426 Fermentation and Biochemical Engineering Handbook 5.0 PROCESS CONSIDERATIONS 5.1 Design Factors The engineer designing an ion exchange column operation usually will prefer to work with the simplest kinetic model and linear driving force approximations. The weakness ofthis approach is that any driving force law only regards the momentary exchange rate as a hnction of the solute concentration in the bulk solution and the average concentration in the particle, neglecting the effect of concentration profiles in the particle. Nevertheless, the linear driving force approach provides an approximation that is sufficiently accurate for the engineer. 5.2 Scaling-up Fixed Bed Operations Rodrigues[66] has presented empirical and semi-empirical approaches which may be used to design ion exchange columns when the solute in the feedstream is co and the flow-rate is uo. The breakthrough point is usually set at the point where the effluent concentration increases to 5% of co. The design equations relate the total equilibrium ion exchange capacity (e) to the volume of resin required (V,) to the time of breakthrough (b). In the empirical approach, the overall mass balance is given by the equation: where Eq. (16) t, = 5 (1 + 5) (the stoichiometric time) Eq. (17) 5 = (1- &)e/& Qo (the mass capacity factor) Eq. (18) 5 = &VhO (the space time) and Vis the bed volume with void space E. It is usually necessary to modify this resin amount by a safety factor (1.2 to 1.5) to adjust for the portion of the total equilibrium capacity that can actually be used at flow rate u and to adjust for any dispersive effects that might occur during operation. Ion Exchange 427 The semi-empirical approach involves the use of the mass transfer zones. This approach has been described in detail specifically for ion exchange resins by Passin0.[~'1 He referred to the method as the operating line and regenerating line process design and used a graphical description to solve the mass transfer problems. For the removal of Ca" from a feedstream, the mass transfer can be modeled using Fig. 21. The upper part shows an element of ion exchange column containing a volume v of resin to which is added a volume V, of the feedstream containing Ca". It is added at a flow rate (Q) for an exhaustion time t, . The concentration of Ca* as it passes through the column element is reduced from xal to x,,. Therefore, the resin, which has an equilibrium ion exchange capacity Cy increases its concentration of Catt fromy,, to yal. In this model, fresh resin elements are continuously available at a flow rate (F') = v/t, , which is another way of saying the mass transfer zone passes down through the column. The lower part of Fig. 21 shows the operating lines for this process. The ion exchange equilibrium line describes the selectivity in terms of a Freundlich, Langmuir or other appropriate model. The equations for the points in the lower part are given by: Eq. (19) Xexl =xo (Ca" in the feedstream) (Ca" in the exhausted streamstream) x dv (average Ca* in the effluent) x,, = - v, Eq. (21) Eq. (22) Yd=O (Ca" in the regenerated resin) and the slope of the operating line: 428 Fermentation and Biochemical Engineering Handbook .- 5 ti; C Q) 01 E -. Ion Exchange 429 The value of c,,, C and v are known so that for any V, value, the slope of the operating line can be calculated from Eq. 23. The specific points: x, is given, x, is obtained by graphic integration fromthe breakthrough curves. After operation and regeneration, the value ofy,, may not be zero but may be between 0.02 and 0.05 ifthe regeneration is not complete. The application of this technique has been described in terms of basic design parameters such as number of transfer units, the height of a transfer unit and mass transfer coefficient. [as1[691 The data generated with the laboratory column may be scaled-up to commercial size equipment. Using the same flow rate (on a mass basis) as used in the laboratory experiments, the appropriate increase in column size over that used in the laboratory is a direct ratio of the volumes to be treated compared to that treated in the laboratory equipment. If a reasonable height to diameter ratio (approximately 1 : 1) is obtained in the scaleup using the bed depth involved in the laboratory procedure, then that bed depth is maintained and the cross-sectional area of the column is increased. However, if the sizing is such that the column is much larger in diameter than the bed depth, scaleup should be done to maintain a height to diameter ratio of approximately 1. The required resin volume is determined by maintaining the same mass flow conditions (liters of feed solution per minute per cubic meter of installed resin) as was used in the laboratory operation. Appropriate tank space must be left to accommodate the backwash operation. This is typically 50% of bed depth for cation exchange resins and 100% expansion in the case of anion exchange resins. 5.3 Sample Calculation The purification of lysine-HC1 from a fermentation broth will be used to illustrate the calculations involved in scaling-up laboratory data. The laboratory fermentation broth, which is similar to the commercial broth, contained 2.0 g/O. 1 1 lysine, much smaller amounts of Caw, Kf and other amino acids. The broth was passed through 500 ml of strong acid cation resin, DowexB HCR-S, in the NH,’ form. The flow rate was 9 ml/min or 1.77 mVmin per cm2 of resin. It was determined that the resin capacity averaged 1 15 g of lysine-HC1 per liter of resin. It may be noted that since the equivalent molecular weight of lysine-HC1 is 109.6 g and the “theoretical” capacity of DowexB HCR-S is 2.0 meq/ml, the operating capacity is 52% of theoretical capacity. 430 Fermentation and Biochemical Engineering Handbook The commercial operation must be capable of producing 9,000 metric tons of lysine (as lysinedihydrochloride-H,O) per year. With a 2.0 g/O. 1 1 concentration of lysine in the fermentation broth, the number of liters of broth to be treated each year are: 9,000 m tons 146.19 (MW of lysine) X Yr 237.12 (MWofL-HCl-H,O) x- Os1 x- lo6g = 27.7 x 107 i/p 2.0g Mton If the plant operates 85% of the time, the flow rate would have to be: 27.7 x lo7 l/yrxlx 1Yr 0.85 365 days 1 day 24 hr -=3.73x104 l/hr At a resin capacity of 1 15 g/1 of resin, the amount of resin that must be available is: l(resin) 219.12(MWof L-HCl) x-x 2.0g 115 g 146.19(MWof L) 0.1 1 3’73x1041/hr =9.71~103 l/hr hr To obtain the maximum utilization of the resin in this operation, series bed operation (Carrousel) operation is recommended. This operation uses three beds of resin in a method having two beds operating in series while the product is being eluted from the third. The freshly regenerated resin is placed in the polishing position when the totally loaded lead bed is removed for regeneration. Ion Exchange 431 The elutionhegeneration step, which includes backwashing, eluting, and rinsing the resin, might take up to four hours. Therefore, enough resin must be supplied to take up the lysine-HC1 presented during that time. The resin requirement for the commercial scaleup operation would be: 9.71 x lo3 l(resin) 4 hr hr bed x-=3.88x104 l(resin)/bed Eq. (27) Thus, three beds of 39 m3 resin each are required to produce 9,000 metric tons of lysine/year. 5.4 Comparison of Packed and Fluidized Beds Belter and co-w~rkers[~~] developed a periodic countercurrent process for treating a fermentation broth to recover novobiocin. They found that they were able to scaleup the laboratory results to production operations ifthe two systems have similar mixing patterns and the same distribution of residence times in the respective columns. The mixing patterns are the same when the space velocity (FN,) and the volume ration (&IC) are the same. This is shown in Fig. 22 for the effluent concentration of novobiocin from laboratory and production columns. The scaling-up of packed beds is subject to the difficulties of maintain- ing even flow distributions. Removal of solution through screens on side walls is not recommended and the flow of resin from one section into another of much greater area could distort the resin flow profile. The problems of scaling-up fluidized bed operations are more difficult in terms of design calculations, but flow distribution is more easily designed because of the mobility of the resin. The degree of axial mixing of the liquid and the resin has to be taken into account when calculating the changes necessary in the bed diameter and bed height. Figure 23 shows the increases in bed height necessary when scaling-up packed and fluidized beds with bed diameter Mass transfer coefficients have been correlated for packed and fluid- ized beds.[72] The mass transfer coefficients for packed beds are 50 to 100% greater than for fluidized beds. The volume of resin in a packed bed is about halfthat in a fluidized bed, but the packed bed column may be up to eight times smaller. Despite this, a complete fluidized bed operation may still be smaller than a fixed bed operation. Also, fluidized bed columns do not operate at high pressures so they can be constructed more economically. 432 Fermentation and Biochemical Engineering Handbook 01 ‘1 1’ I’ ” 0 30 60 90 120 150 110 210 240 210 TIME IN HlNUTfS Figure 22. Comparison of experimental curves for laboratory and production column^.^^^] Figure 23. Scaleup relationships for fluidized beds.[7’] Ion Exchange 433 5.5 Chromatographic Scale-up Procedures The aim of scaling-up a chromatographic process is to obtain the same yield and product quality in the same time period on laboratory, preparative and industrial scale, Laboratory analytical purifications tend to be optimized only for resolution of individual solutes, however, at the preparative and production scale, it is necessary also to maximize throughput. Table 14[731 shows the effect of chromatographic operating parameters on resolution and throughput. While column length is a critical factor for resolution with gel filtration and isocratic elutions, it has little effect on resolution with gradient elution in adsorption chromatography. The wall effects on resolution are very noticeable with small radius columns, but decrease as the column length is increased. Table 14. Effects of Process Parameters on Resolution and Throughput[73] Resolution Throughput Parameter varies with varies with Column Length (L) L I/L Column Radius (r) Some effect r2 Temperature (T) Positive effect T Viscosity (q) Negative effect 1h Sample Volume (V) 14 ‘-‘optimum I V Flow Rate (J) 1WJoptimuml J Voser and Walli~er[~~] viewed scaleup as a three step process involving selection of a process strategy, evaluation of the maximum and optimal bed height and finally column design. The selection ofa process strategy involves choosing the direction of flow, frequency of backwashing, the operation with positive head pressure or only hydrostatic pressure, and the use of a single column or a series of columns in batch or semi-continuous operation. The 434 Fermentation and Biochemical Engineering Handbook maximum feasible bed height is determined by keeping the optimal labora- tory-scale specific volume velocity (bed volumehour) constant. The limiting factors will be either pressure drop or unfavorable adsorptioddesorption kinetics since the linear velocity also increases with increasing bed height. Bed heights from 15 cm to as high as 12 m have been reported. The column diameter is then selected to give the required bed volume. The column design combines these column dimensions with the practical considerations of available space, needed flexibility, construction difficulty and flow distribu- tion and dilution for the columns. The scaleup considerations of column chromatography for protein isolation has been described by Charm and Matte0.1~~1 When several hundred liters of a protein feedstream must be treated, the resin may be suspended in the solution, removed after equilibration by filtration and loaded into a column from which the desired proteins may be eluted. Adsorption onto a previously packed column was not recommended by them since they feared suspended particles would clog the interstices of the column, causing reduced flow rates and increased pressure drops across the column. The reduced flow rate may lead to loss of enzyme activity because of the increased time the protein is adsorbed on the resin. It is important that all of the resin slurry be added to the column in one operation to obtain uniform packing and to avoid the formation of air pockets in the column. An acceptable alternative, described by Whatman,[76] allows the addition ofthe adsorbent slurry in increments. When the resin has settled to a packed bed of approximately 5 cm, the outlet is opened. The next increment of the slurry is added after the liquid level in the column has dropped. It is important that the suspended adsorbent particles do not completely settle between each addition. Stacey and coworkers[77] have used the relationships shown in Table 15 to scaleup the purification from an 8 mg protein sample to a 400 mg sample. The adsorbent used in both columns was a Delta-Pak wide-pore C- 18 material. When eluting the protein, the flow rate should change so that the linear velocity ofthe solvent through the column stays the same. The flow rate is proportional to the cross-sectional area of the column. The gradient duration must be adjusted so that the total number of column volumes delivered during the gradient remain the same. As with size exclusion chromatography, the mass load on the preparative column is proportional to the ratio of the column volumes. Figure 24 shows that the chromatograms from the 8 mg separation is very similar to that obtained for the 400 mg sample. Ion Exchange 435 .4 . Table 15 Scale-up Calculations[77] 8mg Sample Small Scale Preparative Scale Column Dimensions 0.39 x 30 cm 3 x 25 cm Flow Rate Scale Factor (3.0)* = 59 (0.39)2 Sample Load Scale Factor = 49 (3.0)2 x 25 (0.39)2 x 25 1.5 mVmin 90 ml/min 400 mg Gradient Duration Calculation 40 min 33 min column 90 x (Gradient Duration) - 16.7 1'5x40 46.7 - (0.195)2 zx 30 Volume (1.5)2lcx25 Grad. Duration = 33 min. 0.8 f 1 O" 0 8 Minutes Figure 24. Laboratory and preparative scale separation of cytochrome C digest.[77] 436 Fermentation and Biochemical Engineering Handbook Ladisch['*I has worked with a variety of column sizes ranging from 2 to 16 mm in diameter and 10 to 600 cm in length. His experience is that published semiempirical scaleup correlations are usefbl in obtaining a first estimate on large scale column performance. When scaling-up a chromatographic process, it may be necessary to change the order of certain steps from that used in the laboratory. Gel filtration, though a frequent first step at the laboratory scale, is not suitable for handling large scale feedstream volumes.[79] When gel filtration is used to separate molecules of similar molecular weights, sample sizes may range from 1% to 5% of the total gel volume. Thus, a 100 liter feedstream would require a gel filtration column of 2,000 to 10,000 liters. When one is separating a large molecule from small molecules, as in desalting operations, the applied volume may be up to 30% of the gel volume. On the other hand, ion exchange chromatography is a very good first step because its capacity is approximately 30 mg of protein per ml of resin. This capacity is relatively independent of feed volume. For the same 100 liter feedstream, only a 20 liter ion exchange column would be required. 5.6 Pressure Drop The pressure drop across an ion exchange bed has been represented by an equation[80] which depends on the average particle diameter, the void fraction in the bed, an exponent and a friction factor dependent on the Reynolds number, a shape factor, the density of the fluid, the viscosity of the fluid and the flow rate. While that equation has internally consistent units (English system), the variables are not normally measured in those units. Another disadvantage is that one must check graphs of the exponent and friction factor versus the Reynolds number to use the equation. For laminar flow with spherical particles, the equation can be simpli- fied to: AP O.0738p(cp)V0 (Ilmin) (1- E)~ -(bar/cm)= Eq. (28) L Dp2 (mm2) E3 For most ion exchange resins, the void volume is about 0.38, so that (1- 2)l 2 = 4.34 and: Ion Exchange 437 AP - (bar / cm) = Eq. (29) L D; (mm2) 0.32~ (cp) V, (I / min) Table 16 shows the agreement between results from experiments and those calculated with this equation for several ion exchange resins. Table 16. Pressure Drop for Commercial Ion Exchange Resins Resin Flow Rate Mean Bead Pressure Drop (barkm) (L/min) Diameter (mm) Calculated Measured Dowex SBR-P 15 1.4 0.750 5.28 5.08 Dowex WGR-2 22.7 0.675 1 .oo 1.06 Dowex MWA-1 15.1 0.675 0.67 0.77 DowexMSA- 1 37.8 0.650 1.79 2.13 DowexMSC- 1 30.3 0.740 1 .os 1.22 Figure 25 shows the use of a size factor for resins which is used to develop a pressure factor. This pressure factor can then be used to calculate the pressure drop under conditions of different solution viscosities, flow rates or particle size distributions once the pressure drop is known at one viscosity, flow rate and particle size distribution. 438 Fermentation and Biochemical Engineering Handbook .d B x LDoooooooo O~N+Wod~d$ !I %I 2 I: n a Ion Exchange 439 5.7 Ion Exchange Resin Limitations When ion exchange resins are used for an extended period of time, the exchange capacities are gradually decreased. Possible causes for these decreases include organic contamination due to the irreversible adsorption of organics dissolved in the feedstream, the oxidative decomposition due to the cleaving ofthe polymer cross-links by their contact with oxidants, the thermal decomposition of functional groups due to the use of the resin at high temperature and the inorganic contaminations due to the adsorption of inorganic ions. When a resin bed has been contaminated with organic foulants, a procedure is available that can help to restore the resin.[*l] Three bed volumes of 10% NaCl - 2%NaOH solution are passed through the resin bed at 50°C. The first bed volume is passed through at the same flow rate, followed by a through rinsing and two regenerations with the standard regenerate. If the fouling is due to microbial contamination, the same authors recommend backwashing the bed and then filling the entire vessel with a dilute solution (< 0.05%) of organic chlorine. This solution should be circulated through the bed for 8 hours at a warm (50°C) temperature. After this treatment, the resin should be backwashed, regenerated and rinsed before returning it to service. This procedure may cause some oxidation of the polymeric resin, thereby reducing its effective cross-linking and strength. Therefore, treating the resin in this fashion should not be a part of the normal resin maintenance program. Physical stability of a properly made cation or anion exchange resin is more than adequate for any of the typical conditions of operation. These resins can be made to have a physical crush strength in excess of 300 grams per bead. Perhaps more important are the limitations inherent in the structure of certain polymers or functional groups due to thermal and chemical degrada- tion. Thermally, styrene-based cation exchange resins can maintain their chemical and physical characteristics at temperatures in excess of 125°C. At temperatures higher than this, the rate of degradation increases. Operating temperatures as high as 150°C might be used, depending on the required life for a particular operation to be economically attractive. Strong base anion exchange resins, on the other hand, are thermally degraded at the amine functional group. Operating in the chloride form, this is not a severe limitation, with temperatures quite similar to those for cation exchange resins being tolerable. However, most strong base anion exchange resins used involve either the hydroxide form, the carbonate or bicarbonate form. In these ionic forms, the amine functionality degrades to form lower 440 Fermentation and Biochemical Engineering Handbook amines and alcohols. Operating temperatures in excess of 50°C should be avoided for Type I strong base anion exchange resins in the hydroxide form. Type I1 strong base resins in the hydroxide form are more susceptible to thermal degradation and temperatures in excess of 35°C should be avoided. The amine functionality of weak base resins is more stable in the free- base form than that of strong base resins. Styrenedivinylbenzene weak base resins may be used at temperatures up to 100°C with no adverse effects. Chemical attack most frequently involves degradation due to oxida- tion. This occurs primarily at the cross-linking sites with cation exchange resins and primarily on the amine sites of the anion exchange resins. From an operating standpoint and, more importantly, from a safety standpoint, severe oxidizing conditions are to be avoided in ion exchange columns. Oidizing agents, whether peroxide or chlorine, will degrade ion exchange resins.[82] On cation resins, it is the tertiary hydrogen attached to a carbon involved in a double bond that is most vulnerable to oxidative degradation. In the presence of oxygen, this tertiary hydrogen is transformed first to the hydroperoxide and then to the ketone, resulting in chain scission. The small chains become soluble and are leached from the resin. This chain scission may also be positioned such that the cross-linking of the resin is decreased, as evidenced by the gradual increase in water retention values. The degradation of anion resins occurs not only by chain scission, but also at the more vulnerable nitrogen on the amine functionality. As an example, the quaternary nitrogen on Type I strong base anion resins is progressively transformed to tertiary, secondary, primary nitrogen and finally to a nonbasic product. Oxidative studies on resins with different polymer backbones and hctionalities have been performed as accelerated tests.[83] The data is shown in Table 17 for polystyrene and polydiallylamine resins. Although the polydiallylamine resins have a higher initial capacity, they are much more susceptible to oxidative degradation. When the polystyrene resin has a mixture of primary and secondary amino groups or when a hydroxy- containing group is attached to the amine of the functional group, the susceptibility to oxidation is enhanced. Thus one can understand the lower thermal limit for Type I1 anion resins compared to Type I resins. The effect of thermal cycling on strong base anion resins has been studied by Kysela and bra be^.[*^] The average drop per cycle in strong base capacitywas 2.1 x 10-4mm01 (0H~)/m1overthe480cyc1esbetween20"Cand 80°C. Figure 26 shows the decrease in total exchange capacity (open circles) and in strong base (salt splitting) capacity (solid circles) for each of the individual resins included in the study. Ion Exchange 441 Table 17. Oxidation of Polystyrene and Polylamine Resins in a One-Week Accelerated Test at 90°C[831 Polystyrene Polystyrene Polystyrene Polydiallylamine Polydiallylarnine Polydiallylamine Initial Base R-CHzN(CH3)z 4.4 R-CH~N(CHZCH~)~ 4.5 R-CH~N(CZH~OH)~ 3.6 N-H 8.3 N-CHzCHs 7.3 N-CaH7 6.7 Base Capacity Lost 1 2 17 37 37 31 As would be expected, the combination ofthermal and osmotic shocks has been shown to have a large effect on the osmotic strength of anion resins.[85] This is illustrated in Fig. 27, where the decrease is shown to be over 15% in just 150 cycles. The important processing point to remember from this is that if the regeneration of a resin is performed at a temperature more than 2OoC different from the temperature of the operating (loading) stream, it is advisable first to adjust the temperature of the resin with distilled water until it is at the temperature of the regenerant. 5.8 Safety Considerations Safety considerations in the use of ion exchange systems include understanding the reactions which will take place during the contact of resin and solution being treated. The possibility of reactions being catalyzed in the presence ofthe hydrogen form of a strong acid cation or of the hydroxide form of the strong anion must be studied. These resins produce acid or base, respectively, during treatment of salt solutions. If reactants or products of the ion exchange reaction are hazardous, appropriate protective equipment should be included in the system design. For example, during the demineralization of metal cyanide by-products, hydrogen cyanide solution is an intermediate formed in the cation exchange unit. The processing equipment must be designed to ensure containment of this solution until it has passed through the anion exchange unit which picks up the cyanide ions. 442 Fermentation and Biochemical Engineering Handbook I I I I I WIW AI 1.3 I I I I Figure 26. Dependence of the strong base capacity on the number of cycles between 20°C and 80°C.[841 Ion Exchange 443 \4 83 - 80 0 0 50 100 150 NWER OF CYCLES Figure 27. The effect of thermal and osmotic cycling on the osmotic strength of a strong base anion resin, AV-17. (I) Effect of thermal cycling; (2) effect of osmotic cycling; (3) effect of thermal and osmotic cycling (cold alkali to hot acid); (4) effect of thermal and osmotic cycling (hot alkali to cold acid).rs5I Since the ion exchange resins are synthetic organic compounds, they are susceptible to decomposition reactions with strong oxidizing agents, such as nitric acid. Close control of reactant or regenerant concentrations, temperature, and contact time may be required. When the treated solution is to be used as food, food additives, or in pharmaceutical applications, resins must be used which are acceptable for such processing. 6.0 ION EXCHANGE OPERATIONS The typical cycle of operations involving ion exchange resins include pretreatment of the resin and possibly of the feed solution; loading the resin 444 Fermentation and Biochemical Engineering Handbook with the solutes to be adsorbed by contacting the resin with the feed solution; and elution of the desired material from the resin. The scale of operation has ranged from analytical applications with a few milligrams of resin and microgram quantities of material to commercial production units containing several cubic meters of resin to produce metric tons of material. The loading may be applied batchwise, to a semi-continuous batch slurry, or to a column filled with resin which may be operated in a semi-continuous, continuous, or chromatographic manner. A list of designers and manufacturers of ion exchange equipment is given in Table 18. The list is weighted toward those companies that operate in the United States. Table 18. Ion Exchange Process Designers and Manufacturers omDanv Location Belco Pollution Control Corp. Chern Nuclear Systems Cochrane Divison of Crane Cop. Downey Welding & Manufacturing Co. Envirex Corp. Epicor Emo Water Conditioners Hittrnan Nuclear & Development Hungerford & Terry, Inc. Industrial Filter and Pump lnfilco Degrernont, Inc. lntensa Kinetico, Inc. Kurita Liquitech, Div. of Thermotics, Inc. Mannesrnann Mitco Water Labs, Inc. Permutit, Division of Zum Rock Valley Water Conditioning Techni-Chern, Inc. Unitech, Divsion of Ecodyne United States Filter, Fluid System Corp. U. S. FiIterllWT Wynhausen Water Softener Company Parsippany, NJ Columbia, SC King of Prussia, PA Downey, CA Waukesha, WI Linden, NJ St. Louis, MO Columbia, MD Clayton, NJ Cicero, IL Richmond, VA Mexico City, Mexico Newbury, OH Japan Houston, TX Gemany Winter Haven, FL Warren, NJ Rockford, IL Belvidere, IL Union, NJ Whittier, CA Rockford, IL Los Angeles, CA Ion Exchange 445 6.1 Pretreatment Filtration, oil flotation and chemical clarification are the pretreat- ment operations commonly employed with the feed stream. This pretreat- ment is primarily concerned with the removal of excess amounts of suspended solids, oils, greases and oxidative compounds. Suspended solids, including bacteria and yeasts, in amounts exceeding approximately 50 mg/L should be removed prior to applying the fluid to the column to prevent excessive pressure buildup and short operating cycle times. The presence of oils and greases in excess amounts would coat the resin particles, thereby dramatically reducing their effectiveness. Oils and greases in concentrations above 10 mg/L should not be applied to resins in either column or batch operations. Synthetic resins are subject to de- cross-linking if oxidative materials are present in the feed solution or eluant. For many biochemical recovery applications, it is necessary to pretreat the resin to ensure that the extractable level of the resin complies with Food Additive Regulation 2 1 CFR 173.25 of the Federal Food, Drug and Cosmetic Act. The pretreatment recommended for a column of resin in the backwashed, settled and drained condition is: 1. Add three bed volumes of 4%NaOH at a rate sufficient to allow 45 minutes of contact time. 2. Rinse with five bed volumes of potable water at the same flow rate. 3. Addthreebedvolumes of 10%H2S0,0r5%HC1ataflow rate that allows 45 minutes of contact time. 4. Rinse with five bed volumes of potable water. 5. Convert the resin to the desired ionic form by applying the regenerant that will be used in subsequent cycles. Ifthe column equipment has not been designed to handle acid solutions, a 0.5% CaCl, solution or tap water may be used in place of H,SO, or HCl for cation resins. Similarly, for anion resins, a 10% NaCl solution could be used in place of acids. 6.2 Batch Operations The batch contactor is essentially a single stage stirred reactor with a strainer or filter separating the resin from the reaction mass once the reaction 446 Fermentation and Biochemical Engineering Handbook is complete. This type of contactor has an advantage in some fermentation operations because of its ability to handle slurries. Additional advantages include the low capital cost and simplicity of operation. In a batch operation, an ion exchange resin in the desired ionic form is placed into a stirred reaction vessel containing the solution to be treated. The mixture is stirred until equilibrium is reached (about 0.5 to 3 hours). Then the resin is separated from the liquid phase by rinsing with the eluting solution. An additional step may be required to reconvert the resin to the regenerated form ifthis is not done by the eluting solvent. The cycle may then be repeated. The batch system is basically inefficient since the establishment of a single equilibrium will give incomplete removal of the solute in the feed solution. When the affinity ofthe resin for this solute is very high, it is possible that the removal is sufficiently complete in one stage. The batch process has the advantage over fixed bed processes that solutions containing suspended solids may be treated. In these cases, the resin particles, loaded with the adsorbate and rinsed from the suspended solids, may be placed in a column for recovery of the adsorbate and for regeneration. The batch contactor is limited to use with reactions that go to completion in a single stage or in relatively few stages. Difficulties may also arise with batch contactors if resin regeneration requires a greater number of equilibrium stages than the service portion of the cycle. 6.3 Column Operations Fixed Bed Columns. Column contactors allow multiple equilibrium stages to be obtained in a single unit. This contactor provides for reactions to be driven to the desired level ofcompletion in a single pass by adjusting the resin bed depth and the flow conditions. The main components of a column contactor are shown in Fig. 28. At the end of its useful work cycle, the resin is backwashed, regenerated and rinsed for subsequent repetition of the work cycle. Typically, this nonproductive portion of the cycle is a small fraction of the total operating cycle. Column contactors may be operated in cocurrent, countercurrent or fluidized bed modes of operation. The cocurrent mode means that the regenerant solution flows through the column in the same direction as the feed solution. The countercurrent mode has the regenerant flowing in the opposite direction as the feed solution. Countercurrent operation of a column may be preferred to reduce the ion leakage from a column. Ion leakage is defined as the amount of ion being Ion Exchange 447 removed from solution which appears in the column effluent duringthe course of the subsequent exhaustion phase. The leakage caused by re-exchange of non-regenerated ions during the working phase of cocurrently regenerated resin is substantially reduced with countercurrent regeneration. Figure 28. Ion exchange column contactor. The fixed bed column is essentially a simple pressure vessel. Each vessel requires a complexity of ancillary equipment. Each column in a cascade will require several automatic control valves and associated equip- ment involving process computer controls to sequence the proper flow of different influent streams to the resin bed. Combinations of column reactors may sometimes be necessary to carry out subsequent exchange processes, such as in the case of demineralization (Fig. 29). As this figure shows, a column of cation exchange resin in the hydrogen form is followed by a column of anion exchange resin in the hydroxide form. A mixed bed, such as shown in Fig. 30, may also be used for demineralization. Mixed bed operation has the advantage of producing a significantly higher quality effluent than the cocurrently regenerated beds of Fig. 29, but has the added difficulty of requiring separation ofthe two types of resin prior to regeneration. 448 Fermentation and Biochemical Engineering Handbook Na'CI' Solution- Cation Exctungor- Packed Cot. Anion Exchanger $+cr 0+cr @ +OH' H+OH- @+CI' mod col. &+OH' H+OH- @+OH' H+OH HOH Solution Figure 29. Demineralization ion exchange column scheme. Anion Anion Service Backwash Simultaneous Mixing Cycle Regeneration Figure 30. Operation of a mixed bed demineralization ion exchange column. An important requirement for the successful operation of a mixed bed is the carefbl separation of the strong base anion resin from the strong acid cation resin by backwash fluidization, This is followed by contact of each type of resin with its respective regenerant in a manner which minimizes the Ion Exchange 449 cross contamination of the resins with the alternate regenerant. This requires that the quantity of resin, particularly the cation resin, be precisely maintained so that the anioncation interface will always be at the effluent distributor level. Typically, matched pairs of resins are used so that an ideal separation can be repeatedly achieved during this process. Inert resins are marketed which enhance the distance between the anioncation interface and allow less cross contamination during regeneration.Ig61 The air mixing of the anion and cation must also be performed in such a manner that complete mixing of the resin and minimum air entrapment are obtained at the end of the regeneration cycle. Ion exchange is usually in a fixed bed process. However, a fixed bed process has the disadvantage that it is cyclic in operation, that at any one instant only a relatively small part of the resin in the bed is doing usefbl work and that it cannot process fluids with suspended solids. Continuous ion exchange processes and fluid bed systems have been designed to overcome these shortcomings. Continuous Column Operations. When the ionic load of the feed solution is such that the regeneratiodelution portion of the operating cycle is nearly as great or greater than the working portion, continuous contactors are recommended instead of column contactors. Continuous contactors operate as intermittently moving packed beds, as illustrated by the Higgins conta~tor[~~] (Fig. 3 l), or as fluidized staged (compartmented) columns, as shown in Fig. 32, by the Himsley In the Higgins contactor, the resin is moved hydraulically up through the contacting zone. The movement of resin is intermittent and opposite the direction of solution flow except for the brief period of resin advancement when both flows are cocurrent. This type of operation results in a close approach to steady-state operations within the contactor. Elegant slide valves are used to separate the adsorption, regeneration and resin backwash stages. The contactor operates in predetermined cycles and is an ideal process for feedstreams with no suspended solids. The Higgins type of contactor is able to handle a certain amount of slurry due to the continued introduction of fresh resin material to act as a filter media during the operation. A lower resin inventory should result with continuous contactors than with column reactors handling the same ionic load feedstream. contactor, [881[891[901 450 Fermentation and Biochemical Engineering Handbook tion 4 Regenerant 5 Section Figure 31. The Higgins contactor for continuous operation.[87] Figure 32. The Himsley Continuous Fluid-Solid contactor.[88] Ion Exchange 451 The major disadvantage of the Higgins contactor is the lifetime of the resin. Estimates range from as high as 30% resin inventory replacement per year due to attrition and breakage of the resin as it passes through the valves. Lgll In the Himsley type of contactor, the resin is moved from one compartment or stage to the next countercurrent to the feed solution flow on atimed basis that allows forthe rate ofequilibrium resin loading in each stage. Thus each compartment or tray is designed to accommodate specific feed compositions and effluent requirements. Equipment from different commer- cial suppliers differ in the manner of resin transfer. Fluidized Column Operations. Column contactors, when operated with the feedstream in a downflow mode, are poorly suited to handle fermentation slurries because of the excellent filtration characteristics of packed resin beds. For such slurries it is preferable to use a fluidized bed of resin such as is shown in Fig. 33.[921 While most of the commercial applications discussed in the literature pertain to slurries of uranium tailings and paper mill effluent, the equipment may be adapted for use in treating fermentation broths. The shape of the fluidized bed is important in controlling the position of the resin in the column. The effluent from the column should pass over a vibrating screen, e.g., as SWECO, to retain entrained resin, but allow mycellia to pass through. The Ashai contactor, shown in Fig. 34,Lg11 uses conventional pressure vessels as the resin column. These vessels have a resin support grid at the base and a resin screen at the top. The feedstream is fed in up-flow through the packed bed. Periodically the liquid contents of the column are allowed to drain rapidly which causes the resin to flow from the bottom of the bed to a similar vessel for regeneration. At the same time, fresh resin is added to the top ofthe active column from a resin feed hopper. This hopper contains a ball valve which passes the resin in during downflow operation and seals itself during the up-flow portion of the adsorption cycle. The Cloete-Street ion exchange equipment is amultistage fluidized bed containing perforated distributor plates (Fig. 35).fg3] The hole size in the plates is greater than the maximum resin particle. The countercurrent movement of resin occurs due to the controlled cycling of the feedstream. With each cycle, the entire amount of resin in one chamber is transferred to the next chamber. Equipment with 4.5 m diameter columns and eight stages for a total height of about 20 m are in commercial operation. 452 Fermentation and Biochemical Engineering Handbook The USBM equipment, shown in Fig. 36,[941 is very similar to the Cloete-Street system. The differences are in the plate design, the method of transferring solids and in the method of removing the resin. This system and the Cloete-Street system are able to handle slurry feeds with up to 15% by weight solids. The advantage of these fluidized bed columns is the relatively low capital cost, low operating cost, small space requirement, simple instrumen- tation and control compactability with conventional solvent extraction Only those systems that can accommodate slurries with suspended solids are commercially feasible for biotechnology and fermenta- tion operations. Otherwise, the small volume of fermentation feedstreams which need to be processed are not the scale of operations necessary to make continuous ion exchange processes cost effective. + RESIN SUPPLY .RESIN ADDITION PROTECTIVE UYER EFFLUENT OUTLET LOADED LAYER CONE REMOVAL DEVICE FOR RESIN FEEDSTREAM INLET f Figure 33. Fluid bed ion exchange Ion Exchange 453 h s: m 454 Fermentation and Biochemical Engineering Handbook vi B k 3 s G b tA al 0 b 0 Ion Exchange 455 2 0' soluloon I I I /o 0 0 0 o\ 1 00000 00000 f thick PV 8 1 Plot0 Oisiribution Distribution @loto Figure 36. Schematic diagram of the USBM ion exchange The principal disadvantage of fluidized bed columns is the mixing of resin in various stages of utilization. This mixing means that breakthrough occurs sooner and the degree of resin capacity utilization is much lower than in packed bedcolumns. By placing perforated plates in acolumn (Fig. 37),rg5] the resin beads only mix within a restricted area allowing more complete utilization of the ion exchange resin's capacity. 456 Fermentation and Biochemical Engineering Handbook 0 ION EXCHANGE BEAD FINE SOLID PARTICLE IN FEEDSTREAM Figure 37. The perforated plates in this fluidized bed column allow fine solid particles to flow through the column while the ion exchange beds are, for the most part, confined to individual compartments. Lgs1 Smaller continuous fluid bed systems, like the one shown in Fig. 38,1g61 have been developed which operate with a high concentration of ion exchange resin and suspended solids. These units are 80% smaller than the conven- tional resin-in-pulp plants of the type which are used in the treatment of uranium ore slurries.[97] The pilot plant unit, which would probably be the size needed for processing commercial fermentation broths, had dimensions percontactchamberof0.82 m x 0.82mwithafluidbedheightof0.82mand an additional 0.16 m for free board. The unit has been successfully operated with 25 to 50% resin and up to 45% suspended solids. Ion Exchange 457 Figure 38. Pilot plant for resin-in-pulp contactor The effect of the degree of regeneration of the resin on the degree of extraction ofa solute was measured by Slaterr9*I using a seven stage unit. The results are shown in Fig. 39 for the extraction of uranium using a fluidized bed slurry of a strong base resin in a 10% uranium ore leach slurry. However, the fluidized bed column, even with perforated plates separating it into as many as 25 compartments, may not be appropriate for applications in which there are strict requirements (<1% of influent concen- tration) on the effluent. For readily exchangeable ions, the optimum utilization of this technique occurs when an actual effluent concentration of 5% ofthe influent is the breakthrough point. At such times, the ion exchange resin capacity would be 70% utilized.[g5] Should it be acceptable that the breakthrough point occur when the average effluent reaches a concentration of 10% of the influent, the utilization of the ion exchange capacity is 90%. If the ions are not readily exchangeable (low selectivity), the resin utilization would be significantly less and fluidized bed operations should not be used. 458 Fermentation and Biochemical Engineering Handbook '0) - 99.9% ExtmaiIm QPM 0.020 0.015 0.00 0.005 I' 1 I234567 FEED STASE Figure 39. Effect of residual feedstream ions on the resin on the efficiency of a staged system. QPQ) is the concentration of the feedstream ions on the resin at the end of the cycle; QPW) is the concentration of the feedstream ions on the resin at the beginning of the cycle; q is the cycle time; x is the fraction of the resin bed volume in a stage which is removed per cycle.[98] 6.4 ElutiodRegeneration Elution of proteins from ion exchangers can be achieve with buffers containing salts such as sodium chloride or ammonium acetate or by an appropriate pH change, provided that the pH change does not result in denaturation of the eluted protein.[99] The elution may be performed with a series of stepped changes or with a continuous gradient change in the eluting power of the eluant, With such changes, it is possible to separate different proteins or protein fractions from each other based on their different affinities for the ion exchange resin. Ion Exchange 459 Elution of compounds such as penicillin with either acids or bases will render the penicillin inactive. Although aqueous salt solutions can elute penicillin without inactivating it, the large volumes required make this option impractical. Wolf and co-workers[loO] developed an elution solvent combi- nation of organic solvent, water and salt that can elute the penicillin with a minimum volume and no inactivation. The mixture of organic solvent and salt is chosen so that the salt is soluble in the resulting organic solvent-water mixture and the organic substance eluted from the resin is soluble in the elution mixture. Table 19 shows the elution volume required to recover the indicated amount of antibiotic when the elution solvent is 70% methanol and 5% or 7.5% ammonium chloride in water. Table 19. Amount of Antibiotic Recovered with Increasing Volumes of Methanol in Aqueous Ammonium Chloride[100] Antibiotic Eluant Total Volume Amount of of Eluant Antibiotic Recovered (Bed Volumes) (%) Dihydronovobiocin 70% MeOH 1 50.0 Novobiocin Penicillin with5%NH4Cl 2 3 4 70% MeOH 3.5 with 5% NH4CI 70% MeOH 0.5 with 7.5% m4cl 1.0 1.5 2.0 2.5 83.3 93.3 96.6 99.8 11.0 69.0 94.0 97.1 99.7 460 Fermentation and Biochemical Engineering Handbook Regeneration alone is not sufficient to prevent fouling or microbial growth on ion exchange resins. If the resin is left standing in the regenerant during nonoperating times, it is possible to suppress the microbial growth.[1o1] Theregenerant inthis instancewas 10or20%NaC1. Whentheinitialmicrobe count was 10 per milliliter, at the end of three weeks in 20% NaCI, the count had risen tojust 800/ml compared to 200,00O/ml forthe resin stored in water. When an alternate regenerant is used (NaOH or HCI), it is preferable to change the storage medium to 20% NaCl since extended time in an acid or base media can adversely affect the resin matrix. Since most fluids contain some suspended matter, it is necessary to backwash the resin in the fixed bed column on a regular basis to remove any accumulation of these substances. To carry out a backwashing operation, a flow of water is introduced at the base of the column. The flow is increased to a specific rate to classifl the resin hydraulically and remove the collection of suspended matter. Figures 40 and 4 1 show the types of flow rate which provide certain degrees of expansion of cation and anion resins, respectively. Since the anion resins are significantly less dense than the cation resin shown, it would be necessary to have different amounts offreeboard above the normal resin bed height so that backwashing may be accomplished with only a negligible loss of ion exchange resin. Typically, an anion resin bed may be expanded by 100% during backwashing, while a cation resin bed will only be expanded by 50%. It is also necessary that the water used for the backwashing be degassed prior to use. Otherwise, resin particles will attach themselves to gas bubbles and be carried out of the top of the column to give an unacceptable increase in resin losses. When treating fermentation broth filtrates, frequent backwashing of the resin bed is necessary to prevent accumulation of suspended matter. In such cases, the column height should be designed of such a size that the bed is regenerated at least every 10 hours. Shorter columns have been designed to be regenerated at least every hour.[lo21 These shorter beds canthen use finer resins and achieve a high level of efficiency with lower capital costs. This may be taken to the point of using very fine resins, as with the Powdex@ system[103] which discards the powdered resin after a single use. After a bed is backwashed, unless it is air-mixed as the level of water is drained down to the surface, the beads or particles classify according to size. The fine beads end up on top and the large beads on the bottom of the column. In cocurrent operations, the regenerant first contacts the top of the bed. The fast kinetics ofthe fine particles gives a high regeneration efficiency. However, the large beads on the bottom will regenerate more slowly and may Ion Exchange 461 end up only partially regenerated. Thus, when the feedstream is next passed through the resin bed, leakage of undesirable ions may occur from the large beads in the bottom of the column. This may be overcome by using a countercurrent flow arrangement described earlier or by using an air-mixing system during the post-backwash 0 16.32 32.64 48.96 BACKWASH FLOW RATE (HL/HIN /cn2) Figure 40. Backwash expansion characteristics of a macroporous strong acid cation resin, Dowex 88. 462 Fermentation and Biochemical Engineering Handbook TO DETERMINE FLOW RATE AT TERERATlRE T: / 0 I I FT FZ5(1 - 0.044(7 - 2.5)) 0 4,08 8.16 U,24 16.32 BACWASH FLOW RATE (ML/MlN/CH2) Figure 41. Backwash expansion characteristics ofregenerated and exhaustedmacroporous weak base anion resin, Dowex 66. 7.0 INDUSTRIAL CHROMATOGRAPHIC OPERATIONS Packed bed, process scale high performance liquid chromatography (HPLC) equipment was first introduced by Millipore and Elf Aquitaine in 1982.[1051 Since then, several companies, listed in Table 20, have enteredthe large scale HPLC market. The systems use packed beds at moderate pressures (30-140 bar). While there are substantial time-savings in using these systems compared to other purification techniques, the short life of the packing material and its high cost continue to restrict this technique to applications that warrant the $ 1OOkg separation cost. Ion Exchange 463 Table 20. Manufacturers of Process Scale HPLC Equipment Amicon Danvers, MA Dorr-Oliver Stamford, CT Elf Aquitaine (Varex in U.S.) Rockviile, MD Millipore Corporation Bedford, MA Pharmacia Piscataway, NJ Separations Technology Wakefield, RI YMC Morris Plains, NJ The Waters Kiloprep Chromatography pilot plant is one example ofthe successful extension of an analytical chromatography process to the process scale. The ability to control the various operation parameters to scaleup directly from the laboratory to the pilot plant and beyond to commercial production has been Figure 42 illustrates how the perfor- mance of this larger system can be predicted from the data generated in an equivalent laboratory apparatus. Voser and Walli~er['~I have described the approaches different compa- nies have devised sothat fine and soft adsorbents may be utilized in large scale chromatography operations. The scaling-up has usually been achieved by increasing the column diameter and using stack columns. The problem then becomes one of achieving uniform distribution of the feed solution over the entire resin bed surface, particularly when gradient elution is involved. In the Pharmacia approach, the fluid input is split and distributed through six ports on the column end plates.['07] At the entrance of each stream, an anti-jetting device spread the liquid over a fine mesh net. A coarse net between the net and the end plate acts as both a support and a spacer. For the soft Sephadex gels G-50 to G-200, the maximum feasible bed height is 15 cm. Scaleup operations have used as many as six such squat columns in series.["'] The drawbacks of this approach are the cumbersome adsorbent in the column. In AmiconNright columns, the flow is distributed through a carefully designed system of radial ribs cut out ofthe end plate. There is a single central port with a suitable anti-jetting device to reduce and divert the high velocity of the entering stream. The adsorbent bed is covered with a sintered plate. Sintered plates are claimed to be more efficient than most nets to achieve an 464 Fermentation and Biochemical Engineering Handbook (r 0 Ion Exchange 465 even distribution. The larger the pore size of the sintered plate, however, the less efficient the system. Drawbacks of sintered plates are their tendency to adsorb substances on their very large surfaces and the possibility of fouling. These columns may also be stacked. Whatman has developed a column with a new flow distribution system specifically designed for Whatman’s cellulosic ion exchange resins. The bed height is 18 cm. The diameter of the first commercial unit was 40 cm and contained 25 liters of adsorbent. The resin bed is covered with a perforated plate with ahigh free surface. The slightly conical head plate covers an empty space and has a steep cone in the middle. The total empty head space is about 5% of the bed volume. The feed solution enters the steep cone tangentially in its upper part. The resulting rotary movement efficiently mixes the supernatant liquid and allows gradient elution. It is claimed that filling and equilibration take only one hour. AMF has developed an unconventional new approach with its ZETA- PREP cartridge. The cartridges consist of concentric polymer screens which bear the ionic groups and are supported by cellulosic sheets. The flow is radial from the outer rim toward a perforated central pipe. The available nominal cartridge lengths are from 3 cm to 72 cm with a constant diameter of about 7 cm throughout. Scaling-up with this approach is quite straight- forward. Single cartridges, each mounted in a housing, can be combined to a multi-cartridge system. For such a system, flow rates up to 12 liter/min and bovine serum albumin capacities of 1400 g are claimed. The present ion exchange functionalities available are DEAE, QAE and SP. The step-wise transition from high pressure liquid chromatography to medium pressure chromatography, such as described for the preparation of pectic illustrate the progression toward large scale industrial application of the techniques developed in analytical laboratories. The pressure in these medium pressure chromatography applications is only 6 bar instead of the 100 to 150 bar associated with HPLC. The lower pressure results in longer processing times (about one hour) compared to the 5 to 20 minutes required for an analytical determination with HPLC. Studies, such as the one by Frolik and coworkers,[llO] which examine the effect and optimization of variables in HPLC of proteins, can be expected to contribute tothe implementation ofthis type ofprotein resolution technique into future commercial biotechnology processes. The first chromatographic systems capable of handling more than 100 kg/day were merely scaled up versions of laboratory chromatography.[111][112] Even with some of these systems it was necessary to recycle a portion of the overlap region to have an economical process. A typical example of such a 466 Fermentation and Biochemical Engineering Handbook system would be the Techni-Sweet System of Techni~hem["~I used for the separation of fructose from glucose. The unique distributors and recycle system are designed to maximize the ratio of sugar volume feed solution per unit volume of resin per cycle while at the same time minimizing the ratio of volume of water required per unit volume of resin per cycle. The flow through the Technichem system is 0.56 m3/(hr-m3) with a column height of 3.05 m. The feed solution contains 45% dissolved solids, and a feed volume equal to 22% of the volume of the resin is added to the column each cycle. The rinse water added per cycle is equal to 36% of the volume ofthe resin. This is much less rinse water than the 60% volume that was required by the earlier systems. This technique is known as the stationaryport technique since the feed solution and the desorbent solution are always added at the same port and the product streams and the recycle stream are always removed from another port. Technichem and Finn Sugar manufacture chromatography systems which utilize the stationary port technique. One of the earlier attempts['l4I at industrial chromatography used an adaptation ofthe Higgins contactor for the ion exclusion purification of sugar juices. The physical movement of the low cross-linked resin caused attrition as it was moved around the contactor. It was also difficult to maintain the precise control needed on flow rates because ofthe pressure drop changes and volume changes of the resin as it cycled from the mostly water zone to the mostly sugar solution zone. An alternate approach[' 151 utilizes movingport orpseudo-moving bed techniques. With these techniques, the positions on the column where the feed solution is added and where the product streams are removed are periodically moved to simulate the countercurrent movement ofthe adsorbent material. At any given time the resin column can be segmented into four zones (Fig. 43). Zone 1 is called the adsorption zone and is located between the point where the feed solution is added and the point where the fast or less strongly adsorbed component is removed. In this zone the slow or more strongly adsorbed component is completely adsorbed onto the ion exchange resin. The fast component may also be adsorbed, but to a much smaller extent. The second zone, Zone 2, is the purification zone and is located between the point where the fast component is removed and the point where the desorbent solution is added. Zone 3 is called the desorbent zone and is between the point where the desorbent is added and the point where the slow component is removed. In this zone the slow component is removed from the resin and exits the column. The final zone, Zone 4, is called the buffer zone and is located between the Ion Exchange 467 point where the slow component is removed and the point where the feed solution is added. There is a circulating pump which unites the different zones into a continuous cycle. Figure 43. Moving port chromatographic column with four zones for continuous chromato- graphic separation.[115] Different sections of the column serve as a specific zone during the cycle operation. Unlike the stationary port technique, the liquid flow is not uniform throughout the column. Because of the variations in the additions and withdrawals of the different fluid streams, the liquid flow rate in each of the zones will be different. With such a system, one must slowly develop the chromatographic distribution pattern through the different zones. It may take from 8 to 36 hours for the pattern to be established. Other practical considerations are that the recirculation system must represent a small (< 10%) portion of a single 468 Fermentation and Biochemical Engineering Handbook zone to prevent unacceptable back-mixing which would alter the established chromatographic pattern. The flow rate and the pressure drop per unit length of the chromato- graphic column are much lower for the stationary port compared to the moving port system. Also, the moving port system is much less capital intensive. The moving port technique, however, is calculated to require only one-third of the column volume and ion exchange volume and two-thirds of the desorbent volume compared to the stationary port technique. After the expiration of the UOP patent covering the rotary valve, there have been several modifications to the moving port technique by Amalgam- ated Sugar,[116] Illinois Water Treatment,["'] and Mitsubishi.[1181 Each manufacturer has its own proprietary approach for the establishment and control of the chromatographic pattern. proposed a hybrid system which has some of the charac- teristics of both elution chromatography and the pseudo-moving bed system. During the feed pulse, the feed position was moved continuously up into the column at a velocity that lies between the two solute velocities. The eluting solvent was continuously fed into the bottom of the column. Elution development with solvent was used when the feed pulse was over. This method reduces irreversible mixing of solutes near the feed point. Wankat and Ortiz[120] have used this system for gel permeation chromatography and claim improved resolution, narrower bands and higher feed throughputs compared to conventional systems. McGary and Wankat[121] have had similar results applying it to preparative HPLC. This technique uses less adsorbent and produces more concentrated products compared to normal preparative chromatography, but more adsorbent and less concentrated products than pseudo-moving bed systems. Wankat[122] has proposed that his system will be of most value for intermediate size applications or when only one product is desired. The key items to identifjl when considering an industrial chromato- graphic project are the capital for the equipment, yield and purity of the product, the amount of dilution of the product and waste stream, the degree of flexibility the computer controls allow, the expected life of the ion exchange material and whether the equipment allows for periodic expansion of the resin. New techniques are continuing to be developed which can be expected to be used in fiture specialized industrial applications. Multi-segmented columns have been demonstrated for the preparative purification of urokinase.[123] Begovich and ~oworkers['~~1['~~] have developed a technique for continuous spiral cylinder purifications which allow separation of the Ion Exchange 469 basis of electropotential in addition to the selective affinity of the adsorbent resin for the components in solution. A schematic ofthis device is shown in Fig. 44. Ll)r ma Figure 44. Schematic of the pressurized continuous annular chromat~graph.['*~I Another new technology that offers promise for commercial biotech- nology purifications is the use of parametric pumping with cyclic variations of pH and electric field. This has been described by Hollein and cowork- ers.[126] They worked with human hemoglobin and human serum albumin protein mixtures on a CM-Sepharose cation exchanger. The extensive equations they reported for parametric separations allow analysis of other systems of two or more proteins which may be candidates for this type of separation, 470 Fermentation and Biochemical Engineering Handbook Applications of ion exchange and column chromatography techniques have been incorporated into the commercial purification scheme for fermen- tation products, biomaterials and organic chemicals. While the majority of these applications are on the small scale (less than 500 kg/month but greater than 10 glmonth), several large industrial scale applications have arisen in the last decade. The extraction of sugar from molasses, the separation ofglucose from fructose, the separation of polyhydric alcohols, the separation of xylene isomers and the separation of amino acids are carried out in industrial scale operations preparing thousands of metric tons of purified material each year. Two recent books [1271[1281 provide extensive examples of these applications. Additional examples, mostly of laboratory studies, are available in books specifically on the HPLC of peptides and protein~.['~~1['~~] LC-GC and Chromatography are two periodicals with helpful operational suggestions. REFERENCES 1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. Thompson, H. S., Roy, J.,Agr. Soc., Eng., 11:68 (1850) Bersin, T., Natunvissenschaften, 33: 108 (1946) Winters, J. C. and Kunin, R., Ind. Eng. Chem., 41:460 (1949) Polis, B. D. and Meyerhoff, O., J., Biol. Chem., 169:389 (1947) Carson, J. F. andMaclay, W. D., J. Am. Chem. Soc., 67:1808 (1945) Nagai, S. and Murakami, K., J. SOC. Chem. Ind., Japan, 44:709, (1941) Wieland, T., Berichfe, 77539 (1944) Bergdoll, M. S. and Doty, D. M., Ind. Eng. Chem., Ancl. Ed., 18:600 ( 1946) Lejwa, A., Biochem. Z., 256:236 (1939) Cruz, E. C., Gonzales, F., and Hulsen, W., Science, 101:340:(1945) Jackson, W. G., Whitefield, G., DeVries, W., Nelson, H., andEvans, J., J. Am. Chem. SOC., 73:337 (1951) Moore, S. and Stein, W. H., J. Bid. Chem., 211:893 (1954) Puente, A., Microbiol. Espan., 14:209 (1961) Ishida, M., Sugita, Y., Hori, T., and Sato, K., U. S. Patent No. 3,565,95 1 (Feb. 23, 1971) Gordienko, S. V., J. Appl. Chem., 39: 10, USSR (1966) Shirato, S., Miyazaki, Y., and Suzuki, I.,FermentationIndustry, 45(1):60 Japan (1 967) Wheaton,R.M. andBauman, W. C.,Ann. N.Y. Acad. 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Phys., 28:418 (1958) Schogl, R., and Hemerich, F. G., J. Chem. Phys., 265 (1957) Glueckauf, E., and Coates, J. I., J. Chem. SOC. (London), p. 13 15 (1947) Hemerich, F. G., Liberti, L., Petruzzelli, D., and Passino, R., IsraelJ. of Chem., 26:3 (1985) Tsai, F. N., J. Phys. Chem., 86:2339, (1982) Vermeulen, T., Klein, G., and Hiester, N. K., in Perry’s Chemical Engineers’ Handbook, Sec. 16, (R. H. Perry and C. H. Chilton, eds.), McGraw-Hill, New York (1 973) Hemerich, F. G.,Ion Exchange, p. 255, McGraw-Hill, New York (1962) Samsonov, G. V. and Elkin, G. E., Ion Exchange and Soh. Extract., 9:211 (1985) Friedlander, S. K., A. I. Ch. E. J., 3:381 (1957) Marchello, J. M. and Davis, M. W., Jr., I & EC Fund., 2(1):27 (1963) Wilson, J. N., J. Am. Chem. Soc., 62:1583 (1940) DeVault, D., J. Am. Chem. Soc., 6.5532 (1943) Glueckauf, E., J. Chem. SOC., 69:1321 (1947) Glueckauf, E., Discussions Faraday Society, 7:42 (1949) Martin, A. J. P. and Synge, R. L. M., BiochemJ., 35:1358 (1941) Mayer, S. 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