record construction cost index and the iVIarshall and Stevens index (23) were used to correct all cost data t o June 1952. The cost of most of the equipment map estimated from published data (1, 6, 8, If?, 13, 22, 27), but the cost of some small items was merely guessed on the basis of prices for pilot plant equipment. The cost of fermentors and propagators was estimated from the weight of t h r vessels ( 9 , 11). Where possible, data giving the installed cost of equipment were used but where such data were not available, the cost of the basic equipment item was multiplied by 1.43 (16) t o obtain the installed cost. Installed cost includes instdlation labor, foundations and supports, installation of drive equipment, insulation, and painting. I n the second column of Table VIII, which is a summary of the initial cost estimates, each item is expressed as a percentage of the total cost. These percentages are compared with those given by Lang ( 1 4 )for fluids processing plants costing$1,000,000 to $5,000,-
A quantitative flow sheet and cost estimates were prepared for a. hypothetical plant with a daily capacity of 60,000 pounds of molasses. This capacitl- is approximately equivalent in carbohydrate content t o the capacity used for estimating the cost of production from wheat (33). The flow sheet is presented in Figure 3 and in Tables V I and VII. A total capacity of 30,000 gallons and a fermenting capacity of 20,000 gallons were selected for each fermentor. .4t a mash rate of about 40,000 gallons per day, two fermentors could be filled in 24 hours. A fermenting time of 48 hours was required (31) for maximum yield, but the last 12 hours, during which there was no aeration and only a slight increase in yield, could be provided for in the beer well. T o give a fermenting time of 36 hours, a BO-hour cycle would be required for each fermentor. Thus, a minimum of five fermentors would be needed, but seven were included in the cost estimate. The size of the propagators (Figure 3) would give the same proportion of inoculum to mash used in the pilot plant (31). The propagator cycle would be 36 hours, requiring a minimum of three vessels, but seven were included in the estimates. Diol lost in each recovery operation was set a t a low amount without calculating economic optimum recoveries. Reflux ratios were also chosen arbitrarily. Using a Murphree plate efficiency ($0)of 60% for plates and 80% for stills and reboilers and the vapor-liquid equilibrium data of Blom et bZ. (f$, the number of plates for each distillation colTable VI. Composition of Materials Containing Diol umn was calculated. Allowable (Basis, 24 hours) vapor velocities were determined Weight, L b . from the relationship given b>Total Perry (20), assuming 18-hch neutralizing Vol., Gal. Sp. Gr. Total Diol Solids Acid equiv. plate spacings and ll/*-inch liqMash 39,250 1.016 398,800 10,603 15,000 490 2008 uid seals. Over-all heat transfer 38,780 2003 390,800 10,603 15,000 489 Slops 1.016 Distillate 34,150 359 338,300 103 coefficients were taken from ex0 0 I * ooc 1055 Sirup 4,713 1.114 52,500 10,500 15,000 528 perimental results or from pub17 Recycle 483 1.064 5,137 1,236 3,136 0 Total sirup 1072 5,196 57,637 13,636 16,236 528 1 109 lished data (20). The diameter 333 Residue 5,349 59,695 100 16,236 0 1.116 Scrubbed vapor 447 .... 0 245 of the stripping column was cal178,542 100 656 1 oi2 56,406 13,436 Scrubber product 5,574 0 339 culated from the pilot plant sirup 0 ..... Scrubber product Aa 5,658 1 014 57,368 13,436 1,180 n 31 Distillate 1 3,153 1 000 31,530 0 24 4 rate of 280 pounds per (hour) 12,5 56.7 0 Crude diol 2,054 1 021 20,976 10,488 16.7 0 2,917 1,180 1 065 (aquare foot) and, since it was not Recycle 1 457 4,862 14.8 1 000 Distillate 2 1,035 10,348 73 0 0 possible to calculate transfer unit 24.2 12.9 Impure diol 1,063 1 000 10,628 0,415 0 0 ..... Impure diol A5 1 000 1,066 10,664 10,415 56 heights, the height was assumed 1 060 275 219 56 0 0 Recycle 3 26 10.2 I 000 10,389 10,196 0 0 t o be 50 feet. It was thought Final product 1,039 advisable to use a packed stripAfter adding sodium hydroxide. ping column instead of a plate
Ammonia 217 216.9 96.4 98.1 0.7 98.8 34.7 20.6 61.6 ,..
69.5 3.3 0.7 2.5 0.7 '0' 0
pH 5 82 5.82 8.31 6.75 8.80 6.93 8.00 4.45 5.08 10.00 9.57 5.57 8.65 7.40 6.20 9,20 11.20 8.50
Q
2392
INDUSTRIAL AND ENGINEERING CHEMISTRY
Vol. 45, No. 1 1
ENGINEERING AND PROCESS DEVELOPMENT Table VII. Fluid
u
Molasses Triple superphosphate Acetic acid (80%) Water Slurry steam Raw mash Cooker steam Cooker mash Cooler vapor Cooled mash Condenser water Cooler water Ammonium hydroxide (26") Corn oil Gas Fermented mash Ethanol column vapor Condensate, feed preheater Condensate, final condenser Reflux High wine Liquid below feed Vapor below feed Condenser water Steam Slops Second stage feed Sirup Scrubbed vapor Fresh steam Total steam Condenser water
Rates of Flow Lb./Day 60,000
200 990 346,501 18,133 425 824 84,820 510,644 98,280 412,364 1,692,000 468,970 1,354 82 15,000 398,800 88,000 41,100 38,900 80,000 8,000 500,180 109,380 530,770 117,300 390,800 224,445 52,500 178,542 67,858 246,400 1,980,400 232,126 60,180 171,945 224,579 58,224 166,355 ~
Vol. Rate Gal./min. Cu. ft./min.
2.98
.....
0.064 32.6
....
..... .....
29.2
.....
38.1
.....
28.0 117.5 32.57 0.104
..... .....
27.26
.....
3.13 2.97 6.10 0.61 34.19
.....
38.66
.....
26.71 15.21 3.27 .....
.....
.....
137.5
.....
4.18 11.95
.....
4.04 11.55
..... ..... 73.2 ..... 136 .....
5500
..... ..... .....
..... .....
93.1 ..... 1374
..... ..... ..... ..... .....
2040 I
.
473'
..... ..... .....
965 274 1239
2466' ..... ..... 4182 ..... .....
Gal./hr. Fresh sirup Solids recycle Total sirup Steam to sirup Steam to stripper Residue Vapor Water Steam to water Scrubbed vapor Scrubber product 50% sodium hydroxide Scrubber product A Steam to evaporator 1 Vapor from evaporator Solids recycle 1 Vapor from column 1 Reflux 1 Distillate 1 Water to condenser 1 Steam to reboiler Crude diol Vapor from column 2 Reflux 2 Distillate 2 Water to condenser 2 Steam to calandria Impure diol 50% sodium hydroxide Solids recycle 3 Steam to still Water t o eondenser 3 Final product Solids recycle
e
52,500 5 137 57:637 6,946 182,658 59,695 180,600 52,608 1,740 178,542 56,406 962 57 368 53 :308 52,506 4,862 46,712 15,182 31,530 353,992 2,592 20,976 14,448 4,104 10 348 720:OOO 16,565 10,628 36 275 2,874 45,700 10,389 5,137
196.2 20.2 216.6
.....
..... ..... ..... 28 738
.....
222.8
..... .....
219.2
.....
'
975
.....
7.03 965
232.2 2.62 234.8
.....
..... 19.1 ....
215 978 ..... 870 .....
.....
63.3 131.4 1470
.....
85.8
.....
17.1 43.1 3000
.....
44.3 0.098 1.08
.....
43.2 43.3 20.2
...
.....
..... .....
10.5
iiii' .....
..... ..... 67.2
.....
..... 11.6
.....
Table VIII.
Summary of Initial Cost of Plant Cost,
Per Cent of Total
Dollars hand and improvements Buildings Process equipment Process oiDine Instrume&s Electrical installations Service facilities Total physical cost Contingency Insurance and taxes Engineering Total other costs Total cost
51,180 126,550 1,133,620 380.550 109,340 81 ,040 205 370 2,087,650 196,890 59,060 435,830 691,780 2,779,430 ~
1.84 4 5.5 40.79 13.69 3.93 2.92 7.39 __ 75.11 7.09 2.12 15.68 24.89
Lang (141,
% of Total 3 6
31 20
...
5
8 73
-
~
27 100
100.00
.....
.....
000. The equipment cost is higher than that given by Lang because of the extensive use of stainless steel. In general, the cost distribution given in Table VI11 appears t o be reasonable. Daily Production Cost Is Estimated at 25.6 Cents per Pound of Diol
A summary of daily production costs is given in Table IX. Insurance was assumed to be 0.5% and taxes 2% of the total physical cost. Buildings were depreciated a t 2y0 per year and equipment a t 7%. Total energy and labor requirements were estimated as follows: Steam 600,000 pounds per day a t 55 cents per 1000 pounds Water 650,000 gallons per day a t 8 cents per 1000 galIons Electricity 6300 kw.-hr. per day a t 0.5 cents per kw.-hr. Operators 135 man-hours per day a t $1.82 per man-hour Foremen 32 man-hours per day a t $2.28 per man-hour November 1953
The labor costs include an allowance for vacations, pensions, etc. Since the production of diol is 10,389 pounds per day, the production cost is 25.58 cents per pound of diol. Working capital, defined by Vilbrandt ( 2 7 )as the total operating cost for one year, is $797,340. The total investment is then $3,576,770 and a selling price of 48.53 cents per pound would be required t o give a return of 20% on this investment. The effect of changing the plant capacity was estimated by assuming that most of the costs vary as the 0.6 power of capacity ( 7 ) . Diol produced, raw material costs, energy costs, and miscellaneous costs were assumed t o vary directly with capacity. Decreasing the plant capacity to one third of the original estimate would increase the production cost by 25.6% and the selling price for a 20yo return on total investment by 36.3%. Increasing the capacity by a factor of 3 would decrease the production cost by 16.5% and the selling price by 23.3%. Comparative Cost of Production from Wheat. The estimated initial cost of the wheat plant (SS), when corrected t o June 1952 prices, is $865,000, compared with $2,780,000 for the molasses plant. The difference is accounted for by the extensive use in the molasses plant of stainless steel equipment, which costs six t o ten times as much as steel equipment. The initial cost of the molasses plant could be reduced by using less expensive materials of construction. Resin lined, or possibly steel fermentors, might be used and resin linings could be used to advantage elsewhere in the plant. The stripping column could be made of steel, b u t stainless steel is one of the cheapest materials that would withstand the high temperatures and organic acids and ammonia in the evaporator, scrubbing column, and purifying unit. T h e use of plastic pipe and tubing might further reduce the initial cost of the molssses plant. On the basis of the original estimate for production from wheat,
Table IX.
Summary of Daily Production Costs Cost, Dollars/Day
Management Insurance Taxes Depreciation Total fixed costs Molasses Acetic acid Ammonium hydroxide Corn oil Triple superphosphate Total raw materials Labor Steam Water Electricity Total energy >Iiscellaneous Maintenance
106.67 34.79 139.18 466.03 855.90 120.88 75.15 33.69 7.76 ___
746.67
1093,38 320.48
330 00 52.00 31 50
Total operating cost Total production cost
I N D U S T R I A L AND E N G I N E E R I N G C H E M I S T R Y
Percentage of Total Production cost
4.01 1.31 5.24 17.53 32.20 4.55 2.83 1.27 0.29 -
28,09
41.14 12.06
12 42 1.97
413.50 15 00 68.77 1911.13 2657.80
1 17 -
15.56 0 56 2 59
71 91 100 00
2393
ENGINEERING AND PROCESS DEVELOPMENT the production cost of levo-butanediol would be 34.1 cents per pound corrected to June 1952. The selling price for a return of 20% on the investment would be 47.7 cents per pound. If the Name materials of construction were used for the two plants, the selling price of diol from molasses would be considerably lower than the selling price of diol from wheat. Commercial Production May Lead to Expanded Industrial Use of Butanediol
Until recently, 2,3-butanediol has not been available in commercial quantities, so industrial uses for i t have not been developed. However, in 1951 following completion of the work described in this paper, Celanese Corp. of America announced production of synthetic 2,3-butanediol in tank car quantities ( 4 ) at a price of 14 cents per pound (6). The hygroscopic and solubility properties of 2,3-butanediol suggest its use in printing inks and pastes, dyes, soaps, ointments, and wood and leather stains. With polybasic acids, it condenses to form polyesters of the alkyd type ($39). In these more general uses, 2,3-butanediol would have t o be compared with competitive pioducts such 8 6 ethylene glycol, glycerol, and other glycols. Prices reported ( I ; ) for tank car quantities of four such products are as follows: Cents per Pound 16 17 18.5 39.5
1,3-Butanediol Ethylene glycol Cellosolve Synthetic glycerol
Because of its longer chain length, 2,3-butanediol might have greater compatibility or, because of specific properties, result in a better product in some preparations. h more specific use of 2,a-butanediol would be its use as a chemical intermediate. Diacetyl, a food and butter flavoring, could be readily prepared by oxidation. Cyclic acetals and ketals (1 7 ) ,which might have interesting chemical and solubility properties. are another possibility. Acknowledgment
Vapor-liquid equilibria data were determined by V. P. Milo, whose assistance in other phases of the work is acknowledged.
Literature Cited
(1) Bliss, H., Chena. Eng., 54, No. 5, 126-38 (1947); 54, No. 6, 100-2 (1947). (2) Blom, R. H., et al., IND. ENG.CHEM.,37, 870-2 (1945). (3) Blom, R. H., et al., Ibid., 37, 865-70 (1945). (4) Chem. Eng., 58, No. 3, 164 (1951). (5) Chem. Eng. News, 30, 5110, 5210, 5323 (1952). (6) Chilton, C. H., Chem. Eng., 56, No. 6 , 97-106 (1949). (7) Ibid., 57, No. 4, 112-14 (1950). (8) Considine, D. M., Ibid., 56, No. 3, 124-6 (1949). (9) FOX,L. E., Ibid., 54, KO.8, 100-1 (1947). (10) Preeman, G. G., and Morrison, R. I., J . SOC. Chem. Ind. (London), 66,216-21 (1947). (11) Happel, J., Aries, R. S., and Borns, W. J., Chem. Eng., 53, No. 10, 99-102 (1946); 53, NO,12, 97-100 (1946). (12) Jackson, D . H., Ibid., 54, No. 5, 123 (1947). (13) Lang, H. J., Ibid., 54, No 9, 130-3 (1947). (14) Ibid., 54, NO. 10, 117-21 (1947). (15) Ibid., 56, NO. 6 , 112-13 (1948). (16) Neish, A . C., Natl. Research Council Can., Div. dppl. Biology, Rept. 46-8-3 (June 1946). (17) Neish, A. C., and MacDonald, F J . , Can. J . Research, B25, 709 (1947). (18) Othmer, D . F., et al., IND.E m . CHEM.,37, 890-4 (1945). (19) Owen, W. L., Facts About Sugar, 33, No. 6 , 45-8 (1938). (20) Perry, J. H., ed., “Chemical Engineers’ Handbook,” 3rd ed., New York, McGraw-Hill Book Co., 1950. (21) Reich, G. T., Trans. Am.Inst. Chem. Engrs., 38,1049-66 (1942). (22) Sieder, 0. E., Chem. Eng., 54, KO. 5, 117 (1947). (23) Stevens, R. W., Ibid., 54, No. 11, 124-6 (1947). (24) Tollefson, E. L., Wheat, J. A , , and Leslie, J. D., Can. J . Research, F24, 300-10 (1946). (25) Tomkins, R. V., Scott, D. S . , and Simpson, F. J., Ibid., F26, 497-502 (1948). (26) Tomkins, R. V., Wheat, J. A,, and Stranks, D. W., Ibid., F26, 168-74 (1948). (27) Vilbrandt, F. C., “Chemical Engineering Plant Design,” 2nd ed., New York, McGraw-Hill Book Co., 1942. (28) Walmesley, R. A., and Davis, W. R . , U. S . Patent 2,397,065 (March 18, 1946). (29) Watson, R. W., Grace, iK. H., and Barnwell, 3. L., Can. J. Research, B28, 652-9 (1950). (30) Wheat, J. A., Can. J. Technol., 31, 42-56 (1953). (31) Ibid., pp. 73-84. (32) Wheat, J. A., Natl. Research Council Can., Div. Appl. Biology, Rept. 51/1/1, December 1950. (33) Wheat, J. 24~, Leslie, J, D., Tomkins, R. V., hlitton, H. E., Scott, D. S.,and Ledingham, G. A . , Can. J . Research, F26, 469-96 (1948). (34) Whitney, R.P., and Vivian, J. E., Chern. E.ng. P w g r . , 45,323-37 (1949). for review M a y 22, 1953. ACCEPTEDAugust 31, 1953. RECEIVED Issued as Paper 162 on the “Uses of Plant Products,” and as N.R.C. 3099.
Demineralization of Water Electrol and Ion Exchange Processes I.STREICHER
A. E. B O W E R S
AND
Metropolitan W a f e r Districf of Soufhern California, l a Verne, Calif.
R. E. BRIGGS, P . 0 .
BOX
247, San Dimas, Calif.
SCREASIKG demands foi augmented Tater supplies to meet the needs of a growing population and expanding industry and recurrent droughts within recent years in important metropolitan areas have accented the need for efficient methods for producing potable and industrially usable waters from presently unsuitable sources. Among the methods which have aroused considerable interest and discussion in recent months is the electrolytic process for demineralization of water. The publicity accorded to the development of ion evchange membranes (4,5, 8) 23’34
and Langelier’s discussion of electrocheniical processes for desalting sea water (6, 7 ) have done much to renew general interest in this field of investigation. Demineralization of water by the application of direct current in specially designed cells with diaphragms of canvas or similar materials is not new. Bartow and associates (1, 2 ) described an electrolytic apparatus for removing positive and negative ions from mater and listed references t o earlier experiments of this nature performed in this country and in Europe in the middle
I N D U S T R I A L A N D E N G I N E E R I N,G C H E M I S T R Y
Vol. 45, No. 11
