Ind. Eng. Chem. Process Des. Dev. 1004, 23, 733-737
733
Recovery of Ethanol from Fermentation Broths by Catalytic Conversion to Gasoline. 2. Energy Analysis Gary A. Aldrldge, Xenophon E. VeryklO8, and Raja Muthararan' Department of Chemical Engineerlng, Drexel University, Philadelphia, Pennsylvanla 19 104
A process for recovering ethanol from fermentation beers by catalytic conversion to hydrocarbons is described and analyzed in terms of energy requirements. The process thermally integrates an energy-consuming initial distillation step with an energy-producing reaction step resulting in significantly reduced net energy requirements. It is found that the optimum distillation overhead composition is 60 wt % ethanol. The net energy requirement of this process is found to be 1800 Btu/gai of ethanol processed, which compares very favorably with atternatbe separation processes for simply recoverlng ethanol.
Introduction The topic of conversion of plant carbohydrates to ethanol has received widespread attention in recent years. This has occurred as a result of the energy crisis and the realization of the need for dependence on renewable resources. The process economics of fermentation alcohol production center around the energy intensive step of ethanol separation from the fermentation broth (Chambers et al., 1979; Scheller, 1978). Recently we reported the technical feasibility of converting 90% aqueous ethanol solutions to gasoline (Whitcraft et al., 1983). The authors reported that aqueous ethanol solutions can be converted to hydrocarbon fractions in the gasoline boiling range over a shape-selectivezeolite catalyst (ZSM-5) at a temperature in the range of 300 OC and a t pressures between 1 and 8 atm. The product distribution was found to be a function of temperature, pressure, and space velocity. It appears that cocurrent research in Europe also confirm the feasibility of this approach (de Boks et al., 1982; Oudejans et aL, 1982). In this paper we report the energy requirements for converting fermentation derived ethanol to gasoline and compare it with other competing processes such as liquid-liquid extraction, dehydration, and others. Background Various processes for the separation of low grade ethanol/water mixtures containing approximately 10% alcohol have been investigated. These include two-step distillation, solvent extraction (Scheibel, 1950; Boekeler, 1948; Dreyfus, 1936; Leeper and Wankat, 1982), extraction with critical carbon dioxide (de Filippi, 19821, dehydration (Ladisch and Dyck, 1979, Fanta et al., 1980, Pocius et al., 1982) and others. Eakin et al. (1981) reviewed several of the above described ethanol/water separation methods both in terms of energy requirements as well as the available technology. From the viewpoint of energy requirements the carbon dioxide extraction process requires significantly less energy compared to the conventional distillation process (90oO vs. 27000 Btu/gal of ethanol). Even though the capital cost is high, the much lower operating cost could provide a rapid payback. Vacuum distillation can produce 100% ethanol but the energy requirements are much higher, 37 OOO Btu/gal of ethanol. Dehydration employs a solid dehydration agent such as CaO, waste cellulose, calcium sulfate and others, to concentrate the fermentor beer to about 95% ethanol. The saturated dehydrant may be either regenerated or sent to a fermentation step as in the case of cellulose hydrolysis. The energy requirements are estimated to be about 1200 Btu/gal of ethanol in the case 0196-4305/84/1123-0733$01.50/00
of CaO. One disadvantage of this approach might be the cost associated with solids handling on a large scale. All of the methods discussed above deal with physical separation of ethanol-water mixtures to obtain ethanol of a particular purity. Our approach is to convert the ethanol of the fermentor beer to hydrocarbons, which, being immiscible with water, can be easily separated. The question then arises as to what the energy requirements of such a process are. The primary objective of this work is to develop the energy analysis of the ethanol (10%) to hydrocarbons process and to compare the energy requirements with those of other separation processes reported in the literature. It should be emphasized that the ethanol to gasoline conversion is proposed only as a means of separation of ethanol from the fermentor beer. Additionally, this separation scheme has other advantages which include production of a conventional fuel which can be used without modification of existing machinery and transportation via the existing pipeline network.
Integrated Process Configuration A process for the recovery of ethanol from fermentation beer by converting to gasoline was proposed by the authors earlier (Whitcraft et al., 1983). This process integrates an energy-consuming distillation step with an energy-producing conversion step and, as a result, the net energy requirement for the overall process is significantly reduced. The process flow diagram for the recovery of ethanol by catalytic conversion to hydrocarbons in the gasoline boiJing range is shown in Figure 1. The fermentor output, which is assumed to contain 10% ethanol and which is assumed to be the feedstock in this process, is brought to saturation temperature by three stages of heat recovery from the overhead condenser, the column bottoms, and the final product from the separation chamber. The overhead product vapor from the distillation column which contains 60% ethanol is compressed to the reactor operating pressure (8 atm) and raised to the reactor temperature (300 "C). The ethanol vapor in the reactor feed is dehydrated and isothermally converted to gasoline vapor over a suitable zeolite catalyst. In the present study the product distribution obtained over the ZSM-5 catalyst was assumed (Whitcraft et al., 1983). The reactor product is partially condensed and separated. The gaseous fraction which is primarily composed of C1-CI hydrocarbon gases is compressed and recycled back to the reactor for conversion to high hydrocarbons. Thus the product from the reactor consists only of liquid fractions. The liquid fraction is cooled and separated into hydrocarbon and water by de1984 American Chemical Society
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Ind. Eng. Chem. Process Des. Dev., Vol. 23, No. 4, 1984
_..._
1 I J L - I u
0
0
I
--
Figure 1. Ethanol recovery by conversion to hydrocarbons. Process flow diagram.
I8Ooo
t
I000 0.60
0.65 MASS
om
o.n
ow
085
090
5
FRZTION OF ETHANOL IN OVERHEAD PROWCT
Figure 3. Energy required for various distillate compositions from 10 wt % ethanol feed. Basis: reflux ratio is (1.2) (minimum reflux ratio); feed is 10 w t % ethanol at saturation temperature; 99% ethanol recovery in distillate.
!$ I,
\\
F
$
6000
006
008 00 012 014 016 MASS FRACTMN OF ETHANOL IN FEED
018
OZC
Figure 2. Energy required for producing 90 wt % ethanol as a function of feed composition. Basis: reflux ratio is (1.2) (minimum reflux); 99% ethanol recovery in distillate; feed a t saturation temperature.
cantation. The energy integration of the process is achieved by an oil loop which recovers the exothermic heat of the conversion reaction and supplies part of the energy required in the distillation column reboiler and in the preheating of the reactor feed. Energy Requirements and Energy Optimization Energy requirements for distilling ethanol from aqueous mixtures were estimated for various feed and overhead product compositions. Since the ethanol/water system is highly nonideal, the method of Ponchon-Savarit was used in estimating the distillation energy requirements. Equilibrium data for this system were obtained from Henley and Seader (1981). Initial computations showed that the degree of ethanol recovery has a minor effect on the energy requirements of the system. For this reason an ethanol recovery of 99% was assumed. Furthermore, in all cases, the feed was assumed to enter the column in saturated liquid state. The effect of ethanol content of the feed stream on the energy required to produce a 90 wt % ethanol product by distillation is shown in Figure 2. This energy requirement does not include the heat required to bring the feed mix-
ture from ambient temperature to the saturated liquid state. As expected, as the concentration of ethanol in the feed mixture increases, the energy required for distillation decreases. Since the slope of the curve begins to flatten out around a feed composition of 10% and since 10% alcohol feedstocks are the highest concnentration levels produced economically by conventional fermentation techniques, this feed composition was chosen for further analysis. The effect of overhead product composition on the energy requirements for distillation of a 10 w t % feed mixture is shown in Figure 3. As expected, the energy requirement increase with overhead product purity. Nevertheless, the incremental energy expenditure for producing overhead product of concentration greater than about 75% increases rapidly. The above computations provide general guidelines as to the regions in which optimum feed and overhead compositions may lie but do not provide the optimum values for these variables. The energy requirements of the overall integrated process was computed for distillation overhead composition in the range of 50 to 80 wt % ethanol. The energy requirements of the overall process include net energy for the distillation step, the reactor step, and electrical energy required to operate the compressors. In Figure 4, the net energy required for the integrated process is given as a function of the distillation overhead product composition. A feed stream of 10% ethanol was assumed. A sharp minimum of net energy requirement exists at an overhead product composition of 60% ethanol. This optimum value exists because the amount of energy required for the distillation (asthe overhead composition is lowered) decreases a t a rate lower than the rate of increase of the energy required to drive the compressors. The energy for the compressors increases due to the increased amount of water vapor in the various streams. The assumptions involved in these computations are shown on Table I. Component analysis, energy content, as well as temperature and pressure of each of the streams shown in Figure 1 are reported in Table 11, based on the optimum conditions described above. The power requirements of compressors C1 and C2 are 167.7 and 0.94 hp, respectively.
Ind. Eng. Chem. Process Des. Dev., Vol. 23, No. 4, 1984 735
perature. A fraction of the heat recovered from the reactor is used to preheat the reactor feed mixture while a significant portion of the remaining heat is supplied to the reboiler of the distillation column; the excess heat is rejected to the environment through the heat exchanger,
Table I. Assumptions for Material and Energy Balances 1. Excess enthalpy of mixing of gaseous ethanol and water is independent of temperature. 2. Compressor efficiencies are 90% with a 5% heat loss. 3. Compression follows the ideal gas equations for an isentropic compression. 4. The heat of mixing is negligible for hydrocarbon-water systems. 5.The equilibrium in the flash chamber follows Raoult's law. 6.Vessel and piping heat losses are negligible. 7. The complex product distribution from the reactor can be represented by their closest, simpliest analogue (i.e. all hexanes will be represented by normal hexane). 8. Temperature of approach in heat exchangers is 10 O F .
H-10.
Heat Integration The heart of the proposed process for ethanol recovery from fermentation beers is a heat integration loop which recovers the exothermic heat of the conversion reaction and supplies it to the distillation step. The conversion reactor is assumed to be of the heat exchanger type in which a heat transfer oil circulates absorbing the heat of reaction, thus maintaining the reactor at the desired tem-
The exothermic heat of the reaction is not the only heat recovery scheme in the proposed process. A number of other heat integration steps supply a significant amount of the heat required in the process: the distillation column feed stream is brought to saturation temperature by three stages of heat recovery, from the overhead condenser, the column bottoms, and the product from the separation chamber. A fraction of the heat requirement of the column reboiler is supplied by the condensing of the gaseous products from the reactor. The energy balance of the overall process shows a net heat input of 57 Btu/lb of ethanol. The net heat input is supplied by steam as the final preheating of the reactor feed mixture to the reactor temperature in H-7. An additional energy requirement of the process is the electrical energy for operating the compressors. The total electrical energy requirement is esti-
Table 11. Material a n d Energy Balance of t h e Integrated Process (Figure 1) component stream 1 stream 2 stream 3 stream 4 ethanol, lb/h 2000 2000 2000 2000 water, Ib/h 18000 18OOO 18000 18Ooo 0 0 0 0 C2H4 0 0 0 0 C6H12 0 0 0 0 C6H14 0 0 0 0 C7H16 0 0 0 0 C8Hl8 C~HS-CHS 0 0 0 0 0 0 0 0 CeHdCH3h
total,mol/h total, lb/h pressure, psia temperature, "F enthalpy, M M Btu/h HHV,M M Btu/h component ethanol, lb/h water, lb/h
C2H4 C6H14 C7H16
CBHl8 CeH5-CH3 CsHdCH3)2 total, mol/h total,lb/h pressure, psia temperature, "F enthalpy, M M Btu/h HHV,M M Btu/h component ethanol, lb/h water, lb/h CZH4 C5H12
CsH14 C7H16
CBH18 C6Hs-CH3 CeH,-(CH& total, mol/h total, lb/h pressure, psia temperature, O F enthalpy, M M Btu/h H H V , M M Btu/h
1042.6 2oooo 14.7 77 0.33 25.53 stream 8 5 4091 0 0 0 0 0
1042.6 20000 14.7 116.3 1.08 25.53 stream 9 20 16680 0 0 0 0 0 0 0
0 0
227.2 4096 14.7 210.0 4.61 0.06
926.3 16700 14.7 126.3 1.11 0.26 stream 14 0 2094 0 228 240 133 348 102 154 129.1 3300 105.0 165.5 0.28 24.38
1980 1320 0 0 0 0 0 0 0
stream 6 489 326 0 0 0 0 0 0
stream 7 20 16680 0 0 0 0 0 0
0
0
1042.6 1042.6 116.3 28.7 20000 20000 815 3300 14.7 14.7 14.7 14.7 188.6 197.6 192.0 192.0 2.49 2.66 3.38 0.59 25.53 25.53 25.27 6.24 stream 10 stream 11 stream 12 1980 0 0 2138 1320 2094 0 301 0 253 0 228 251 0 240 0 136 133 351 0 348 104 0 102 155 154 0 116.3 3300 117.6 611.1 4.04 25.27 stream 15 0 2094 0 228 240 133 348 102 154 129.1 3300 105.0 68.0 0.02
24.38
142.9 129.1 3689 3300 110.0 105.0 572.0 230.0 3.36 0.46 24.38 31.82 stream 16 1980 1320 0 0 0 0 0 0 0
116.3 3300 125.0 461.7 3.79 25.27
926.3 16700 14.7 210.0 2.51 0.26 stream 13 0 44 301 25 11
3 3 2 1 13.7 390 105.0 230.0 0.08 7.44 stream 17 0 44 301 25 11 3 3 2 1
13.7 390 117.6 242.9 0.08 7.44
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Ind. Eng. Chem. Process Des. Dev., Vol. 23, No. 4, 1984
i
050
OS5
060
065
I
010
M A S FRACTION OF ETHANOL
IN
OB
080
085
DISTILLATE
Figure 4. Energy required for the integrated process as a function of distillate composition. Basis: feed is 10 w t % ethanol at 25 “C; see Tables I and 11.
Table 111. Integrated Heat and Net Energy Requirements ~~
~
~~
source distillation overhead distillation bottoms reactor product condensation product cooling total heat recovered
heat recovered, Btu/lb of ethanol 381.4 710.2 662.4 1425.3 88.5 3267.8
% of total heat recovered
11.7 21.7 20.3 43.6 2.7
net energy requirements (Btu/lb of ethanol) steam: 57 .O electricity: 215.4 total net: 272.4 total process energy requirements: 3540.2 Btu/lb of ethanol (total energy recovered)/(total process energy requirement) = 92.3%
mated to be equivalent to 215.4 Btu/lb of ethanol. The relative amounts of heat recovered from the various sources described above as well as the net energy requirements of the process are shown in Table 111. The condensation of the gaseous products from the conversion reactor provides the largest fraction of the recovered heat. All the other steps supply a significant fraction of the process heat requirement with the exception of the heat recovered by initial cooling of the hydrocarbon products, which accounts for only 2.7% of the total heat recovered in the integrated process. Comparison with Competing Processes For the various processes evaluated by Eakin et al. (1981), the net process energy requirement (starting with 10% alcohol) and energy effficiency are summarized in Table IV. Energy efficiency is computed as energy efficiency = gross heating value of output products/[ (gross heating value of input material) + (net process energy input)] Conventional two-stage distillation is the most energy intensive process, requiring 27 400 Btu/gal of ethanol with
Table IV. Comparison with Competing Processesa net process energy requirement, energy Drocess Btu/eal of ethanol efficiencv. % solvent extraction 3600 95.9 distillation (to 60%) followed 1800 94.9-97.0* by reactions to gasoline (based on 1 gal of ethanol processed which produces 0.678 gal of gasoline) 90.4 C 0 2 extraction 8000-11 000 84.3 vapor recompression (to 95%) 15 800 followed by azeotropic distillation conventional distillation 19 200 81.5 (95% ) followed by dehydration 81.1 conventional distillation 21 000 (95%) followed by low temperature blending with gasoline conventional two-stage 27 400 75.5 distillation “Data obtained from Eakin et al. (1981) except for distillation/reaction scheme. Value depends on actual hydrocarbon product distribution.
*
an energy efficiency of 75.5%. The proposed integrated distillation/reaction to gasoline system appears to require the lowest heat input per gallon of ethanol processed, that being one-half of its nearest competitor. However, the loss in heating value due to the conversion reaction can lower its overall energy efficiency to slightly less than that of solvent extraction depending on the final product distribution. The other processes require significantly more energy per gallon of ethanol processed and consequently their efficiencies are lower. Conclusion The proposed integrated distillation/reaction to gasoline process for the recovery of ethanol from fermentor feedstocks is very promising as an energy efficient means of producing commercial fuel. The main advantage of this scheme is that it produces a traditional fuel from a renewable resource which can be used in present day equipment without modification. The process uses relatively simple unit operations which are easy to control and maintain. The vast majority of the process energy required (92.3%)is recovered internally by heat interchange. This process could present an attractive alternative to gasohol when ethanol supplies become plentiful. Registry No. Ethanol, 64-17-5.
Literature Cited Boeckler, 8. C. U.S. Patent 140925, 1948. Chambers, R. S.;Heredeem, R. A.; Joyce, J. J.; Penner, P. S. Science 1979, 206, 789. de Boks, P. A.; Huber, T. F.; Van Nes, W. J.; van den Oosterkamp, P. F.; van Bekkum, H.; van Eybergen, 0. C.; van de Hende, J. H.;Kossen, N. W. F.; Moorlng, C. I.; Oudejans. J. C.; Snaterse, A. C.; Wessellngh, J. A. Biolechnol. Lett. 1982, 4(7), 447-452. De Flllppl, R. P. “Extraction of Organics from Aqueous Solutions wlth CriticaCFluM Carbon Dioxide", Paper presented at IVth Symposium on Biotechnology in Energy Production and Conservation at Gatlinburg, TN, May 11-14, 1982. Dreyfus, H. U S . Patent 2 053770, 1936. Eakln, D. E.; Donovan, J. M.; Cysewskl, G. R.; Petty, S. E.; Maxham, J. V. Preliminary Evaluation of Alternative Ethanol Water Separation Processes”; Paclflc Notthwest Laboratory. Richland, WA, Battele Memorial Institute Report PNL-3823, May 1981. Fanta, A. F.; Burr, R . C.; Orton, D. L.; Doane, W. M. Science 1980, 270. 646.
Ind. Eng. Chem. Process Des. Dev. 1904, 23 I 737-74 1
Henley, E. J.; Sea&, J. D. "Equilibrium Stage Separations in Chemical Engineering"; Wiley: New York, 1981. Ladisch, M. R.; Dyck, K. Science 1979, 205, 898. Leeper, S. A.; Wankat, P. C. Ind. €ng. Chem. Process D e s . D e v . 1982,
737
Scheller, W. A. "Energy Requirements for Grain Alcohol Production"; Presented at 176th National Meeting of the American Chemical Society, Miami Beach, FL, Sept 10-15, 1978. Scheibel, E. G. Ind. Eng. Chem. 1950, 42, 1497-1508. Whitcraft, D. R.; Verykios, X. E.; Mutharasan, R. Ind. Eng. Chem. Process D e s . Dev. 1983, 22, 452-457.
27, 331.
Oudejans, J. C.; Van Den Oosterkamp, P. F.; Van Bekkum, H. Appi. Catai. l.B - -8-,2 . -3 ,. 109-11s. .- - . .-. Poclus, D.; Ladlsch. M. R; Tsao. A. T. "Calcium Sulfate as a Selective Adsorbent of Water"; Paper 87d. AIChE Annual Meeting, LA, Nov 14-19, 1982.
Received for review March 28, 1983 Accepted November 14, 1983
Mixing Characteristics of the Positive and Negative Pressure Air Mixers Tetsuo Aklyama" and Isao Tada Department of Chemical Engineering, Shizuoka Universw, Hamamatsu, 432, Japan
This paper compares the mixing efficiency of the negative pressure air mixer (NPAM) and the positive pressure air mixer (PPAM). Experiments were carried out using five pairs of particles which were chosen from five kinds of particles. The quantity of air and work necessary to mix 1 kg of solid particles to a state of M (the degree of mixing) = 0.1 were compared. Experimental results indicated that the NPAM can become more advantageous than the PPAM in the range of practical use. This fact was then backed by a simple theoretical analysis through calculation of the work that can be done by the jetting air stream.
Introduction Mixing of particulate material is widely practiced in industry. It is often required to mix two or more kinds of particulate materials or to homogenize raw or manufactured material. Attainment of fine homogeneity is highly desired in ceramics and other industries and the rising energy cost calls for more energy-effective mixing techniques. When large quantities of solids are to be handled, an air mixer is often used. A drawback of this mixer is that it requires filter bags. To overcome this drawback, a negative pressure air mixer (NPAM) was proposed in the previous paper (Akiyama et al., 1982). The NPAM is suitable to handle smaller volumes of various kinds of particulate material which are sensitive to contaminants. Thus the NPAM was recommended for mixing of foodstuffs or medicine. A quantitative comparison of the mixing efficiency among the NPAM, the positive pressure air mixer (PPAM), and the V-shaped tumbling mixer has been reported by Akiyama et al. (1983). The study indicated that under limited conditions, the NPAM was more advantageous than the PPAM and that the V-shaped tumbling mixer was better suited for rather coarse mixing but it was unable to attain as fine a homogeneity as was possible through the use of the air mixers. The mixing characteristics of the nonuniform fluidization mixers were investigated by Ando et al. (1970) and Sano (1973), and comparisons among various types of mixers were reported by Miles et al. (1970) and Shah (1973). These studies, however, were concerned with a steady-state type of operation, and direct comparison with the NPAM was not possible. The operation of the NPAM is inherently transient because it needs repetitive air injections. The objective of the present investigation is to make a quantitative comparison of the mixing efficiency between the NPAM and PPAM over a broader range of conditions than those studied in the previous work. As criteria of comparison, the air quantity and work that are necessary 0196-4305/84/ 1123-0737$01.50/0
Table I. Physical P r o m r t i e s of Particles true bulk av diam density, density, Particle X lo3. m ke/m3 ke/m3 glass bead A 0.214 2520 1600 glass bead B 1.705 2520 1600 millet seed 1.573 841 1386 polythylene 3.70 940 563 pellet polystyrene 0.997 1070 651 pellet
void ratio 0.365 0.365 0.393 0.401
angle of repose, rad 0.262 0.209 0.611 0.436
0.392
0.436
Table 11. Physical Properties of Systems
system G G : glass bead A-glass bead B M-G millet seed-glass bead A P-G: polyethylene-glass bead A PS-G: polystyrene-glass bead A P-M: polyethylene-millet seed
diam ratio 7.97 7.35 17.3 4.66 2.35
angle of repose, rad 0.349 0.576 0.384 0.462 0.733
to secure the desired degree of mixing were experimentally determined. Then it was shown that the NPAM can be more advantageous than the PPAM. A simple theoretical analysis was also made to show that the NPAM can indeed become more advantageous than the PPAM in the range of practical interest. Experimental Section Schematics of the NPAM and PPAM are shown in m3, Figures 1and 2, respectively. A column (8.16 X 0.3 m i.d.) made of transparent cylindrical acrylic resin with a cone of 65' was used as a mixing vessel. This column was interchangable with the rest of the test equipment. We used five kinds of particles, from which five pairs weighing 7.5 kg each (15 kg in total)were chosen to carry out experiments. Physical properties of these particles are listed in Table I, and some pertinent properties of the pairs are shown in Table 11. The initial 0
1984
American Chemical Society
