Crystallization of Organic Compounds || Mixing and Crystallization

Chapter 6

Mixing and Crystallization

6.1 INTRODUCTION

Much of this chapter assumes operation in stirred vessels. Several alternative designs(fluidized bed and impinging jet crystallizers) are summarized early (Table 6-1) for compari-son with stirred tanks, and are described later in this chapter and in other parts of this book.An alternative feed addition geometry (mixing elbow) is described in Example 7-1.

Mixing in crystallization involves all elements of transport phenomena: momentumtransport, energy transport, and material transport in both the solution phase and the solidphase. In many cases, the interactions of these elements can affect every aspect of a crystal-lization operation including nucleation, growth, and maintenance of a crystal slurry. Tofurther complicate the problem, mixing optimization for one aspect of an operation mayrequire different parameters than for another aspect even though both requirements mustbe satisfied simultaneously. In addition, these operations are intrinsically scale dependent.

For these and other reasons to be discussed below, it might be stated that crystallizationis the most difficult of the common unit operations to scale up successfully.

Successful operations depend on identifying the mixing parameters for the most criticalaspects of the process and then evaluating whether those parameters will be satisfactory forthe others. Although this approach may be satisfactory in most cases, there will be crystal-lization procedures that require operation under conditions that are not optimum for mixingfor some aspects of the operation, as discussed below.

Note: Successful scale-up implies that both physical and chemical properties have beenduplicated between pilot plant and plant operations. These rigid criteria are not alwaysrequired but are, for example, for final bulk active pharmaceutical ingredients (APIs).In these cases, the width of the particle size distribution (PSD), the average particle size,the bulk density, and the surface area may all be required to fall within specified ranges.It is prudent to apply these criteria, if possible, in bench scale developmental planningand experimentation to reduce the risk of a dramatic failure. Examples of failures thatcould result from scale-up issues are (1) increased impurity levels, (2) small crystal sizecausing drastically reduced filtration rates, large PSD including a bimodal distribution,and (3) a poor-washing and slow-drying product.

For final bulk active pharmaceutical compounds, failures could also include physicaland chemical properties that do not meet regulatory requirements to meet biobatchspecifications.

Crystallization of Organic Compounds: An Industrial Perspective. By H.-H. Tung, E. L. Paul, M. Midler, andJ. A. McCauleyCopyright # 2009 John Wiley & Sons, Inc.

117

6.2 MIXING CONSIDERATIONS

The following is a brief discussion of mixing issues that can be expected to influence crystal-lization processes. Extended discussions of these and other mixing topics may be found inseveral references including Baldyga and Bourne (1999), Harnby et al. (1992), and Paul et al.(2003), as well as the crystallization texts of Mersmann (2001), Myerson (2002), Sohnel andGarside (1992), and Mullin (2001). A recent overview of mixing strategies for crystallizationis provided by Genck (2003).

As mentioned earlier, mixing requirements for crystallizers involve all aspects of trans-port properties: momentum, energy, and mass.

For momentum transport, i.e., velocity profile and impact on crystallization, theyinclude

† homogeneity of crystal slurry

† entrainment of gas/vapor from the head space (foaming)

† secondary nucleation through impact

† shear damage to crystals and impact on agglomerate formation/breakup

† satisfactory discharge of the slurry without excess retention of product crystals andpossible operation over a wide volume range

For energy transport, i.e., temperature profile and impact on crystallization, they include

† rate of heat transfer through the jacket wall

† avoidance of encrustation—solid scale on walls and baffles

Table 6-1 Mixing in Crystallizers for Pharmaceutical Processes

Function

Type of crystallizer

Stirred vessel Fluidized bed Impinging jet

Continuous and/or batch Both Continuous Semi-continuousType of mixer Variety of impellers Fluidization Kinetic energyCooling Good Excellent NAEvaporative Good NA NAAntisolvent Good NA ExcellentReactive ppt/cryst Good NA ExcellentCirculation-macromixing Poor to good� Excellent PoorMesomixing Poor to good� NA SatisfactoryMicromixing Poor to good� NA ExcellentMicromixing time tE, ms �5 to 40 NA 0.05 to 0.2Scale-up Can be difficult Good with good seed ExcellentSupersaturation range Wide Low HighControl of supersaturation Achievable Excellent at low S Excellent at high SSeeding Wide range Massive None or lowNucleation Wide range Minimum MaximumGrowth Wide range Maximum Minimum

�Dependent on localized conditions in the vessel

118 Chapter 6 Mixing and Crystallization

For mass transport, i.e., solution concentration profile and impact on crystallization,they include

† blending of solution and antisolvent components to the molecular level to achieveuniform supersaturation

† blending of reagents to the molecular level to achieve reactive crystallization/precipitation

For example, the initial blending of components to the molecular level, while avoidingregions of high supersaturation, requires consideration of the mesomixing and micromixingenvironments of the contactor. Other requirements involve the macromixing capabilities ofthe crystallizer. Extensive discussions of these fundamentals of mixing may be found in thereferences cited above.

Although many variations of mixing systems have been used for crystallization pro-cesses, the three primary types that are discussed in this book are the stirred vessel, fluidizedbed, and impinging jet devices. Each of these utilizes different mixing environments toachieve the desired local and global conditions.

The predominant system in the pharmaceutical industry is the stirred vessel. Fluidizedbeds (Chapters 7 and 11) and impinging jets (Chapter 9) fill specific mixing requirements, asindicated in Table 6-1.

6.3 MIXING EFFECTS ON NUCLEATION

6.3.1 Primary Nucleation

The effects of mixing on primary nucleation are exceedingly complex. The overall result is areduction in the width of the metastable region when this width for a static solution is com-pared to that for an agitated solution (see Chapter 2). Therefore, an unagitated solution can,in general, be cooled further before the onset of nucleation than an agitated solution. Since anindustrial system with few exceptions will always be agitated, this is of theoretical interestonly. (Exceptions are for operations such as melt and freeze crystallization in which suchissues are key factors; see Chapter 11 in this book; Mullin 2001, pp. 343ff.; Mersmann2001, pp. 663ff.)

In a mixed solution without crystals present and at constant supersaturation, increasedmixing intensity can reduce the induction time—the time elapsed after mixing to createsupersaturation to the time crystals first appear. Induction time decreases up to a criticalspeed, after which it remains unchanged (Myerson 2001, p. 145). Additional discussionmay be found in Chapter 4.

6.3.2 Secondary Nucleation

Since secondary nucleation is dominant as soon as nuclei appear, the nucleation mechanismsbecome virtually impossible to characterize in an industrial operation. In addition, anyseeded crystallization is by definition secondary even though some nuclei may simul-taneously form by a primary or other secondary mode mechanism. Therefore, the majorityof this discussion will focus on secondary nucleation.

6.3 Mixing Effects on Nucleation 119

The effect of agitation on secondary nucleation has been reported in the literature, andseveral references are discussed by Mullin (2001). Secondary nucleation is mixing depen-dent as follows:

† Crystal-crystal impact: a function of both the local micromixing environment and theoverall macromixing circulation

† Crystal-impeller and crystal-wall impact: functions of impeller speed, shape of theblade, and material of construction

These factors, along with the other intrinsic nucleation properties of the crystallizingsubstrate, affect the rate of nucleation, which in turn determines the number of nucleiformed and their size. This complex relationship is strongly dependent on the specificsystem characteristics but, in general, the nucleation rate increases rapidly with increasingenergy input. This high dependence is especially true for reactive crystallization with fastreactions. The reader is referred to Chapter 10 for a discussion of this topic, as well as toSohnel and Garside (1992).

6.3.3 Damkoehler Number for Nucleation

It is helpful to visualize the relationship between mixing and nucleation rates through an ana-logy with the reaction Damkoehler number (Da). The Da number for reaction is defined as

Da for reaction ¼ mixing time=reaction time

in which mixing time and reaction time can be the time required from the initial state (timezero) to 95% or other percentages of the final state. The mixing time is usually a measure oflocal mixing where the reaction is occurring. However, local mixing times vary over a widerange in a stirred vessel in which the reactants are subject to varying local mixing times.

As shown in Fig. 6-1, at low values of the Da number, the reaction yield is insensitiveto mixing, whereas at high values of the Da number, the reaction yield is sensitive tomixing. The reader can find a more detailed discussion on this concept in chapter 13 ofPaul et al. (2003).

Increasing k2

Y/Yexp

DaM = kR1CBotM

1.0

Figure 6-1 Reaction yield as a function of the reaction Da number; mixing sensitivity for chemical reactionsis analogous to mixing sensitivity in crystallization nucleation and growth, from low sensitivity at low Da tohigh sensitivity at high Da.


By analogy, a Da number for nucleation could be visualized as a ratio between mixingtime and induction time:

Da for nucleation ¼ mixing time=induction time

Mixing time and induction time can be the time required from the initial state to 95% or otherpercentages of the final state.

As shown in Fig. 6-2, at low values of this ratio—fast mixing and long induction time—mixing would have a minimal effect on nuclei size, whereas at high ratios—slow mixing andshort induction time—mixing effects would be critical to nuclei size. Local supersaturationcould be visualized as a parameter as shown.

For slow-nucleating systems that have long induction times or equivalently wide meta-stable zone widths, the location of the feed point and the impeller energy input may not be asimportant as for higher values of the ratio in determining the size—and number—of nucleithat can be generated in regions of the feed stream. This analogy also parallels that used inreacting systems in evaluating the requirements for micromixing of reagents.

For reactive crystallization processes, in addition to mixing time and induction time, wemay need to consider the reaction time in the analysis. A fast reaction (high Da for reaction)could be readily seen as sensitive to the local mixing environment of two reactive streams.It could generate locally a high product supersaturation due to the fast reaction. This, in turn,could generate small particles as in precipitation which can have very short induction time.However, if the product has a long induction time, i.e., low Da for nucleation, the local highsupersaturation regions can be mixed with the rest of the system and distributed throughoutthe vessel before nucleation. The PSD would not be sensitive to micromixing or mesomixingbut may still be sensitive to macromixing and other mixing issues, such as impact with theimpeller and other crystals.

If the reaction rate is slow compared to the mixing rate, the reactants can be dispersedthroughout the vessel before they react. The resulting product solution concentration will bemore uniform throughout the vessel. The operation could be thought of as being determinedby crystallization parameters rather than the reaction rate.

For pharmaceuticals, most of the reactive crystallization processes are salt formationfrom an acid and base. In this situation, the reaction rate is generally much faster than themixing or crystallization rate, and the mixing sensitivity depends primarily upon the magni-tude of the induction time, or Da for nucleation.

Increasing S

dp/d

p max

DaC = tM/tind = mixing time/induction time

Figure 6-2 Particle size as a function of a crystallization Da number (the ratio of mixing time/nucleationinduction time).

6.3 Mixing Effects on Nucleation 121

It should always be kept in mind that mixing and nucleation time can be arbitrarilydefined as the time required from the initial state to, for example, 95% of the final state.These values are also affected by other operating factors, such as the type of solvent, temp-erature, degree of supersaturation, flow ratios of two streams, and vessel geometry.

It cannot be overemphasized that the Da number also depends upon the “scale.” On theone hand, we can see the mixing time as primarily determined by the scale of the mixingdevice. Mixing time in micro-scale devices, which can be considered equivalent to thetime for micro-scale mixing, is always faster than the mixing time in large-scale vessels.On the other hand, induction time is not considered to be a function of scale. Thus, a nuclea-tion process may perform successfully in the laboratory using a small vessel (mixing timeshorter than induction time) but fail to produce the expected mean particle size and PSDin the pilot plant or factory vessels (mixing time longer than induction time) because ofthe local non-uniformity of concentration and its effect on nucleation induction time.

Clearly, the key to a successful scale-up of this process is to maintain the mixing timeat the pilot plant and factory scale below the threshold which would cause the process tomove down the curve of Fig. 6-2. This can be best accomplished by using a specialmixing device, such as an impinging jet, to approach the same mixing time at allscales. The reader can find more information on impinging jets and crystallization inSection 6.6.3 and Examples 9-5 and 9-6.

6.3.4 Scale-up of Nucleation-Based Processes

Nucleation events can dominate the entire crystallization operation with respect to bothphysical and chemical purity attributes. Ultimate crystal size as a function of the numberof nuclei generated is summarized in Table 4-2, where the nominal dimensions of resultingcrystals (spherical, flat plates, needle-shaped) are shown as a function of the number of nucleior seed particles added. It can be seen that the number of nuclei generated by the variouscauses of nucleation—including agitation—has a negative exponential effect, as expectedfrom this purely geometrical relationship, on the ultimate size that can be achieved bygrowth subsequent to nucleation. The nucleation rate, particularly contact nucleation, canincrease on scale-up because all key parameters of (stirred tank) mixing cannot be held con-stant. For example, scaling up at equal power per unit volume results in an increase in impel-ler tip speed. The resulting average particle size could then be reduced because, after theincreased nucleation, there are more particles for a reduced amount of substrate to growon. In addition, other mixing factors that affect growth could increase the size distributionfurther, as discussed in Section 6.4. Oiling out and agglomeration also can further compli-cate a nucleation-based process, as discussed in Section 5.4.

The critical nature of these interactions is the key factor in causing difficulty in scale-upof nucleation-based crystallization processes—even with small quantities of seed.

The critical mixing factors in a stirred tank are impeller speed and type, as well as theirinfluence on local turbulence and overall circulation. Since all aspects of these factors cannotbe maintained constant on scale-up either locally or globally, the extent to which changes inthe crystallizing environment will affect nucleation is extremely difficult to predict. To themixing issue must be added the uncertainties caused by soluble and insoluble impuritiesthat may be present in sufficiently different concentrations from batch to batch to cause vari-ation in induction time, nucleation rate, and particle size.

The severe problems associated with nucleation-based operations, some of which,especially in stirred vessels, are directly caused by mixing issues, lead to the conclusion


that dependence on nucleation can be problematic in achieving reproducible results onscale-up and/or in ongoing production.

If no process alternative is possible to avoid dependence on nucleation, mixing scale-upcan be based on equal power per unit volume, assuming that the same impeller type is used.In most cases, however, this approach will result in changes in PSD on scale-up that may ormay not be acceptable. In general, as suggested by Nyvlt (1971), the PSD will be broader andthe average particle size will be smaller if this scale-up criterion is used. As often experi-enced in crystallization scale-up, however, the opposite can occur, depending on the specificnucleation characteristics of the system. A further generalization is that rapid nucleationtends to produce the smaller size distribution on scale-up, whereas slow nucleation cangive the opposite result.

A more robust way to accomplish consistent scale-up of a nucleation-based process is toapply impinging jet crystallization. The energy of impinging jets can achieve rapid mixing oftwo streams at high local energy dissipation rates to the molecular level, possibly beforenucleation occurs. By matching the energy dissipation rate at different scales, the desiredmixing time from laboratory-scale devices to production scale can be achieved. This tech-nology is discussed in Chapters 9 and 10.

Several references on nucleation provide excellent insight into this complex phenom-enon, including Mersmann (2001), Mullin (2001), and Myerson (2001).

6.4 MIXING EFFECTS ON CRYSTAL GROWTH

6.4.1 Mass Transfer Rate

Mixing can obviously have a large effect on the mass transfer rate of growing crystals throughits effect on the film thickness. This influence is dependent on both the size of the crystals andthe mixing intensity. As mixing intensity increases, mass transfer rate increases and filmthickness decreases up to a limit, beyond which the effects approach limiting values.

The concentration gradient in the film is illustrated in Fig. 6-3, where it can be seen thatsupersaturation conditions are present throughout the film. When crystallization is diffusionlimited, concentration in solution drops significantly from that in the bulk (Cb) to that in thefilm (Cf), which is close to that at the crystal growth surface (Csurf). When growth rate is pri-marily limited by resistances to surface incorporation, the larger drop in concentration is

Bulk

Csurf

Cb

Cf

Cb

Cf

Csurf

Diffusion control Surface incorporation control

Figure 6-3 Schematic representation of concentration gradients from bulk solution to growing surface.

6.4 Mixing Effects on Crystal Growth 123

across the film. The thickness of the film will determine the time available for nucleationwithin it. The new nuclei are either shed into the bulk or attach as shards to growing surfaces.In both cases, the effect will be to limit growth.

Large crystals (.100 microns) may be more subject to film mixing issues than smallerones because crystals less than 10–20 microns are approximating the Kolmogoroff eddy sizeand tend to follow these eddies. These differences in crystal size can change the growth ratelimitation from film control in large crystals to intrinsic growth rate control in small ones.

6.4.2 Da Number for Crystallization

The Da number concept has been applied to crystallization by Garside (1971) and Garsideand Tavare (1985) by the development of an effectiveness factor for growth by comparingthe growth rate at the interface conditions to the growth rate expected if the interface wereexposed to the bulk solution conditions. The Da number in their work is defined as theratio of the surface integration rate to the mass transfer rate through the film. This definitionis different from the Da number defined in this book, which is the ratio of mixing time tocrystal growth or nucleation time. The effectiveness factor varies in an S-curve from 1.0at low Da to zero at high Da. A value of 1.0 would indicate pure surface integrationgrowth where mixing effects are unimportant, while a low value approaching 0 would indi-cate pure diffusion growth with high sensitivity to mixing. Regions between these limitsrepresent a combination where both growth and mixing have an effect.

The Da number for crystal growth can be defined as

Da for crystallization ¼ mixing time=overall crystal growth time

or equivalently

Da ¼ mixing time=supersaturation release time

Again, mixing time and supersaturation release time can be those required to achieve 95% orpercentages of the final state.

Similar to the discussion of the Da number for nucleation, at low Da values forcrystallization—fast mixing and slow release of supersaturation—mixing would not affectcrystal growth, whereas at high ratios—slow mixing and fast release of supersaturation—mixing effects would be critical to PSD, as fast local crystal growth may occur.

For systems with a slow release of supersaturation, the location of the feed point and theimpeller energy input may not be as important as for those with higher values of the ratio indetermining the size of crystals that can grow in regions of the feed stream.

Depending upon the nature of the crystals and the operating environment, authors haveobserved a wide range of release rates of supersaturation. The time scale for the release ofsupersaturation can vary from fractions of seconds to days (or longer). Due to the wide vari-ation in release rate of supersaturation, the inherent concepts of the Da number for crystal-lization provide a useful way to readily identify the potential mixing sensitivity issue of aparticular crystallization process.

6.4.3 Conflicting Mixing Effects

Factors that improve with increased mixing in a stirred tank are (1) heat transfer, (2) bulkturnover, (3) dispersion of an additive such as an antisolvent or a reagent, (4) uniformity


of crystal suspension, (5) avoidance of settling and minimization of wall scale, and (6) mini-mization of impurity concentration at the crystallizing surface.

However, these factors must be balanced against the possibly negative results ofovermixing that can result in crystal breakage and/or shedding of nuclei as well as increasedsecondary nucleation.

These concerns lead to the conclusion, referred to above, that it is often necessary to choosea mixing condition (impeller speed, type, etc.) that may not be optimum for every aspect of thecrystallization and may actually not be optimum for any of them. In many cases, however, oneend result (i.e., PSD, bulk density, uniformity of suspension, and approach to equilibrium solu-bility [yield]) may dictate the choice of mixing conditions. In this case, it becomes essential todetermine if the negatively affected aspects can be tolerated. If these problems are occurringin operation in a stirred vessel, a different type of crystallizer, such as a fluidized bed, mightbe used to promote crystal growth and minimize nucleation. Readers can find more informationon fluidized bed crystallizers in Section 6.6.2 and Examples 7-6 and 11-6.

6.4.4 Experimentation on Mixing Effects

All of these factors are properties of a given crystallization system, thereby requiring choicesfor each specific operation. Experimentation is required to determine the key responses tomixing for each system and could include determination of the following:

† Effect of impeller speed and type on PSD at a minimum of two seed levels and twosupersaturation ratios. These results should indicate the sensitivity of the system tomixing.

Note: A small response could indicate that other system properties were controlling (i.e.,inherent crystal growth rate or nucleation rate). A large response would indicate sensitivity tosecondary nucleation and/or crystal cleavage and require additional experimentation andevaluation of scale-up requirements. The laboratory results should be evaluated relative toeach other since scale-up can be expected to make additional changes in PSD, especiallywhen a large response is experienced in these simple experiments.

† Effect of impeller speed on crystallization rate and approach to equilibrium solubility(yield). Failure to achieve equilibrium solubility may indicate accumulation of impu-rities at the crystallizing surfaces. An increase in impeller speed resulting in furtherreduction in solution concentration could indicate resumption of growth or additionalnucleation (see Example 6-1).

† Suspension requirements, as indicated by the settling rate to achieve off-bottomsuspension (see Section 6.6 below).

† Effect of feed pipe location (for antisolvent and reactive crystallizations) (see Section6.6.1.5 below).

For nucleation-dependent operations, it is recommended that additional information beobtained as follows:

† Effect of impeller speed and type on the width of the metastable region.

† Effect of impeller speed and type on the rate of nucleation.

Note: This experimentation is focused primarily on evaluation of mixing sensitivity.Other experimentation on crystallization issues is beyond the scope of this discussion.

6.4 Mixing Effects on Crystal Growth 125

6.4.5 Effects of Mixing on PSD

The effect of mixing on PSD has been experimentally examined for the reactive crystalliza-tion of calcium oxalate (Marcant and David 1991). This work is an excellent example of themultiple dependencies on mixing that can be experienced in a crystallization operation. Thefactors noted above that are mixing dependent are shown to have positive or negative influ-ences on the resulting physical characteristics, thereby illustrating the necessity of selectingthe most important result to be achieved. Increasing the agitator speed is shown to initiallycause an increase in particle size, followed by passing through a maximum and then decreas-ing particle size. This result is attributed to changes in controlling factors resulting from thechanges in mixing. This experimental result provides an excellent justification for bothvariable-speed drive and subsurface feed.

Experimental results indicating the sensitivity of particle size and PSD to thelocation of the feed stream may be used to confirm the sensitivity of a compound tomixing. If no sensitivity is observed, it may be concluded that intrinsic factors suchslow growth or low nucleation rate are dominant. For additional discussion, seeSection 6.6.1.5 below.

6.5 MIXING SCALE-UP

6.5.1 Power

The compromises in mixing optimization that may be required on scale-up often result in theuse of the common mixing criterion of equal power per unit volume or, in some cases, equaltip speed. Both of these recommendations are more relevant for utilization of the sameimpeller type as well as geometric similarity. Laboratory evaluation of the mixing require-ments for a crystallization operation should be carried out in a minimum 0.004 m3 liter vessel(4 liters) for preliminary data and a further evaluation at 0.1 to 1 m3 as practical.

Smaller-scale operations will generally produce a more uniform PSD and a larger meancrystal size than the manufacturing scale (typically �10 m3) when using equal power perunit volume. These changes typically are caused by the local differences in impeller shear(an unavoidable result of the equal power per unit volume criterion) that cause increasednucleation leading to a larger number of particles, an increased spread in PSD, and a smallerparticle average diameter.

Guidelines other than equal power per unit volume were suggested by Nienow (1976)that can be helpful in avoiding this local over-mixing. Using this guideline, the agitator speedat the manufacturing scale would be selected to be sufficient to just maintain off-bottom sus-pension, thereby resulting, in addition to reduced shear damage, in reduced nucleation, fewerparticles, and more growth. In general, this speed would be considerably less than equalpower per unit volume, depending primarily on the density difference between the suspend-ing solvent and the crystals and their size.

Limitations on this guideline would be high-density crystals that require possibly dama-ging higher speeds to achieve the just-suspended condition. In addition, antisolvent and reac-tive crystallization applications may require higher speeds to prevent local supersaturation atthe point of addition. In the latter case, scale-up based on equal local energy dissipation at thepoint of addition may be necessary.

A further caution on reduced speed is a possible increase in encrustation caused by crys-tal contact with the bottom surface with insufficient fluid force to prevent wall growth.


6.5.2 Off-Bottom Suspension

The requirements for particle off-bottom suspension are also discussed in many literaturereferences including Zwieterling (1958), Chowdhury et al. (1995), and Paul et al. (2003,Chapter 10). Methods of calculating various degrees of homogeneity in solids are presented.These calculations are an important part of crystallization scale-up studies.

6.6 CRYSTALLIZATION EQUIPMENT

The three types of crystallizers summarized in Table 6-1 achieve operation in extreme mixingenvironments, thereby providing a wide variety of contacting capabilities. Equipmentchoice can be tailored to specific needs of supersaturation control—high or low—andmixing intensity—high or low—as well as other aspects of the operation. However, asdiscussed in the sections that follow, it may be difficult to find one mixing system for a stirredtank that can satisfy all the needs of a specific crystallization operation. Fluidized beds andimpinging jets can be designed and operated in narrower regions of mixing intensity, lowand high, respectively, and may therefore be tailored to meet all the needs for a particularoperation when matched to the specific needs of that system.

6.6.1 Stirred Vessels

6.6.1.1 Alloys and Stainless Steels

A stirred vessel crystallizer is shown in Fig. 6-4. Included are a dual-impeller pitched-bladeturbine with a “tickler” blade (see Section 6.6.1.6), a subsurface addition line, baffles, and aram-type bottom outlet valve to aid in discharge of slurries.

The workhorse impeller is the pitched-blade turbine because of its ability to create goodcirculation at relatively low shear. These attributes help reduce secondary nucleation andcrystal breakage while achieving good suspension and circulation. The flat-blade turbineis less versatile because of high shear and less overall circulation. The Ekato Intermig hasproven to have superior performance in some crystallization operations because of itscombination of excellent circulation and low shear.

Baffles are required in all cases to prevent poor mixing due to swirling as well as entrain-ment of vapor which can provide nucleation sites. Baffles may also have an important effectin minimizing foaming. Propellers and hydrofoils are not normally suitable for multipurposeservice in heterogeneous systems—especially in the event that nucleation and growth maypass though a stage involving sticky solids, as discussed above in Section 5.4.1.

Computational fluid dynamics representations of flow in these vessels may be helpful invisualizing flow patterns and particle paths.

6.6.1.2 Glass-Lined Vessels

The versatility of the glass-lined vessel in a large variety of chemical environments has madeit the most common in the industry. For these reasons, a crystallization step may be carriedout in an equipment train in a glass-lined vessel whether or not this is required to preventcorrosion. These reactors range in size from 80 to 20,000 liters or more. One limitation inthe use of glass-lined vessels related to mixing and heat transfer is that the limit of the temp-erature difference between jacket and batch is �1258C. However, such temperature extremes

6.6 Crystallization Equipment 127

are not often encountered in crystallization. (The manufacturer should be consulted forspecific limitations on the type of glass lining and base metal in use.)

The retreat-blade and anchor impellers that have been widely used for many yearsare now being replaced by glass-lined turbines and other shapes that have been recentlydeveloped by the manufacturers using sophisticated methods of applying the glass tomore sharply angled shapes. These turbines are available for vessels as small as 80liters, although shafts with removable, interchangeable impellers are not currently avail-able for tank sizes smaller than about 1200 liters. These limitations are subject toadditional improvement, and it is recommended that the specific vendors be consultedfor updated information.

These new impellers, especially in multitier configurations, have greatly improvedthe mixing capabilities of the glass-lined reactor by providing increased shear and cir-culation. A number of glassed impeller types are now available including curved- andpitched-blade turbines. Some examples of these are shown in Fig. 6-5. In addition,two or three turbines mounted on a single shaft are now available for larger vessels(.3 m3). The lower turbine can be positioned within �10 cm of the vessel bottom.For single turbines in larger vessels, however, the low turbine position may not providethe desired overall circulation.

The three manufacturers of glass-lined mixing systems use different methods of attach-ment of the blades of the impeller(s) to the shaft as summarized in Table 6-2.

Subsurfacefeed

Splitjacket

Above-surfacefeed

Baffle

Multi-agitators

Flush bottom valve

Figure 6-4 Typical stirred vessel crystallizer.


6.6.1.3 Baffles

Although glassed baffle design has significantly improved mixing performance, most glass-lined vessel applications utilize only a single baffle in order to maximize the number of tanknozzles available for other purposes. Two manufacturers, DeDietrich and Tycon, offer modi-fications to include multiple baffles in their mixing systems.

6.6.1.4 Variable-Speed Drive

The use of variable-speed drives in pilot plant and manufacturing plant vessels is recommendedfor development and scale-up of crystallization processes. This capability provides the oppor-tunity for critical experimentation at the pilot plant scale to determine the effect of impellerspeed on PSD and other variables. On the manufacturing scale, the ability to change impellerspeed is the most readily adjustable parameter for manipulation on scale-up. Modern variable-frequency drives provide an excellent means to vary speed over a wide range. The added costof variable-speed capability is minimal compared to all other methods of changing mixing

Figure 6-5 (a) GlasLockw glass-lined steel impellers.Courtesy of De Dietrich Process Systems.


parameters and may be one of the most effective ways of solving scale-up problems andproviding versatility for multipurpose operations. In addition, speed changes can be utilizedin troubleshooting in plant operations that are not meeting expected results.

6.6.1.5 Subsurface Addition Lines

Another key factor in successful pilot plant and manufacturing plant operation is the appro-priate use of subsurface addition of antisolvents and reactive reagents for crystallization andprecipitation. The primary purpose of subsurface addition is the introduction of the agentcausing supersaturation in the region of intense micromixing in the vessel to avoid localexcess supersaturation with the associated potential for increased nucleation.

The effect of the feed point has been shown to be dramatic in the reactive crystallizationof calcium oxalate, as reported by Marcant and David (1991). The conclusion was reachedthat determination of the sensitivity of crystallization to feed point location can provide infor-mation on controlling factors. In a laboratory development program, therefore, an effectivemethod of determining the importance of mixing effects is to run experiments with two ormore feed points—a surface feed and an impeller feed—and at two different impeller speeds.If no difference can be found in the PSD and mean particle size, mixing effects may not becontrolling and other aspects such as inherent growth rate may be dominant. As discussed in

Figure 6-5 (b) Cryo-Lockw impellers.Courtesy of Pfaudler, Inc.


Sections 6.3.1.4 and 6.4 on the use of the Damkoehler number concept, the need for effectivelocation of a feed line may be evaluated from a comparison of nucleation and mixing rates.

The mixing texts referenced above contain extensive discussion of the importance of thelocation of feed streams. While these studies are primarily concerned with reagent feed forchemical reactions and the influence of local turbulence on reaction selectivity, the sameissues are encountered in the addition of antisolvents and reagents for reactive crystallizationbecause nucleation is a function of supersaturation, whether local or global.

The location of the point of introduction and the diameter and flow rate of the additionstream are key to successful operation and scale-up. Generally, the optimum location is at the

Figure 6-5 (c) ElcoLockw and fixed impellers.w impellers.Courtesy of Tycon Technologies, a Robbins & Myers Company.


point of maximum turbulence, as characterized by the shortest mixing time constant. Thispoint is in or near the impeller discharge flow. For the down-pumping pitched-blade impel-ler, this point would be just above the impeller so that the flow would be rapidly mixed bypassing through the impeller blades. For a radial flow flat-blade turbine, it would be directedinto the discharge flow.

Feed introduction on or near the surface of a stirred vessel can have dramatic effects on acrystallization process because unwanted oiling out, nucleation, and/or agglomeration canoccur in these poorly mixed zones. It is recognized that mechanical design and cleaningissues may make the use of subsurface lines more difficult. However, it cannot be overem-phasized that the negative effects on a manufacturing-scale operation can be far more costlythan provision of the necessary mechanical requirements. For alloy vessels, the mechanicalissues are minimal. For glass-lined vessels, the use of Teflon tubing attached to the baffle hasbeen effective.

Cleaning issues can be resolved by providing for removal of the subsurface line betweenbatches or during turnaround.

The diameter of the subsurface line is also a key factor with crystallization, as it is withreactions. If the diameter is too large, the incoming stream may produce regions of highsupersaturation before it can be effectively blended to the molecular level. Large-diameterpipes can also be subject to reverse flow in which the crystallizing mixture is forced intothe pipe, where it is subject to meeting the incoming antisolvent in a nearly stagnantregion, leading to nucleation and possible solids accumulation and plugging.

Extensive analysis and experimentation on dip-pipe design has been carried out byPenney and co-workers (Jo et al. 1994), and the reader is referred to this work for specificdesign recommendations for diameter and flow rate. Their recommendations are summar-ized in Table 6-3 for six-blade flat-blade and three-blade high-efficiency down-pumpingimpellers. These recommendations should be followed to prevent dip-pipe back-mixing,enhance local blending, and prevent solids plugging.

In the Table 6-3, D and T are the impeller and tank diameter, respectively, G is the dis-tance from the feed pipe to the impeller, and Vf/Vt,min is the recommended minimum ratioof the velocity in the feed pipe to the velocity in the tank at the feed pipe location.

The importance of the feed location for chemical reactions has been clearly establishedby the work referenced above and many others. The literature contains less data on crystal-lization. However, undesired nucleation is potentially present for all crystallization systems,depending on the nucleation rate and the degree of local supersaturation. The analogybetween reaction sensitivity and supersaturation sensitivity can be visualized through theconcepts represented by the Damkoehler number, as discussed in Section 6.3.1.4 above.

Table 6-2 Glass-Lined Impellers and Their Methods of Attachment

Blade attachment Removable Variable pitch

DeDietrich GlasLockw

Friction fit of blades into holesin the hub

Yes Yes(Within power limit of motor anddrive)

Pfaudler Cryo-Lockw

Liquid N2 cooling to shrinkthe shaft

Yes No(Different pitches available)

Tycon Elcolockw Yes No(Different pitches available)


As in scale-up of chemical reactions, it is important to consider the relative time constants ofmixing and nucleation in order to estimate which part of the S-curve applies to a particularsystem. Although this determination is conceptual, the insight provided by such analysis canbe helpful in establishing operational design criteria.

6.6.1.6 Discharge of Slurry

Efficient discharge of the crystal slurry to the next process vessel—usually a filter—isrequired to prevent yield loss by settling or to achieve the necessary degree of crystalslurry homogeneity for satisfactory filter loading or other subsequent operations. Themixing obviously changes as the slurry is discharged and can result in an unsatisfactorydegree of settling after passing the impeller or an unsatisfactory degree of homogeneityfor the subsequent operation. “Tickler” impellers are often used, as shown in Fig. 6-4, toprovide mixing until the discharge is nearly complete. These may be custom designedand fitted by the manufacturer. The diameter of the tickler blade is usually one-half totwo-thirds the diameter of the main impeller, and thereby requires little power and doesnot create sufficient shear to be harmful.

Selection of pumps for discharge may also be a significant issue with regard to energyinput and slurry pumping capability. Crystals that have been carefully grown and are readyfor high rate filtration can be reduced in size by pump shear to drastically reduce the filtrationrate. This is especially problematic for recycle pumps in filter equipment feed loops.

Low-shear pumps can be effective in limiting shear damage and include lobe and dia-phragm pumps. The reader is referred to an analysis of pump and transfer energy byMersmann (2001, pp. 454ff.) to aid in design of slurry transfer and control of flow.

Shear damage on transfer can be eliminated by gravity transfer when equipment layoutincludes the necessary vertical clearances.

6.6.2 Fluidized Bed Crystallizer

An important alternative type of crystallizer is the fluidized bed. One fluidized bed designis shown in Fig. 6-6. As shown, the supersaturated solution enters at the bottom of the crys-tallizer and the clear, partly or fully depleted solution exits at the top of the crystallizer. Theslurry is suspended in the crystallizer by the upward liquid flow, which usually lies in the

Table 6-3 Recommended Minimum Vf/Vt for Selected Geometries for Turbulent FeedPipe Low Conditions

Case Impeller Feed position D/T G/D Vf/Vt,min

1 6BD� Radial/mid-plane‡ 0.53 0.1 1.92 6BD� Above/near shaft§ 0.53 0.55 0.253 HE-3† Radial/mid-plane† 0.53 0.1 0.14 HE-3† Above/near-shaft§ 0.53 0.55 0.15

�Six-blade disk turbine.†High-efficiency three-blade down-pumping turbine.‡Injection radially inward toward the impeller at its mid-plane at a distance G.§Injection downward into the impeller at about D/4 from the centerline of the impeller shaft and G/D above theimpeller mid-plane.


laminar and transition flow domain. When designed with a tapered bottom section of modestincluded angle to minimize flow separation, the hydrodynamic flow pattern is much closer tothe plug flow within a tube than that in a continuously stirred tank. Therefore, the crystallizercan exhibit a high degree of homogeneity with a very low degree of back-mixing at the sameheight of the bed, but the slurry concentration and particle size distribution can vary alongthe height of the bed. This is a sharp contrast to the “homogeneous” mixing of slurry withina stirred tank.

Due to its low turbulence intensity, fluidized bed crystallization is very effective in mini-mizing nucleation by providing very low shear, low energy, and minimum velocity impactbetween crystals. The operation can achieve growth by operating at low supersaturation andin the presence of a large surface area for growth.

This technology is used primarily in cooling crystallization instead of antisolvent orreactive crystallization. Thus, residence time, instead of mixing time of two streams,would be more appropriate for the definition of the Da number, as well as for scale-updesign. A major advantage of continuous and semicontinuous operation that can be realizedwith fluidized beds is that the seed is always present, thereby eliminating the need for timingof seed introduction and other aspects that have been discussed as variables in batch andsemibatch operation. Initial seed preparation becomes critical to success; methods for grow-ing initial seed are presented in Chapter 5. Once grown, however, the ongoing growth on theseed continually renews the supply. Methods of crystal cleavage to maintain seed PSD arealso presented in Chapter 5.

Both continuous and semicontinuous operation can be utilized. This principle has beensuccessfully applied in the resolution of optical isomers in which nucleation must be mini-mized, and preferably eliminated, to achieve isomer separation, as described by Midler(1970, 1975, 1976) and presented in Examples 7-6 and 11-6 in this book. Tools for moni-toring PSD online, as mentioned in Section 2.10.2, are very applicable here.

6.6.3 Impinging Jet Crystallizer

Impinging jet crystallization achieves the opposite extreme of mixing, compared with flui-dized bed crystallization, of high shear and energy input in small regions. As shown inFig. 6-7, two (or multiple) streams of high velocity are impinged upon each other, whichresults in a rapid localized intense mixing of these streams. Depending upon the mixing

CrystallizerDissolver

Cool

Heat

Figure 6-6 Fluidized bed crystallizer and dissolver.


intensity and detailed design of the impinging jet device, the mixing time of the two streamscan be fractions of a second and can be shorter than the induction time. As a result, thisdevice can be used to generate a high level of supersaturation for a nucleation-dominantprocess. At the optimum mixing effectiveness, an impinging jet can mix the feed streamsto the molecular level in less than the induction time, thereby allowing nucleation underuniform conditions throughout the operation.

This principle has been successfully utilized in an industrial application to achieve asmall average particle size (3–5 microns) and a narrow PSD. For impinging jet crystalliza-tion, industrial operation is described by Midler et al. (1994), with variants by Lindrud et al.(2001) (impinging jet crystallization with sonication) and by Am Ende et al. (2003) (specificreference to reactive crystallization). Laboratory studies are reported by Mahajan and Kirwan(1996), Benet et al. (1999), Condon (2001), and Hacherl (Condon) (2003). Johnson (2003)and Johnson and Prud’homme (2003) report on the use of impinging jets to produce nano-particles stabilized by block copolymers.

EXAMPLE 6-1

Process: Crystallization of an intermediate (MW � 700) from a very impure mixture.

Issues: Slow growth rate, shear damage during long approach to equilibrium. Loss of yieldbecause actual equilibrium solubility was not achieved.

Fluid feed Fluid feed

Impinging jetmixer

Figure 6-7 Impinging jet crystallizer.


Probable Cause(s)

† Accumulation of the impurity(ies) at the crystallizing surfaces, limiting the growth rate.

† Thick films around the crystals because of limited impeller speed to prevent crystal damage.

Possible Solutions

† Extend the age time (provides more time for the solute to reach the crystal surface).

† Remove impurities (reduces impurity accumulation at the surface and thereby speeds up thegrowth rate).

† Increase impeller speed (reduce film thickness) or reduce it (less shear damage).

Laboratory evaluation indicated that the expected equilibrium solubility could be achievedthrough longer aging (in the laboratory). However, the plant operation did not improve with longeraging.

Removal of impurities at this stage was not feasible. The intermediate produced in this step isrecrystallized to improve purity, and good growth is achieved in the purified system. Therefore, it isknown that the compound will grow.

SOLUTION: It was observed that adequate growth occurred early in the process and then stopped.To test the ability of fluid forces to reduce the presumed impurity film on the crystal surfaces, impellerspeed was increased. An increase in impeller speed by �20% after growth had stopped was effective inreaching the equilibrium and achieving the expected yield.

Conclusions

An increase in impeller speed was effective either because (1) the impurity film was reduced, allowinggrowth to resume, or (2) the increased shear caused additional nucleation.

Message

Impurity(ies) can have a profound effect on the growth rate and, in this case, on the approach to equili-brium solubility. As in many cases due to experimental time constraints, the actual cause of theimprovement was not clearly established. A very effective method of determining the impact of impur-ity(ies) from the process in question is to recrystallize the compound from a pure solvent(s) and fromits own mother liquor and compare crystal size, growth rate, and morphology.

Incorporation of impurities in the crystals was a major concern that indicated higher impellerspeeds. However, excessive fines and the resulting poor filtration rates were counterbalancing influ-ences on the determination of impeller speed. The two-level agitation rate scheme was a balancebetween these conflicting factors that resulted in passable purity, yield, and filtration rate. As inmany high-impurity systems, however, the average particle size was small because of the need tooperate at relatively high supersaturation to achieve practical growth rates, thereby incurring morenucleation than desired.

The qualitative aspects of this difficult crystallization provide examples of the conflicting require-ments that are often encountered in development and scale-up. B


Documents

Crystallization of Organic Compounds || Mixing and Crystallization