Ind. Eng. Chem. Process Des. Dev. 1083, 22, 356-361
356
LaMler. K. J.: Buntlng, P. S. "The Chemlcai Kinetlcs of Enzyme Action", 2nd ed.; Clarendon Press: London. 1973. Lee. S. 8.; Klm, S. 8.; Ryn, D. D. Y. Bbteclmol. Bloeng. 1979, 21, 2023. Lee, C. K.; Long, M. E. U.S. Patent 3821086. Messlng. R. A. German Patent 2404353. Nlelsen, M. H. Proceedings, "Chemical Englneering in a Changing World". Etsesevler: Amsterdam, 1978; 183. Nordahl, K.; Thompson, R.: Johnson, A.: Lloyd. N. E. Brbh Patent 1371 489. Park, Y. K.; Thoma, M. J . Food Sci. 1975. 40. 1112. Roek, J. A.: van Tburg, R. St8erke 1979, 91, 17. Salni, R.; Vieth, W. R. J . Appl. Chem. 6iotedulol. 1975, 25, 115. Slpos, T.; HIII, N. J. U.S. Patent 3 708 397. Smith. J. H. "Chemlcal Englneefing Kinetics", 2nd ed.; Mc(3raw-Hlll: New York. 1970 Chapter 11. Sneil, R. L. German Patent 2 539 015.
Stanley,W. L.; Waters, 0. L.; Kelly, S. H. Bbte&wl. 6bwg. 1976, 18, 439. Sbaetsm, J.: Vellenga, K.: de Wm. H. 0. J.; Joosten, 0. E. H. Ind. €ng. Chem. Process. Da9. Dev. 1983, fdkwhg paper In tMs ksue. Strandberg, 0. W.; Smlky, K. L. Bbtechnol. Bkeng. 1972, 14. 509. Tomb, W. H.;Weetai, H. H. U.S. Patent 3957587. VeNenga, K. W.D. Thesis, State Unhrerslty of (LMllngen, The Netherlands, 1978. Vefhaar, L. A. Th.; Dlrkx, J. H. M. &My&. R e s . 1977, 58, 1. Vieth, W. R.; Wang, S. S.; Sainl, R. Blofechnol. Bioeng. 1973. 15. 565. Zienty, M. F. German Patent 2223340.
Received for review March 2, 1981 Revised manuscript received July 13, 1982 Accepted September 23,1982
Isomerization of Glucose to Fructose. 2. Optimization of Reaction Condltions in the Production of High Fructose Syrup by IsomWzation of Glucose Catalyzed by a Whole Cetl Immobilized Glucose Isomerase Catalyst J. Straatsma,' K. Velknga,2H. 0. J. de Wllt,s and 0. E. H. Jooden' Chemical Engineering Department, Nl@nborgh 16, 9747 AG h l n g e n , The Netherlands
The resuits of previous studies have been used to calculate the economically optimal reaction conditions in the production of high fructose syrup by isomerization over an Immobilized catalyst in a packed bed reactor. I t is found that the optimum pH of the feed is 7.65. The sendthhty of the total operation costs for small variations in the pH (0.1 pH unit) Is small. The temperatwe should preferably #e between 50 and 55 O C OT as close as possible to this range while avoiding "biobglcal growth in the system. The catalyst particles should be so small that their effectiveness factor is unity. I f unity cannot be reached for practical reasons, the particles should be as small as the pressure drop over the reactor allows.
Introduction High fructose syrup (HFS) is produced commercially in large quantities by the isomerization of glucose. HFS usually consists of a syrup containing 71% dry matter. A normal composition of the dry matter is 42% (w/w) fructose, 55% (w/w) glucose, and 3% (w/w) polysaccharides. The isomerization is usually carried out in a less concentrated solution. The feed stream to the reactor contains for example 50% (w/w> glucose and 1.5% (w/w) polysaccharides. As a catalyst for the isomerization an immobilized enzyme system is very attractive, combining the advantages of an enzyme catalyst (high selectivity, high activity, etc.) with the well-known advantages of a heterogeneous catalyst (easy separation of catalyst and product stream etc.). The activity of the enzyme catalyst, however, declines during operation, so that eventually it has to be replaced by fresh material. The rate of activity decline depends on the conditions in the reactor as described in part l of this contribution (Straatsma et al., 1983; part 1). In the present contribution the economically optimum conditions in the isomerization reactor leading to minimal cost of isomerization per ton of HFS will be indicated. The costs of isomerization vary with operation conditions due mainly to two factors: the cost of the catalyst
NIZO,Ede. Gist-Brocodes, Delft. 3Douwe Egberts, Utrecht. 0196-4305/83/1122-0356$01.50/0
used in the isomerization and the costs of depreciation, maintenance, and operation of the reactor system. Of course, other costa also contribute to the total isomerization costa, e.g. chemicals, heating, and purification steps. They are not considered here, however, as they hardly depend on the conditions in the reactors and therefore are of little interest for the present purpose. The costs of isomerization are therefore defiied here as the s u m of the total catalyst costa and the total reactor costa. These two groups of costa are different functions of the process parameters such as temperature, pH of the feed, catalyst dimensions, the time after which an aged charge of catalyst is replaced by fresh material, etc. For example, it is clear that the integral production of HFS per kilogram of catalyst increases with the total time the catalyst remains in operation. Hence the catalyst costs per ton of HFS decrease with increasing total operating time per catalyst charge. However, the time-averaged production rate decreasea, so that more or larger reactors are required, leading to increased reactor costs. An economic optimization is therefore required. The same holds for other process parameters. In our laboratory extensive studies have been made of the isomerization of D-glucose into fructose catalyzed by an immobilized whole cell glucose isomerase catalyst. Glucose isomerase is immobilized within the cell as well (Van Keulen et al., 1981). Among others, the kinetics of the reaction (Kikkert et al., 1981),the influence of the immobilization on the kinetics (Van Keulen et al., 1981), the interaction between diffusion and reaction in the 0 1983 American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 3, 1983 357
catalyst particle (Boersma et al., 19791, and the aging of the catalyst (Straatsma et al., 1983, part 1) have been considered. The results of those studies have been used to carry out an economic optimization of the conditions in the isomerization reactor, which is the subject of the present contribution. Few data on production rates of fructose as a function of operating time are available. Only Roels and Van Tilburg (1979) calculated the integral production as a function of total operating time. However, they did not take into account a fall in pH across the reactor. Moreover, they did not calculate the processing costs as a function of the process parameters pH, temperature, particle shape, catalyst price, and yearly reactor costs. Catalyst and Kinetics The catalyst investigated was obtained form the former Dutch company KoninkIijke Scholten Honig (KSH) and had been produced by Reynolds Tobacco Co. and by Imperial Chemical Industries. The method of production of the catalyst is described by Van Keulen et al. (1981). The kinetics of the isomerization of glucose over fresh and aged catalyst can be described by a first-order reversible reaction a t the glucose feed concentrations of commerical intereat (Kikkert et al., 1981). The equilibrium constant K of the reaction is given by
K(T) = F,/G, = 1.06 exp
[--'?(k
&)]
---
(l)
The rate of glucose isomerization per kilogram of dry catalyst is given by
(
-r(G) = k,(G - G,) = k, G - l?K)
(2)
The value of the firsborder rate constant k, of the catalyst is a function of pH, total hexose concentration, temperature, the time the catalyst has been in operation (the operating time, t), and the conditions during operation. For the first-order rate constant a t zero operating time (fresh catalyst, t = 0) the values are used which are obtained by extrapolating the activity at operating times larger than 50 h to zero operating time. These extrapolated values are slightly smaller than the values obtained with fresh Catalyst due to a rapid initial activity decline during the first 50 h of operation (Straatsma et al., part 1). Hence somewhat conservative production rates are obtained in the calculations. In case of a total hexose concentration of 50% (w/w) (3.4 kmol/m3) the relationship between k, pH, and T is kr,O = 4.3 X lo* X ( - O . l 1 8 ( ~ H+ ) ~1.96pH - 7.06) X exp[ 57000 T 1(3)
A)]
-7(
The decline in activity of the catalyst with operating time can be described by a first-order process. The rate constant k, of the isomerization at operating time t is then given by kr,t kr,o exp(-&,t) (4) The rate constant of the aging, ,k is a function of pH and temperature; a t a total hexose concentration of 50% (w/w) k, is given by (Straatsma et al., part 1) k, = (2.5 X 10-3(pH)2- 3.66 X lom2pH- 2.93 X
In the present study two shapes of catalyst particles have
been considered: cylindrical extrudates of 1.7 X m in diameter and 5 X m in length and spheres with a diameter of 0.9 X lo9 m. Due to the relatively high rate of reaction and the limited rate of diffusion of glucose in the catalyst particles the effectiveness factor E of the catalyst may deviate substantially from 1. The value of e is given by (Smith, 1970) t
-{
= 1 1 Th tanh3Th
3Th
Here Th i&the Thiele modulus, being given by (7) Here Pk represents the mass of dry catalyst/m3 of wet catalyst, being equal to 680 kg/m3 for the present catalyst (Vellenga, 1978). Boersma et al. (1979) showed that the effective diffusion coefficient of glucose in the catalyst in a 3.4 kmol/m3 hexose solution is given by
D,(V = 8.8 X
exp
- [--24i00( f &)] -
(8)
With these data the effectiveness factor can be calculated for both catalyst shapes. At zero operating time, 60 "C, and pH 8.20 the effectiveness factors are 0.40 and 0.60 for the cylindrical and spherical particles, respectively. It was observed in the experiments that the pH of the hexose solution decreases when passing through the catalyst bed. The decrease in pH was found to be proportional to the fractional conversion q of glucose, being 0.5 unit at a conversion of 0.84. Here q is defined as Go - G
v=-
(9)
Go - Ge
The local pH is therefore given by 7
pH(q) = pH(? = 0) - - X 0.5 0.84
In the calculations the local value of the pH in the reactor was always used. Reactor System The isomerization of glucose to fructose is carried out in a tubular reactor fdled with catalyst particles. The feed rate to the reactor is adjusted occasionally so that the fructose content of the product stream has the desired value. As the activity of the catalyst decreases with operating time the flow rate to the reactor will have to be decreased gradually. To maintain a more or less constant rate of production, a number of reactors may be operated in parallel, each being in a different stage of aging. The dimensions of each reactor are assumed to be 1m in diameter and a height of 4 m. The reactor is filled with catalyst for 80%. In that case it contains 800 kg of dry catalyst. A reactor of this size was operated successfully on production scale (Vellenga, 1978). It is assumed that the liquid moves in plug flow through the reactor. As the reactor is several hundreds of catalyst particle diameters long, this is reasonable (Smith, 1970). It can be calculated from correlations in the literature for mass-transfer coefficients between liquid and solid in packed beds (Colquhoun-Lee and Stepanek, 1974) and reaction rate data under non-diffusion-inhibited conditions that under process conditions no external mass transfer limitations occur. The local concentration of glucose in the bulk of the liquid as a function of position in the
358
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 3, 1983
reactor is given by the following differential equation
8ooo[
Here E , is the total weight of dry catalyst present at operating time t in the reactor, @ is the volumetric feed rate of glucose to the reactor, t is the local value of the effectiveness factor, x is the length coordinate along the reactor, and H is the height of the catalyst bed. The boundary condition is G = Go at x = 0 (Ilb)
It should be noted that the values of t and k,,depend on position as well as operating time. Composition of Feed and Product Stream The feed to the reactor is assumed to consist of a glucose solution containing 50% (w/w) glucose and 1.5% (w/w) polysaccharides. Furthermore, it is assumed that the polysaccharides behave as inert components, are not converted in the reactor, and have no influence on the kinetics of the isomerization and catalyst stability. The composition of the product stream is specified at 21.7% (w/w) fructose, 28.3% (w/w) glucose, and 1.5% (w/w) polysaccharides. Range of Parameters and Constraints In the optimization of the conditions in the reactor the yearly production of the plant is specified at 100 OOO tons of regular grade HFS (i.e., containing 71% (w/w) dry matter). This means that per year 137860 tons of feed have to be processed. It is assumed that the yearly production is realized in 6500 h. The range of pH values considered is 7.0-8.2; the temperature range is 50-65 "C. A maximum of 4000 h is set to the total time a charge of catalyst remains in operation. This maximum value and the lower limit on the temperature range are dictated by the requirement that microbial contamination of the catalyst should be avoided. The price of the catalyst is varied between $10 and $5O/kg of dry catalyst. The yearly cost of one fixed bed reactor is varied in the range of $5000-15 OOO. Assuming that the yearly capital costs (including maintenance and depreciation) are one-third of the costs of the installed equipment, including auxiliary equipment, this corresponds to a price per reactor of $15000-45 0oO. This range seems reasonable. Calculations The object of the calculations is to find the conditions in the reactor at which the total yearly costs of isomerization (i.e., the sum of catalyst and reactor costs) per ton of HFS are minimal. The yearly catalyst costs follow from the price of the catalyst and the amount of catalyst necessary for the yearly production. The yearly reactor costs simply are the costs per reactor times the number of parallel reactors required to realize the yearly production. When the catalyst has been in operation for t hours, the instantaneous rate of production P of regular grade HFS per kilogram of dry catalyst in a reactor is given by 1 P ( t ) = @(t)pw-(12) E,, Here p is the density of the feed (1225 kg/m3 at 60 "C)and o is the correction factor for the difference in dry matter content of the feed and regular grade HFS: o = 51.5/71. The value of f i t ) is calculated by solving eq 12 numerically under the condition that the liquid has the desired composition a t the exit of the reactor ( x = H). This is an iterative procedure. In the calculations the catalyst bed
5O0C
~
4800
IP (kg H W k g d r y IC1 c a t )
1600
0 t (hrs)
Figure I. Integral production of HFS/kg of dry catalyst (cylinders) as a function of operating time, at pH 7.65 of the feed stream.
0
p
o
I
m
I
ma
1
c90
1
am
1
410
t (hrs)
Figure 2. Integral production of HFS/kg of dry catalyst (spheres) as a function of operating time, at pH 7.65 of the feed stream.
was divided into N slices. It was assumed that within each slice the pH has a constant value. N was taken to be 20. An increase of N to 30 resulted in less than a 0.1% variation in P. The integral production (IP)of HFS per kilogram of dry catalyst in a reactor at operating time t is found from
IP(t) = j k0 ( t ) dt
(13)
The integration is carried out numerically, dividing the operating time space in increments of 50 h. It is assumed that during these 50 h the production rate and pH remain constant. This spacing is sufficiently fine: halving the time increments resulted in a change in the calculated IP value after 4000 h of less than 0.1%. The total number of reactors required for the yearly production follows from the time T a charge of catalyst remains in operation and the value of IP at t = T. From the above data the yearly catalyst and reactor costs and hence the total yearly costs of isomerization can be calculated. Results and Discussion The Integral Production of HFS/kg of Dry Catalyst. The IP was calculated for the cylindrical and spherical catalyst particles as a function of reaction temperature, pH of the feed stream, and operating time. Some
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 3, 1983
--
8.20
7.90
7.60
7.30
0
7.0
Figure 3. The half-life time based on the overall column-averaged apparent firsborder rate constant (&), as a function of pH of the feed stream for cylindrical catalyst particles.
low
-
p(o0
gm
359
(ooo
t (hrs)
Figure 5. The yearly processing costa as a function of the operating time after which a charge of catalyst is replaced (7).The pH of the feed ie 7.65, price of catalyst $lO/kg, yearly costa per reactor $15000.
8
I U Lo
0
1' . -
6S0C
d
3.20
J
7.90
7.60
7.33
7.0
7.5
1
1
7.8
7.7
I 7.6
I
7.5
1 7.b
Figure 4. The half-life time based on the overall column-averaged apparent fiiborder rate constant (&), as a function of temperature and pH of the feed stream for spherical catalyst particles.
Figure 6. The influence of the pH of the feed on the yearly operating costa. Cylindrical catalyst, catalyst price $10/kg, yearly costa per reactor $15 000.
results are given in Figures 1and 2. The intersection of the curves for various reaction temperatures is caused by the opposing effects of temperature on activity and stability: with increasing temperature the activity increases, whereas the stability decreases (eq 3, 4,5). A comparison between Figures 1and 2 shows that the integral production using spherical catalyst particles is (considerably) larger than of cylindrical particles. This is due to the fact that the effectiveness factor of the spherical catalyst particles is greater than that of the cylindrical catalyst particles used. The activity of the catalyst in the reactor does not decrease at the same rate throughout the column due to the pH gradient. At any value of t it is possible to calculate a column-averaged apparent first-order rate constant ek,. In Figures 3 and 4 the half life times of the catalyst in the reactor based on tk,are shown as a function of reaction temperature and pH of the feed stream for the cylindrical and the spherical catalyst particles, respectively. From the figures it appears that for each temperature and both particle shapes the maximum half-life time is reached at a pH of the feed stream of 7.55. This value corresponds with a pH of the product stream of 7.05.
It also follows from Figures 3 and 4 that the observed half-life times of the spherical catalyst particles are lower than those of the cylindrical particles. This is due to the higher effectiveness factor of the former. (Straatsma, part 1). Processing Costs. The processing costs have been calculated as a function of a number of process parameters: pH, temperature, particle shape, catalyst price, yearly reactor costs, and the operating time 7 after which a charge of catalyst is replaced by fresh material. It appeared that the processing costs as a function of 7 pass through a minimum for all values considered for the pH, temperatures, catalyst prices, and yearly costs of one reactor. In some cases this minimum is reached for values of 7 larger than 4000 h, however. Figure 5 gives a typical example. At constant valuea of all other parameters, the value of this minimum varies with the pH of the feed. It was found that, virtually independent of the values of the other parameters, the optimal pH is 7.65. The minimum is rather flat so that a pH range of 7.6 to 7.7 of the feed is allowed without influencing the processing costa significantly; see Figure 6. The pH value of 7.65 for optimal operation is 0.1 unit higher than the value for maximum catalyst stability. From an economic point of view it is apparently
360
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 3, 1983
Table I. Economically Optimum Lifetime as a Function of Temperature for Cylindrical Catalyst Particles at the Optimum pH of the Feed Stream, a Catalyst Price of $10/kg, and Yearly per Resctor of $15 000
T," C
optimum operating time, h
50 55 60 65
>4000 2950 1900 1200
column-av residual activity no. of ( % of init activity) reactors 45 31 22 13
a Costs after an operating time of 4000 h. kg of HFS.
tot reactor costs, l o 5 $
tot cat. costs, lo5 $
2.5O 2.2 1.9 1.7
2.3O 2.7 3.6 5.0
17 15 13 12
These costs multiplied by
year1y.b processing costs, lo5 $ 4.8O 4.9 5.5 6.7
are equal to the processing costs per
lo3
Table 11. Economically Optimum Lifetime as a Function of Temperature for Spherical Catalyst Particles at the Optimum DH of the Feed Stream. a Catalvst Price of $10/ka and Yearlv Costa Der Reactor of $15 000
T,"C
opt. operating time, h
50 55 60 65
>4000 2650 1750 1100
column-av residual activity no. of (% of init activity) reactors 37 27 17 9
Costs after an operating time of 4000 h.
tot reactor costs, lo5 $
tot cat. costs, lo5 $
year1y.b processing costs, 10' $
2.1" 1.7 1.4 1.2
l.ga 2.2 2.7 3.7
3.9O 3.9 4.1 4.9
14 11 9 8
These costs multiplied by
are equal to the processing costs per l o 3 kg
of HFS. Table 111. Influence of Catalyst Price and Yearly Costa per Reactor on Economically Optimum Lifetime for Cylindrical Particles. T = 60 " C ;PH of Feed is 7.65 cat. price, $/kg of dry cat.
yearly costs per reactor, $
opt operating time, h
column-av resid activity, % of init no. of activity reactors
tot reactor tot cat. costs, lo5 $ costs, l o 5 $ ~
yearly processing costs,
io5 $
~~
10 20 30 40 50
1 5 000 15 000 15 000 15 000 15 000
1900 2300 2600 2800 2950
22 16 12 10 9
13 15 16 17 18
1.9 2.3 2.4 2.5 2.7
3.6 6.5 10.0 13.0 16.0
12.4 15.5 18.7
10
5 000 5 000 5 000 5 000 5 000
2600 3050 3350 3550 3750
12 9 7 6 5
16 19 20 21 23
0.8 0.9 1.0 1.1 1.2
3.3 6.5 9.5 12.5 15.5
4.1 7.4 10.5 13.6 16.7
20 30 40 50
attractive to sacrifice some catalyst stability in favor of catalytic activity. The activity exhibits a maximum at pH 8.3 (Kikkert et al., 1981). From Figure 5 the economically optimum value of T at a number of temperatures and at the optimum pH of 7.65 can be deduced. The results are given in Table I. Also given are the column-averaged residual catalyst activity at t = 7,the number of reactors required, the total reactor and catalyst cost8, and the processing costs. The residual activity is rather small in a number of cases. The same calculations have been carried out for spherical catalyst particles. These results are given in Table 11. Table I shows that the economically optimum life time T of a catalyst charge decreases with increasing temperature. This is not unexpected considering the integral production of HFS per kilogram of dry catalyst as a function of temperature (Figure 1). From Table I it follows also that for minimal processing costs the isomerization has to be carried out at 50 O C when cylindrical particles are used; for spherical particles the optimum temperature lies between 50 and 55 O C . A comparison between Tables I and I1 shows that the yearly processing costs when spherical catalyst particles are used are about 75-80% of those using cylindrical particles. Thii is caused by the higher integral production of HFS per kilogram of catalyst (Figures 1and 2), which in turn is due to the higher effectiveness factor of the spheres. Hence when altering the shape and size of the
5.5 8.8
particles the aim should be to increase the effectiveness factor. In Table I11 the influence of the catalyst price and the reactor costa on the economically optimal value of T k given for a pH of the feed of 7.65 and a temperature of 60 "C. As expected, the optimum value of 7 increaw with catalyst price. As at a constant temperature the integral production of HFS per kilogram of catalyst varies less than linearly with T it is expected that the number of reactors required increases with T , which is confirmed by the data in Table 111. Table I11 shows also that the number of required reactors decreases when the yearly costs per reactor increase; this is as expected. As a consequence of this, it is expected that the optimal value of 7 decreases and therefore the total quantity of required catalyst increases. This is confirmed by Table 111. From Table I11 it can be concluded that the sensitivity of the yearly processing costs for changes in the catalyst price is much larger than for changes in the yearly costs per reactor; the processing costs are roughly proportional to catalyst price to the power 0.8 and to the yearly costs of one reactor to the power 0.15. Conclusion Based on quantitative information on the kinetics and stability of the catalyst particles, the conditions have been calculated under which the processing costs are minimal. The numerical integration of the relevant equations indicated that: (a) over the whole range of parameters
Ind. Eng. Chem. Process Des. Dev. 1903, 22, 361-366
considered the optimum pH of the feed is 7.65; (b) the sensitivity of the total operation costs for small variations in the pH (0.1 pH unit) is small; (c) the temperature should preferably lie between 50 and 55 O C , or as close as possible to this range, while avoiding microbiological growth in the system; (d) the catalyst particles should be as small as the pressure drop over the bed allows, when the effectiveness factor is smaller than one; and (e) the total operation costs are sensitive to variations in the catalyst price, rather insensitive to variations in the yearly costs per reactor. Nomenclature De = effective diffusion coefficient in a catalyst particle, mz/s E,, = total weight of catalyst in a reactor, kg of dry cat. G = glucose concentration, kmol/m3 Go = glucose concentration in the feed solution, kmol/m3 G, = glucose concentration in a solution which is in thermodynamic equilibrium, kmol/m3 H = height of the catalyst bed, m K = thermodynamic equilibrium constant of the glucose to fructose isomerization k, = first-order rate constant, m3/s kg of dry cat. L = ratio between volume and external surface of a catalyst particle, m f i t ) = production rate of HFS per kg dry catalyst at operating time t, kg of HFS/s kg of dry cat. R = gas constant, J/K mol
36 1
Th = Thiele modulus (see eq 6) t = operating time, h x = axial coordinate in the catalyst bed, m c
= effectiveness factor
Pk = weight of dry catalyst/m3wet catalyst, kg of dry cat./m3
of wet cat. density of HFS,kg/m3 = flow rate, m3/s = fractional conversion (eq 9) = time after which a charge of catalyst is replaced, h Registry No. Glucose isomerase, 9055-00-9.
p = 9 T
Literature Cited Boersma, J. 0.;Vellenga, K.; de Wllt, H. G. J.; Joosten, G. E. H. Blofechnol. B h n g . 1070, 21, 1711. Colquhoun-Lee. I.; Stepanek, J. Chem. Eng. 1974, 1 , 108. Keulen, M. van; Veilenga, K.; Joosten, 0. E. H. Bbfechnol. B h n g . 1981. 23, 1437. Kikkert. A.; Vellenga, K.; ce ! WIR, H. G. J.; Joosten, G. E. H. Blofechnol. B h n g . 1981, 23, 1007. Roels, J. A,, van filburg, R. Sfaerke 1070, 31. 17. Smith, J. M. “Chemical Engineering Kinetics”, 2nd ed.; McGraw-HiII: New York, 1970; Chapter 11. Straatsma, H.; Vellenga, K.; de Wilt, H. 0. J.; Joosten. G. E. H. Ind. Eng. Chem. Process Des. Dev. 1089, preceding paper in this issue. Vellenga, K. Ph.D. Thesis State Unlversky Groningen. The Netherlands, 1978.
Received for review March 2, 1981 Revised manuscript received July 13, 1982 Accepted September 23, 1982
Control of High-Purity Distillation Columns Carmelo Fuentes and Wllllam L. Luyben’ Department of Chemical Engineedng, Lehigh University, Bethiehem, Pennsylvanla 180 75
The dynamic behavior of distlllation columns with high-purity products (down to 10 ppm) has been studied via digital simulation. The effects of product purity, relative volatility, composition analyzer sampling time, and magnitude of disturbance have been explored. Results show that systems with low relative volatility (a= 2) respond slowly enough so that good control can be achieved at very high purity levels with a 5-min analyzer dead time. However, systems with high relative volatility (a= 4) respond so quickly that large deviations in product purities occur before the analyzer can respond. The dynamic response of the column is very nonlinear. Effective control was obtained by using a composition/temperature cascade system. An intermediate tray temperature was controiled to achieve fast dynamic response to disturbances, and the setpoint of the temperature controller was reset from a product composition controller. The secondary temperature controller gave better control for feed composition disturbances when it was proportional only and loosely tuned. The opposite was true for feed rate disturbances.
Introduction The control of distillation columns is probably one of the most studied areas of process control. Tolliver (1980) has recently compiled a useful comprehensive list of references. Most of the columns studied have had product purities that were low to moderate (0.1 to 5 mol 90impurity). Very few papers consider the dynamics and control of highpurity columns, despite their industrial importance. Boyd (1975) used a double differential control scheme to maintain overhead purities of about 10 ppm in a benzene/ toluene separation. ’&reus (1976) reported difficult control problems for a methanol/water column with product purities in the range of lo00 ppm. A highly nonlinear behavior was reported. The control problems with high-purity columns have been so severe that many process designers try to avoid
making high-purity products out of both ends of a column simultaneously. It is a very common practice in the chemical industry to build two columns instead of one and to provide large intermediate tanks to handle recycle flows between these columns. Pure products are produced out of one end of each column. This practice increases both capital investment and energy costs. Thus, there is considerable economic incentive to improve the control of high-purity columns. The purpose of this paper is to explore the dynamics and controllability of high-purity distillation columns. The dynamic behavior was investigated by linear analysis and by digital simulation of a nonlinear mathematical model. The dynamic responses of the open-loop system for changes in various manipulated and disturbance variables were first studied in order to gain some insight into the dynamic difficulties associated with the control of these
01~~-43051831 I 212 - 0 3 ~ 1 ~ 0 1 . 5 0 1 00 1983 American Chemical Society
