I n d . Eng. Chem. Res. 1987,26, 699-705
rates and compositions of all the streams of the process are summarized in Table V. The calculated number of equiIibrium stages and the feed plate location are shown in Figure 3 on each column. The greatest energetic requirements of the process are the latent heat for the vaporization of stream 6 in dehydration column Q (about 5.36 X 40 = 224 kJ/mol of ethanol obtained; the latent heats for water and ethanol are about 40 kJ/mol) and in solvent recovery distillation column R (about 2.2 X 40 = 88 kJ/mol) and a sensible heat of stream 4 (about 0.26 X 2.45 X (184 - 25) = 101 kJ/mol; the heat capacity and boiling temperature of 2-ethylhexanol are 0.26 kJ/(mol "C) and 184 OC,respectively). In addition a significant fraction of the required latent and sensible heat input can be recycled through heat exchangers which are not shown. These energetic requirements represent an important fraction of the energy that would be liberated if the alcohol is burnt as a fuel (1360 kJ/mol). However, it can be possible to find another solvent with properties more favorable than 2-ethylhexanol. These properties should be those decreasing the flow rates of streams 4, 6, and 8. (1) A greater distribution coefficient for ethanol is needed to decrease the solvent-to-water ratio and so decrease stream 4. (2) A greater separation factor is needed to decrease the reflux ratio in dehydration column Q and so decrease stream 6. (3) The enhancement by the solvent of the relative volatility of water to ethanol should be greater in order to decrease both the reflux ratio in the dehydration column (and so decrease stream 6) and stream 5, since less flow rate of 2-ethylhexanol is necessary to keep the water more volatile than ethanol. (4) The solvent must have less heat capacity and a lower boiling point to decrease the enthalpy of stream 11 (and therefore the enthalpy of stream 4). As proposed by Munson and King (19841, a diluent can be incorporated
699
into the solvent so as to provide volatility in the reboiler of solvent recovery distillation column R. However, the solvent volatility must be sufficiently less than that of ethanol to facilitate the separation in this column. (5) The solubility of the solvent in water must be low enough so that it is not necessary to remove and recover the residual solvent from the aqueous raffinate (stream 3) leaving the extractor. Registry No. H 3 C C H 2 0 H , 64-17-5; H3C(CH2)3CH(CHZCHJCHZOH, 104-76-7.
Literature Cited Abrams, D. S.; Prausnitz, J. M. AIChE J. 1975,21, 116. Brown, W. V. Chem. Eng. Prog. 1963, 59(10), 65. Fredenslund, A.; Gmehling, J.; Rasmussen, P. Vapor-Liquid Equilibria Using UNZFAC; Elsevier: Amsterdam, 1977. Gmehling, J.; Rasmussen, P.; Fredenslund, A. Znd. Eng. Chem. Process Des. Dev. 1982, 21, 127. Hunter, T. J.; Nash, A. W. J . SOC.Chem. Znd. 1932, 51, 285T. Ladisch, M. R.; Dyck K. Science (Washington,D. C.) 1979,205,898. Munson, C. L.; King, C. J. Ind. Eng. Chem. Process Des. Dev. 1984, 23, 109.
Nelder, J. A.; Mead, R. Comput. J . 1965, 7, 308. Othmer, D. F. Ind. Eng. Chem. 1958, 50(3),60A. Prausnitz, J.; Anderson, T.; Grens, E.; Eckert, C.; Hsieh, R.; O'Connell, J. Computer Calculations for Multicomponent Vapor-Liquid and Liquid-Liquid Equilibria; Prentice Hall: Englewood Cliffs, NJ, 1980. Roddy, J. W. Ind. Eng. Chem. Process Des. Deu. 1981, 20, 104. Roddy, J. W.; Coleman, C. F. Ind. Eng. Chem. Fundam. 1981,20, 250.
Ruiz, F.; Gomis, V. Ind. Eng. Chem. Process Des. Dev. 1986,25(1), 216. Ruiz, F.; Prats, D. Fluid Phase Equilib. 1983, 10, 77. Ruiz, F.; Prats, D.; Gomis, V. J . Chem. Eng. Data 1984, 29, 147. Stabnikov, V. N.; Matyushev, B. Z.; Protsyuk, T. B.; Yushchenko, N. M. Pishch. Prom. 1972, 15, 49. Received f o r review October 25, 1985 Revised manuscript received September 22, 1986 Accepted December 6, 1986
Progressing Batch Hydrolysis Reactor J o h n D. Wright,* Paul W. Bergeron, and Pamela J. Werdene Solar Energy Research Institute, Golden, Colorado 80401-3393
In the dilute acid hydrolysis of lignocellulose to produce fermentable sugars, conditions severe enough to hydrolyze crystalline cellulose are also severe enough to degrade the product sugars. Countercurrent flow of liquids and solids minimizes sugar degradation and product dilution by removing the sugars from the reaction zone before substantial degradation can occur. The progressing batch reactor uses several percolation reactors in series to simulate countercurrent flows while retaining the simplicity of the percolation reactor. In all dilute acid hydrolysis processes for glucose production, conditions severe enough to hydrolyze crystalline cellulose to glucose are also severe enough to degrade the glucose into undesirable compounds such as hydroxymethylfurfural (HMF), levulinic acid, and formic acid. One way to minimize the sugar degradation is to remove the sugars from the reaction zone before substantial degradation occurs. Sugars are most efficiently removed by a reactor system that uses countercurrent flow of liquids and solids, which allows simultaneous achievement of high yields and high sugar concentrations. The progressing batch hydrolysis process, invented and now under development at the Solar Energy Research Institute (SERI), 0888-5885/87 / 2626-0699$01.50/ 0
uses several percolation reactors in series to simulate countercurrent flow of liquids and solids. In this way, the advantages of countercurrent flow are achieved, and the mechanical and operational simplicity of the percolation reactor is retained. This paper describes the theory and operation of the progressing batch hydrolysis reactor and presents the results of our mathematical modeling of the system. Theory of Operation Hydrolysis Kinetics. The high-temperature acid hydrolysis of the carbohydrate portion of lignocellulose can 0 1987 American Chemical Society
700 Ind. Eng. Chem. Res., Vol. 26, No. 4, 1987
be approximated by the simplified reactions crystalline cellulose
- - -
amorphous glucans
-
xylans
k,
k3
glucose
glucose
- k3
xylose
k4
k2
HMF tars and organic acids
- k2
HMF tars and organic acids
furfural
-
tars
The hydrolysis of crystalline cellulose is difficult because of its highly ordered structure, but the hydrolysis of the amorphous five- and six-carbon sugars (xylans and glucans) proceeds at much faster rates (Harris, 1975). Studies of the kinetics of acid hydrolysis and sugar decomposition have been carried out by many investigators, notably Saeman at the Forest Products Laboratory and Grethlein at Dartmouth College. Saeman (1945) published the first paper describing the pseudo-first-order kinetics of cellulose hydrolysis and glucose degradation. Detailed investigations of the degradation of the five- and six-carbon sugars were carried out by Root et al. (1959) and McKibbins et al. (1962). Grethlein's group extended these studies to higher temperatures and shorter reaction times (Kwarteng, 1983). The primary problem of all dilute sulfuric acid hydrolysis processes is that conditions severe enough to hydrolyze crystalline cellulose are also severe enough to degrade the product sugars. Historically, this problem has been dealt with by two different methods (high-temperature hydrolysis and percolation hydrolysis). A t high temperatures (approximately 240 "C) and 1%H2S04, the hydrolysis reactions are faster than the degradation reactions, and yields of approximately 55% are achieved at residence times of 10 s or less. This approach is used in the plug-flow reactor, which has been studied at Dartmouth (Grethlein and Converse, 1982) NYU (Rugg et al., 1981), and SERI (Wright, 1983). This method produces a moderate yield but can potentially achieve a concentrated sugar solution if a reactor can be built that processes a high solids feed at low residence times. Percolation Reactors. A second approach is the percolation reactor, which has different residence times for the solid and liquid fractions. Conditions in percolation reactors are generally 180 "C and 0.5% H2S04(Harris and Beglinger, 1945). Since the rates of hydrolysis and sugar degradation are similar at these conditions, the yields in a batch or plug-flow reactor would be less than 30%. Although the solids (wood chips) are kept in the percolation reactor for approximately 3 h so that the cellulose has time to hydrolyze completely, the water and acid flow continuously through the reactor with a residence time of approximately 45 min. Thus, the sugars dissolve in the water, are washed out of the reactor, and are cooled before substantial degradation can occur (Cahela et al., 1983; Maloney et al., 1984). The advantages of the percolation reactor are its relatively high yield (6C-70%) and its simplicity (all solids are processed at atmospheric temperature and pressure). The disadvantage is that in order to achieve the high yields, large amounts of water are required to quickly wash the sugar from the reactor, resulting in dilute sugar solutions (Wright and d'Agincourt, 1984). Percolation reactor technology was f i t developed in the late 1920s in Germany. The Scholler-Tornesch process operated at 170 "C with 0.4% acid. In this process the liquid was intermittently fed into the reactors and then drained out. Approximately 12-15 batches of liquid were used over a 15-18-h operating cycle, producing a 40% yield of glucose from cellulose (Wenzl, 1970). A commercial-
scale plant was built with three 25-50-m3 hydrolysis reactors. In the 19409, an improved process was developed at the US Forest Products Laboratory in Madison, WI (Harris and Beglinger, 1945), and a large commercial-scaleplant was designed that was capable of processing 220-300 tons of wood/day. One of the reactors was tested for 6 months, but World War I1 and the fuel shortage ended before the full plant was operating. In the Madison process, acid flowed continuously through the reactor, reducing the residence time to 3 'Iz h. The sugar concentration before the flash valve was increased to 4% and the yield to 50%. The reaction was initially carried out at 150 "C and 0.5% H2S04to recover the hemicellulose sugars at high yields, after which the temperature was raised to 180 "C to hydrolyze the crystalline cellulose. In the early 19509, researchers at the Tennessee Valley Authority resumed work on the Madison process (Gilbert et al., 1952). The yield was increased to 65% and the total sugar product concentrations to 4% before the flash valve. In recent years, further development has been carried out by Inventa A.G. in Switzerland (Mendelsohn and Wettstein, 1981) and by the New Zealand Forest Research Institute (Burton, 1982; Wayman, 1983; Mackie et al., 1982). However by far the largest application of such reactors has been in the Soviet Union, where over 20 commercial plants are in operation (Wenzl, 1970). The percolation reactor is a proven technology. Although the designs, details, and procedures necessary for building a reliable reactor are not all available in the literature, the existence of the Russian percolation industry suggests that reasonable designs exist. The principal operational problems reported in the literature include fouling of the outlet pipes by degradation products and collapse, compaction, and plugging of the chips within the reactor as hydrolysis progresses (Korolkov et al., 1961). Countercurrent Reactors. An alternative flow arrangement, the countercurrent reactor, offers considerable advantages over the batch, plug-flow, or percolation configurations. Because the hydrolysis reactions are essentially first order, the rate of sugar production is greatest when the solids are first introduced into the reactor. With countercurrent flow, the liquids outlet is adjancent to the solids inlet. Thus, the largest fraction of sugars has only a small distance to travel before being washed out of the reactor, which minimizes the time available for sugar degradation and maximizes the yield and sugar concentration. For comparison, in a percolation reactor the sugars are created evenly throughout the reactor. Therefore, on the average the sugars flow halfway through the reactor before they are removed. In a typical countercurrent design, sugars travel only about one-fifth of the reactor length on the average. Thus, the countercurrent reactor with the same nominal residence times for solids and liquids as a percolation reactor suffers less than half the sugar degradation of the percolation reactor. The yield is thereby increased from approximately 65% to about 80%, and the sugar concentration increases by 20%. The advantages of the countercurrent reactor have been described by several groups. Investigators at A. D. Little (Greenwald et al., 1983) and Elf Aquitaine (Ladousse et al., 1981) used simple mathematical models to compare several different reactor configurations. A more detailed model of the three major reactor types was prepared by researchers at Auburn University (Song and Lee, 1982; Song, 1983). Although the continuous countercurrent reactor is attractive on paper, it may be difficult to construct and
Ind. Eng. Chem. Res., Vol. 26, No. 4,1987 701 7
k
C-0A Solid
Solids
Degradation products
Sugar
U
k-
-
x x + A x
Liquids
(>v
x.0
x = L
c, = c,
c, = 0
Figure 2. Material flows and boundary conditions for an ideal countercurrent reactor.
P = p r e h y d r o l y s i s H = h y d r o l y s i s E=emptyinQ F=fillinQ S=soaking
Figure 1. Progressing batch reactor operation.
operate. The movement of a concentrated stream of biomass and water into and out of a pressurized reactor is a very difficult operation, even when the liquid and solids are moving in the same direction (as in a plug-flow reador). The equipment required for continuous countercurrent reactor would be considerably more complex, and no practical designs have been proposed to date. Progressing Batch Reactors. The progressing batch hydrolysis reactor, conceptualized at SERI, operates on the principle that a continuous process can be approximated by several batch processes in series. Figure 1shows several percolation reactors in series. In Figure l a the liquid (acid and water) enters reactor 7 (a hydrolysis reactor) and flows from right to left through reactors 7, 6, 5 , 4 and 3. After passing through reactor 3, the hydrolyzate liquids exit the reactor train and are flashed to near atmospheric pressure to quench the hydrolysis and degradation reactions. During this time, reactor 2 is being filled with chips and prepared for hydrolysis, while reactor 1 is being emptied of the unreacted residue from a previous cycle. After 15-30 min, the flow is changed to the configuration in Figure l b by simple changes in the valve position. Fresh liquids enter reactor 6 and flow through reactors 6, 5 , 4 , 3, and 2. After exiting reactor 2, the liquids leave the train. Reactor 1 is now being filled with fresh wood, and the residual solids in reactor 7 are being dumped. Figure IC shows the next stage in the cycle. The net effect of this process is that fresh solids have entered the reactor train at the left, and spent solids are dumped after leaving the last hydrolysis reactor in the train on the right. The solids appear to move from left to right through the active reactors, while the liquids move from right to left. Thus, the apparent countercurrent motion of the solids and liquids is achieved without the necessity for moving solids into and out of a pressure zone, and all solids handling is done at ambient temperature and pressure. The advantages of countercurrent operation are realized while retaining the operational simplicity of the percolation reactor.
ance and in identifying potential problems and process improvements. In each, the progressing batch reactor was modeled as a continuous countercurrent reactor. This assumption simplifies the modeling considerably and in many cases allows the use of analytical solutions instead of numerical simulations. The validity of this assumption depends on the number of active stages in the system; the assumption becomes more valid as the number of stages increases. The following sections describe the modeling work and show how the models were used to improve the processing system. Ideal Countercurrent Reactor: Effects of Residence Time, Temperature, and Packing Density. In the simplest reactor model, the wood and water flow in opposite directions in perfect plug flow. There is no sugar in the incoming liquid stream. The wood is assumed to be made of infinitely small, nonporous particles, so mass-transfer effects are neglected, and the absorption of liquid by the wood is ignored. The temperature and acid concentrations are assumed to be constant throughout the reactor, which is operating at steady state. For this case, an analytical solution can be developed. Figure 2 shows the physical system and the boundary conditions. A, B, and C represent the solid carbohydrates, sugars, and degradation products, respectively, since the derivation applies equally to the reactions of crystalline cellulose, amorphous glucans, and xylans. Writing a material balance over a differential element of the reactor, we get
and
x=L
cg=o
(2)
where CA and CB are the concentrations of the solid carbohydrate and soluble sugar, kl and k2 are the rate constants for decomposition of A and B, and U and V are the solid and liquid velocities. The solution is
Mathematical Modeling The principles and advantages of the progressing batch reactor can be easily understood from the qualitative arguments presented so far. However, to accurately evaluate the potential of the concept and to confidently design an experimental apparatus, it is necessary to mathematically predict the performance of the reactor system. Therefore, several models of varying levels of sophistication have been developed that are useful in estimating process perform-
and CB t s y = -cAO
(4)
tl
where tl and t , are the liquid and solid residence times.
702 Ind. Eng. Chem. Res., Vol. 26, No. 4, 1987 1.0,
I
09
08 01
07
2
02
06
0,
x
0 53
, 60 3o
03
I
1
1
70 80 90 Yield 01 glucose f r o m crystalline cellulose
0.2 100
1
01
Figure 3. Comparison of percolation and countercurrent reactors. Kinetic parameters for cellulose hydrolysis by Saeman (1945); glucose degradation parameters by McKibbins et al. (1962).
A similar derivation (Song and Lee, 1982) gives eq 5 and 6, which describe the performance of a percolation reactor:
(5) and CBt, y =CAOtl
Using these two equations, we can make some observations about the major variables and about the relative performance of the two systems. The effect of the major variables (solid and liquid residence times and reaction rates) is the same for both reactor types. The most important variable is the liquid residence time. Decreasing the liquid residence time increases the yield by removing the sugars more quickly from the reaction zone. To a first approximation, cutting the liquid residence time in half (doublingthe flow rate) cuts the degradation losses in half. Thus, the yield is sensitive to changes in the flow rate when the flow rates and yields are low but is relatively insensitive at high flows and yields. As solids residence time increases, the yield also increases, because larger fractions of the feed are converted to sugar. Since most of the sugars are produced early in the reaction, the yield is most sensitive to changes in solids residence time when reaction times and yields are low. Two other parameters of importance are the reaction rates and the solids loading. Increasing the temperatures increases both kl and k,, but kl increases more quickly. Therefore, increasing the reaction temperature can increase both the yield and product concentration. Also, the outlet sugar concentration is directly proportional to the inlet concentration, so the reactor vessels should be packed as densely as possible with the feedstock. Unfortunately, those factors that work to increase yields also decrease the product concentration. Increasing the flow rate dilutes the sugars, while increasing the solids residence time prolongs the time in which the sugars are produced. Thus, the proper choice of reaction conditions involves a tradeoff of yield and product concentration. Figure 3 shows the relationship between the sugar concentration and yield produced from crystalline cellulose, with liquid residence time as a parameter. In both cases, the yield increases and the product concentration decreases with decreasing liquid residence time. However, for a given liquid residence time, the countercurrent reactor has half the sugar degradation and a 20% higher glucose concentration than the percolation reactor.
00
'
1
I
2
"
3
Sugar
'
4
'
5
I
'
6
concentration
7
1
8
9 1 0
(wt %)
Figure 4. Countercurrent reactor yield and concentration as a function of porosity (e) and w (residence time ratio).
Porosity and Backmixing. The idealized model of the countercurrent reactor describes the effect of the major variable but leaves out several important effects of the porosity of wood. Depending on the type of wood, up to 80% of the apparent volume of a dry wood chip may be void space. When a chip is placed in a pressurized reactor, it will soak up water. Unfortunately, the water absorbed contains the highest concentration of sugar since the solid inlet is at the liquid outlet. This solution is then carried into the reactor with the solids, thereby increasing the time the sugar is in the reactor, increasing the degradation, and decreasing the yield. A model accounting for these effects was developed at Auburn University (Song and Lee, 1982). The important additional assumptionsare as follows: 1. The temperature, acid concentration, porosity of the chips, and void fraction in the bed are uniform throughout the reactor. 2. The sugar concentrations inside the chips and in the free liquid phase are in equilibrium. Again, when we write a material balance around the system shown in Figure 2, the equation describing the solids is unchanged. Writing a balance on the sugars we obtain
[U€- V(1kz(6
+ (1 - C)o/CB = 0
(7)
where 0 is the porosity of the chip and t is the void fraction of the bed. The first term accounts for the movement of both the free liquid and the liquid trapped within the pores of the wood particle. The second term is the generation of sugars by hydrolysis, and the third term accounts for the degradation of the sugars in both the pores and the free stream. Solving these equations for the concentration of the sugars at the hydrolyzate outlet gives
and
where w is the ratio of the solid to liquid residence times. Figure 4 shows the results predicted by eq 8 and 9. Note that for any given ratio of residence times (constant a), the performance of the reactor improves as the porosity of the wood decreases because less sugar is being carried
Ind. Eng. Chem. Res., Vol. 26, No. 4, 1987 703 120
140
Temperature ( " C ) 160 180 200
220
Acid concentration = 0 7 wt -010
000
020
040
060
080
100
Characteristic length i c m j
Figure 5. Comparison of time constants for diffusion and reaction as a function of length and temperature. Characteristic length for diffusion of H8O1and glucose is one-half the length along the grain. Length for heat is the shortest particle dimension.
back into the reactor. However, in comparing these results with those from the more simplified model, we find a serious discrepancy. For any given yield, this model predicts outlet sugar concentrations 3 times higher than those predicted by the simpler model because of a flaw in assumption 2. As the hydrolysis reaction progresses, up to 80% of the volume of the wood is converted to soluble products. The wood particles lose their rigidity and collapse, and the void fraction of the bed may increase from approximately 35% to 85%. If we ignore this variation, the model grossly underestimates the amount of water needed to achieve a given liquid residence time and therefore grossly overestimates the sugar concentration. If we understand the limitations of this model, we may still extract useful information from it. The model warns us of the undesirable effect of sugars being trapped in the porous wood particles and being carried back into the reactor. Transport of Acid, Sugar, and Heat. Another potentially flawed assumption of the model is that the liquid inside and outside the particle is in equilibrium, which is undoubtedly true for very small particles. However, the size reduction of wood is an extremely energy-intensive process, and it is to our advantage to use the largest possible particle size that does not result in significant reductions in yield and product concentration. While it is difficult to accurately estimate mass- and heat-transfer rates within a wood particle, some simple approximations allow us to understand the important parameters (Sherwood et al., 1985). Figure 5 shows characteristic times for (1) the transport of heat, acid, and glucose as a function of wood particle size; (2) the hydrolysis of crystalline cellulose and hemicellulose; and (3) the degradation of glucose and xylose as a function of temperature. In all cases, the transport problem is modeled as diffusion into or out of a semiinfinite slab having a thickness of 2y0. initially, the concentration everywhere is Co. At to,the concentration outside or inside the slab is changed to Ciand held constant. The fractional approach to the new steady state is
and
An equation with identical form may be written for the case of heat conduction. Since all of these processes are exponential, the characteristic time is that required for the process to go 63.2% of the way to completion. The characteristic length for the diffusion of glucose and H2S04 is one-half the particle length measured along the grain, and the diffusion coefficient is the diffusion coefficient in water multiplied by the square of the porosity of the particle. For the diffusion of heat, the characteristic length is the shortest particle dimension. The particle was modeled as a solid with the thermal conductivity of wood, while the heat capacity was a weighted average of the wood and the water filling the pores. The reactions are all first-order processes as well. The characteristic times are equal to l / k e and are functions of temperature and acid concentration. Figure 5 shows that for small particles the time required for diffusion is small compared to the time required for reaction, and the effects of mass transfer can be neglected. As the particles become larger, the mass-transfer effects may become significant. For example, the wood particles will always have a characteristic length for heat transfer of 0.3 cm or less, and heat-up times will always be 1min or less. These heat-up times are short compared to reaction times (3 h or more), so they can be safely ignored. For large chips, the time required for acid to diffuse into the particle may be 20-40 min. This is a significant fraction of the total reaction time, so it may be useful to force the acid into the chip instead of relying on diffusion. For all but the smallest wood particles, the time required for diffusion of sugar out of the wood is appreciable compared to the rate of xylose degradation at hydrolysis conditions (180 "C). The relative rates of reaction and diffusion will affect the yield. Once the sugar has been produced by hydrolysis, it is subject to degradation by the acid. The amount of degradation is proportional to the residence time of the sugar in the reactor. The total time for sugar degradation is the time spent diffusing out of the chip plus the time required for the sugar in the free liquid to be washed out of the reactor. Therefore, a prehydrolysis would be useful to recover the xylose and amorphous six-carbon sugars at temperatures where the degradation rate is lower. For larger particles (2 cm or 1 in. long), the times for sugar diffusion are similar to the times for degradation of glucose. Therefore, we would expect a tradeoff between the cost of feedstock size reduction and glucose yields. Several assumptions were included in the calculation of the characteristic times for diffusion in Figure 5. For example, the actual diffusivity of sugar in a wood particle has not yet been measured. Also, there may be an appreciable amount of diffusion across the grain in a wood chip, since the cross-grain dimension is typically much less than the length along the grain. Finally, the particles may shrink dramatically as the reaction progresses. With a hardwood feedstock, the chips may disintegrate into a fine mudlike material as the reaction progresses. Therefore, these calculations should be used only as a qualitative tool to anticipate some of the important effects in an actual reactor. Reactor Design. While the modeling effector is severely limited by the necessity for simplifying assumptions and by the uncertainty of the parameters, we have learned several important things: 1. Outlet product concentration is directly proportional to the packing density of the wood
704
Ind. Eng. Chem. Res., Vol. 26, No. 4, 1987
bed. 2. Increasing the liquid residence time improves product concentration but reduces yield. 3. Increasing the hydrolysis temperature increases potential yield and sugar concentration. 4. The absorption of liquid by the dry chips at the reactor entrance can severely reduce the reactor yield. 5 . As the particle size or temperature is increased, degradation of sugars within the chip becomes more important. 6. The time required for acid to diffuse into a large chip may be significant compared to the overall cycle time. These observations allow us to improve our reactor design and performance. The first improvement is to realize that we must use a feedstock with a wide distribution of sizes. With a uniform size feedstock, it is difficult to obtain a bed density greater than 8-10 dry lb/ft3 (depending on the density of the wood species). With a mixture of sizes, the small particles fill the voids between the larger chips, making packing densities of 14-18 lb/ft3 possible. The use of a smaller fraction will also reduce the problems arising from diffusion within the particles. A second improvement is to presoak the chips in a dilute acid-water solution before raising the temperature to begin the reaction. By presoaking the chips, the pores are filled with water, preventing the chip from absorbing the hydrolyzate as it enters the reactor. Sugar will still diffuse into the chip but a t a much slower rate and only until sufficient sugar has been produced within the chip to reverse the direction of diffusion. A second benefit is that the presoak rapidly impregnates the wood particles with acid. In the presoak, liquid is formed into the chip by pressure. Thus, the acid is rapidly delivered to the chip's interior, and the delay of diffusion is eliminated. A third improvement is to incorporate a prehydrolysis step prior to the main cellulose hydrolysis. If the hemicellulose is carried out a t 150 "C instead of 180-185 "C, the hydrolysis rates for the amorphous five- and six-carbon sugars are rapid, but the sugar degradation rates are greatly reduced. Thus, at the beginning of the reaction, where the time required for the sugars to diffuse out of the chips is greatest, the rate of sugar degradation is significantly reduced. Also, the prehydrolysis opens up the interior of the chip, allowing the sugars produced in the higher temperature hydrolysis to diffuse out more quickly. There are tradeoffs in the design of the reactor system; the most notable are between the sugar concentration and yield and between the temperature and particle size. The first tradeoff involves the classic dilemma of the percolation reactor. Simple process optimizations conducted at SERI indicate that a rather broad optimum exists where concentration and yield may be varied with little effect on the overall process economics. This optimum occurs in the region of 183 "C, 0.7% H2S04,a solids residence time of 2.5 h, and a liquid residence time of 70 min. Under these conditions, the simplified models of the countercurrent system predict a glucose yield of 60% with a concentration of 6% and xylose yields in excess of 90% with a concentration of 5 % . The simple kinetic models predict that higher yields and product concentrations may be achieved at higher temperatures. However, this prediction does not account for diffusion time. In the simple theory, as the temperature is raised, the rates of both the hydrolysis and degradation reactions are increased, but the ratio of the two rates becomes more favorable. The increased flow rates needed to remove the sugars quickly from the reactor are offset by the reduced reaction times, and a higher yield and concentration are obtained. However, for any given size of particle, at a certain temperature the time for diffusion
l
h
-
0 7'0 H SO H y d r o y s i s 183-C Frehydrolysls 150 C
08
3 >
1 "
0 so
20
40
60
80
100
Solids r e s i a e n c e I me imin I
L quia E x t i
120
140
L qLid E r l r a r c e Solids E x i t
Figure 6. Development of yields along reactor length for countercurrent hydrolysis of mixed hardwoods.
7-
Irk
P *ale "dULt 5
98
dC
d
Figure 7. Progressing batch reactor schematic
out of the particle is dominant. When this occurs, increases in flow rate will no longer increase the yield. While models have been developed to predict the combined effects of temperature, flow rate, and diffusion (Song, 1983), their simplifying assumptions limit their usefulness for quantifying the tradeoff. Perhaps the best guide to appropriate design tradeoffs is the literature on percolation reactors. Of the five major research groups that have worked with a variety of woods and particle sizes, none report using temepratures over 190 "C. Many models of varying complexity can be used to identify the limiting factors in countercurrent hydrolysis, but we feel that the most useful is one that takes into account the use of both a hydrolysis and prehydrolysis stage but neglects the effects of backmixing and diffusion. Modeling the reactor as a series of stirred tanks allows us to explore the effects of varying the temperature and acid concentration along the reactor length and to see the development of the sugar yields and concentrations along the reactor. Figure 6 shows the development of the yields of the major reaction species at our nominal design conditions. System Design. Using the design information developed above, we can now prepare a preliminary hydrolysis system design (Figure 7 ) and set conditions that allow all solids processing to be done at ambient temperature and atmospheric pressure. When the first reaction vessel is being filled, it is isolated from the adjacent reactors by closing the valves in the inlet and outlet lines. The top hatch is opened, and the chips are dumped or blown into the reactor. The reactor is then sealed, purged of air, and preheated by passing steam through the tank. Removing the air from the system also makes it easier to force water into the interior of the chips and removes oxygen that might react with the product sugars. Reactors 2 and 3 are in the prehydrolysis stage, at 150 "C, while hydrolysis is taking place in reactors 4, 5, and 6. Reactor 7 is being emptied and made ready for a new charge.
Ind. Eng. Chem. Res., Vol. 26, No. 4, 1987 705
Conclusions Mathematical models are useful for showing the important parameters to be considered in the design of a progressing batch hydrolysis system. The important results of our model are that the product concentration increases with bed packing density and liquid residence time, but the yield is increased at reduced liquid residence times. Increasing the reaction temperature improves the kinetics, but the improvements are limited by the rate of diffusion of the product sugars out of the particles. Absorption of hydrolyzate by the fresh solids could significantly reduce the yields, but this problem can be largely eliminated by presoaking the chips. As a first approximation, the progressing batch reactor will have a 20% increase in both the yield and sugar concentration at the same conditions as a percolation reactor. Alternately, the sugar concentration can be doubled if the yield is the same as the percolation reactor. More sophisticated finite difference modeling of the problem would relax some of the limitations of these models, but experimental data are necessary to validate their accuracy.
Nomenclature C = concentration, /cm3 D = diffusivity, cmP/s k = reaction rate constant, min-' L = reactor length, cm m = P/(1- PO/,) n = (1 + PO)/(l- PO/,) P = (1- €I/€ t = residence time, min U = solids velocity, cm/s V = liquid velocity, cm/s x = length from reactor solids inlet, cm Y = yield yo = particle dimension, cm Greek Symbols cx = kz/kl /5 = k , L / U t = void fraction in reactor bed 0 = porosity (void fraction within particle) 0 = t It, { = r5Dt/4yo
Subscripts A, B, C = reactant (solid), intermediate (sugar), degradation product e = effective (inside wood chip) i = boundary conditions at time greater than 0 1 = liquid 0 = initial condition s = solid 1 = hydrolysis reaction 2 = sugar degradation reaction
Registry No. H2S04,7664-93-9; cellulose, 9004-34-6; glucose, 50-99-7.
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