In situ product recovery (ISPR) by crystallization: basic principles, design, and potential...

This is the authors’ version before publication in Applied Microbiology and Biotechnology

(2006) volume 71 pages 1-12

In Situ Product Recovery (ISPR) by Crystallization:

basic principles, design, and potential applications

in whole-cell biocatalysis

Evelyn M. Buque-Taboada1,2, Adrie J.J. Straathof1*, Joseph J. Heijnen1,

and Luuk A.M. van der Wielen1

1 Department of Biotechnology, Delft University of Technology,

Julianalaan 67, 2628 BC Delft, The Netherlands

2 Department of Chemical Engineering, University of San Carlos, Talamban,

Cebu City 6000, The Philippines

*To whom correspondence should be addressed: Tel.: +31-15-2782330.

Fax: +31-15-278 2355. Email: [email protected].

ABSTRACT

Removal of inhibiting or degrading product from a bioreactor as soon as the product is

formed is an important issue in industrial bioprocess development. In this review, the

potential of crystallization as an in situ product removal (ISPR) technique for the biocatalytic

production of crystalline compounds is discussed. The emphasis is on the current status of

crystalline product formation by metabolically active cells, for application in fine-chemicals

production. Examples of relevant biocatalytic conversions are summarized, and some basic

process options are discussed. Furthermore, a case study is presented in which two

conceptual process designs are compared. In one process, product formation and

crystallization are integrated by applying ISPR, whereas a second, non-integrated process is

based on a known conventional process equivalent for the production of 6R-dihydro-

oxoisophorone (DOIP). The comparison indicates that employing ISPR leads to significant

advantages over the non-integrated case in terms of increased productivity and yield with a

corresponding decrease in the number of downstream processing steps as well as in the

quantity of waste streams. This leads to an economically more interesting process alternative.

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Finally, a general outlook on the various research aspects of in situ product removal by

crystallization is given.

Keywords: In situ product removal, crystallization, fermentation, process design, 4-

oxoisophorone, process integration, 6R-dihydro-oxoisophorone

In Situ Product Recovery (ISPR)

Often biocatalytic processes have limited productivity and yield, which may be due to

product inhibition or product degradation (Lye and Woodley 1999). Thus, attempts to

decrease this inhibition and degradation by optimizing physiological and technological

parameters are essential in the development of biocatalytic processes. Keeping the dissolved

product concentration low in the reactor can obviously circumvent these limitations. An

approach that can accomplish this task is the implementation of an in situ product recovery

(ISPR) technique (Alba-Perez 2001; Freeman et al. 1993; Lye and Woodley 1999; Schügerl

2000; Schügerl and Hubbuch 2005; Stark and von Stockar 2003; van der Wielen and Luyben

1992; von Stockar and van der Wielen 2003).

ISPR techniques lead to the processes where the product is removed from the catalyst as soon

as the product is formed in the reactor. In doing so, the waste streams as well as the number

of downstream processing steps can be reduced. ISPR is part of the general concept of

process integration or intensification (Freeman et al. 1993; Lye and Woodley 1999; Schügerl

2000; Schügerl and Hubbuch 2005; Stark and von Stockar 2003; van der Wielen and Luyben

1992; von Stockar and van der Wielen 2003), which represents the improved coordination of

upstream, reaction, and downstream technologies. The selection of a proper ISPR strategy,

appropriate process configuration and best mode of operation must be considered a priori in

the implementation of this technique in biocatalytic processes.

The techniques employed for in situ product removal depend on the properties of the target

product, on the expected benefit that ISPR needs to achieve, as well as on the biocatalyst

involved. Metabolically active cells are the biocatalysts focused on in this mini-review. By

far, extraction, adsorption and evaporation have been the most commonly employed

techniques for ISPR with these whole cells (Alba-Perez 2001; Lye and Woodley 1999;

Schügerl 2000; Schügerl and Hubbuch 2005; Stark and von Stockar 2003). In some cases,

electrodialysis, precipitation, complexation, and membrane-assisted separation techniques

have been employed (Alba-Perez 2001; Fernandes et al. 2003; Lye and Woodley 1999;

Schügerl 2000; Schügerl and Hubbuch 2005; Stark and von Stockar 2003). Product

categories involved are organic solvents, organic acids, aromas and flavors, secondary

metabolites and proteins.

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The basic configurations for in situ product removal as applied to the established techniques

are shown in Figure 1. Separation of the cells and product can be achieved by employing: (a)

an internal configuration, where the fermentation liquid is in contact with a second, product

removing phase that is present in the fermenter; or (b) an external configuration, where the

fermenter liquid is via a loop in contact with a product removing phase in an external unit

(Alba-Perez 2001; Freeman et al. 1993; Stark and von Stockar 2003).

Although the external configuration does not apply in situ product removal in a strict sense, it

often has been considered an ISPR configuration as it intends to reduce the product

concentration in contact with the cells.

In either the internal or external configuration, the cells can be in direct or indirect contact

with the product removing phase. The choice of configuration is also dependent on the

properties of the target product and biocatalyst, the number of relevant phases involved as

well as the type of recovery process employed. In most approaches, introducing a product

removing phase introduces auxiliary components, which complicates the process design and

potentially increases production costs.

[insert Figure 1 here]

A process configuration with indirect contact of the cells is well-established and

straightforward. This can be done by microfiltration or ultrafiltation (for external ISPR) or

by cell immobilization (for internal ISPR), for example. Thus, fermentation can be run close

to optimum operating conditions. However, it is necessary to achieve controlled product

removal at a reasonable rate without making large changes in pH or temperature between

fermenter and separator as this could potentially complicate liquid recycling to the fermenter.

In situ product recovery by crystallization during a reaction catalyzed by metabolically active

cells has hardly been considered. This may be due to complications caused by cell-crystal

interactions. Nevertheless, when in situ product crystallization is employed, it might directly

provide the desired product (in solid or crystal form) without the need for auxiliary

components in the product removing phase.

Crystallization is one of the oldest yet the most effective processes for the preparation of

solid products. High purities and stabilities can be obtained with distinct morphologies, and

functional properties. Crystallization is important in different areas of biotechnology. For

instance, crystallization of proteins, viruses, nucleic acids, and macromolecular complexes

has been done for a variety of reasons such as for X-ray diffraction analysis, as part of the

formulation of pharmaceutical or fine-chemical products, and as a means of purification in

bioseparation processes (McPherson 1999).

Microbial formation of crystalline products

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When a certain substrate S enters a microbial cell, it is converted to product P as

schematically shown in Figure 2. The cell then has at least two options on how to deal with

the product P. In some cases, the cell does not secrete the product and tends to store it.

When this happens, the concentration of the product in the cell (CP,in) increases and will be

greater than the concentration outside the cell assuming a continuous import of substrate into

the cell; then, CP,out is low or zero as CP,in is maximized in the cell (Figure 2A). This leads to

the accumulation and potential crystallization of the product inside the cell, which sometimes

might result in cell damage. However, when the cell needs to maintain a low concentration

of the product inside the cell (e.g. due to its toxicity), it has to secrete the product; then, CP,out

is greater than CP,in as CP,in remains low in this case (Figure 2B). When this occurs, product

crystals may start to form outside the cell when the solubility limit is reached. In this sense,

the crystal formation of certain molecules may be intracellular (in the cell) or extracellular

(outside the cell). If secretion is slow as compared to product formation, both intracellular

and extracellular crystals might be formed.


The crystal formation may involve extremely wide ranges of compounds and conditions. For

example, calcium lactate formation may occur on Cheddar cheeses as a result of microbial

action (Johnson et al. 1990). Plant cells may form storage polymers such as cellulose having

crystalline regions (Astley et al. 2001). Plasmodium parasites accumulate heme crystals (Iyer

et al. 2003). Also, crystalline proteins may accumulate, for example for the toxins of Bacillus

thuringiensis used as insecticide (Agaisse and Lereclus 1995). Paramecium species even

seem to discharge protein-crystal containing trichosyst organelles, as a defense against other

microorganisms (Plattner 2002). An antibiotic of the nifimycin A1-scopafungin group can

spontaneously crystallize during submerged fermentation of Actinomyces hygroscopicus

(Kuimova and Kazakov 1976). More importantly, microorganisms play a role in the

deposition of many minerals, for example for polyphosphates, calcium carbonates, silicas,

metal sulfides, metal oxides, and even gold (Ehrlich 1999). Mineral deposition can be

intracellular as well as extracellular.

Table 1 shows that crystal formation also occurs in a wide range of cases relevant to

industrial biotechnology. Examples from reactions catalyzed by enzyme preparations have

not been included because they are out of the scope of this mini-review. The examples given

in Table 1 cover only a small part of the crystalline industrial biotechnology products,

perhaps because in most cases product toxicity or degradation occurs before saturation

concentrations are achieved. Thus, microbial conversions can be exploited for the direct

production of crystalline products of organic, inorganic, or biological nature.

[insert Table 1 here]

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In situ product removal by crystallization

In chemical technology, processes are often referred to as reactive crystallization

(Stankiewicz and Moulijn 2004) if an internal configuration is used as shown in Figure 1. For

several enzymatic processes, a crystallization step has been integrated with the reaction step,

using either crude or pure enzyme (Arimatsu et al. 2004; Bao et al. 2001; Fernandes et al.

2003; Furui et al. 1988; Furui et al. 1996; Harano et al. 1986; Michielsen et al. 2000a-b;

Nakayama 1985; Takamatsu and Ryu 1988a-b; Tosa et al. 1988). In contrast, in

fermentations and biotransformations using resting cells, crystalline product formation has

hardly been seen as a potentially beneficial phenomenon for an efficient product recovery.

When crystals are formed and stored intracellularly, this requires cell disruption to release

and recover the crystals. Then, product crystal removal cannot be performed during the

microbial conversion. On the other hand, when the product is secreted and crystals are

formed extracellularly, product removal can be performed during the microbial conversion by

1) inducing crystallization as soon as supersaturation is reached, for example by seeding, and

2) separation of the crystals from the cells. This separation may be based on size, density, or

surface tension differences. Complexation can also be exploited to give a distinct property to

the target product (Lye and Woodley 1999). Centrifugation, flotation and interfacial

partitioning may be considered (Jauregi et al. 2000; van Hee et al. 2006). In case of release

of supersaturation, such in situ product removal would lead to a local decrease of dissolved

product concentration, and therefore probably to increased productivity and yield in case of

product inhibition or degradation. If the presence of extracellular crystals causes stress and

mixing problems, removal of product crystals may be favorable to the cells.

However, the separation of product crystals from the cells is usually not carried out in situ

but, for example, by addition of organic solvent, hence precipitating the cell debris and

dissolving the crystals. Subsequently, the product is crystallized again (Fukuoka et al. 2001;

Leuenberger 1985; Leuenberger et al. 1976). Such procedures lead to high amounts of waste

streams and high production costs.

On the other hand, in a large number of whole-cell processes, significant product inhibition or

degradation occurs already at a concentration below the solubility limit and product crystals

are hardly or not formed in the fermenter. In these situations, ISPR by external

crystallization with indirect contact between cell and product-containing phase will be an

attractive solution. The product is removed from the fermenter as soon as it is formed,

thereby, lowering intracellular and extracellular product concentrations and maximizing the

fermentation process performance. The subsequent discussion will focus on ISPR by

external crystallization with indirect contact.

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Cell retention options

Figure 3 shows some options for cell retention in ISPR by crystallization with external

configuration. In these schemes, product crystallization is performed outside the bioreactor.

In Figure 3A, a membrane filter separates the product-containing supernatant (liquid) from

the cells. While the retentate is recycled to the reactor, the filtrate is directly fed to the

crystallizer where the target product is crystallized. The product-depleted mother liquor in

the crystallizer is also recycled to the reactor. In this case, it is also important to consider the

residence time of the fermentation broth in the membrane filter as the cells may suffer from

nutrient or oxygen limitation. In scheme B, a spin filter or an acoustic perfusion system is

installed in the bioreactor for cell retention. This strategy, which is usually applied for

processes employing mammalian cells (Zhang et al. 1998), allows the direct and continuous

withdrawal of the cell-free liquid broth from the bioreactor while the cells are retained,

avoiding cell stress and potential nutrient or oxygen limitation. However, complete cell

retention is difficult and would require a second-stage filtration unit, making this option less

attractive in combination with a crystallization process. Immobilized cells can also be used

(scheme C) where cells are retained in the bioreactor. This leads to a relatively simple

configuration, but requires an immobilization step. Of course, Figure 3 is by no means

exhaustive.


In conventional processes, ultrafiltration precedes crystallization from a cell-free

fermentation broth. Ultrafiltration removes lysed cell components and biopolymers of

>10,000 g/mol. These may otherwise inhibit crystallization (Li et al. 2004). Therefore, an

attractive option is to use the set-up of Figure 3A with ultrafiltration instead of microfiltration

(cf. Buque-Taboada et al 2004).

Criteria for process implementation

ISPR by external crystallization is logically applied if the target commercial product is solid

or crystalline at storage conditions. The following criteria must be considered in deciding

whether it is applicable or beneficial for a certain whole-cell catalyzed process. Some are

related and readily known. It must be noted that the goal is to remove the product as soon as

it is formed in the reactor to avoid product toxicity or degradation, thereby, increasing

fermenter productivity and yield.

Product toxicity, degradation, and solubility

When the product is toxic to the microorganism or is degrading, the dissolved product

concentration in the reactor level must be kept low because the concentration that leads to

inhibition or degradation will usually be below the solubility. However, product

supersaturation must be achieved in the crystallizer unit. Thus, it is imperative to determine

7

the allowable dissolved product concentration in the reactor in relation to the solubilities of

the product and also of other compounds involved in the process at the conditions in the

reactor as well as at the conditions in the crystallizer.

For 1/3 of a series of 24 liquid and solid substrates and products involved in microbial

biotransformations, the toxicity concentration was in the range of 50-90% of these products’

aqueous solubility (Straathof 2003). The subset of microbial products that are room

temperature solids is not sufficiently represented in this series, but suppose that also for 1/3 of

all solid microbial products the toxicity range is 50-90% of their aqueous solubility. Also,

suppose that the microbial process for production of a solid is carried out up to its toxicity

concentration. A reduction of solubility by 50% in an external loop, for instance by cooling

from 30 to 5 oC, might allow crystallization in that loop for this 1/3 fraction of the products

considered. This preliminary estimate indicates that ISPR by crystallization might be

applicable in a significant number of cases.

In any case, the crystallization rate must be able to keep up with the net production rate to

implement ISPR by external crystallization in the process. Otherwise, in situ product

removal by crystallization is less efficient, if not illogical. Thus, a typically reasonable

crystallization rate must be assumed, if not yet determined.

Product formation rates and volumetric productivity

The production rate must be fast enough to achieve the product solubility limit (saturation) in

the reactor or in the crystallizer rather quickly. For economically attractive biocatalytic

processes, typical final product concentrations are 50-100 g.L-1 (Straathof et al. 2002). For

biocatalytic processes involving oxygen transfer, a typical volumetric productivity of 1 g.L-

1.h-1 is assumed. This implies a typical residence time of 50-100 h (2-4 days) for such a

process. Furthermore, crystal growth rates of well-soluble compounds are typically in the

range of 10-8 – 10-7 m.s-1. Simple separation units are capable of handling particles ranging in

size of 0.2 - 1 mm. To reach these sizes, generally a residence time of crystals in the

crystallizer of 0.5 to 5 h is needed. This implies that the product formation rate will be rate-

limiting and integrated fermentation-crystallization processes will operate at low

supersaturation ratios such that seeding may be required in the crystallizer.

Biocatalyst

When whole-cells are employed as biocatalyst, the best physiological state of the cells must

be considered. The start of the reaction, which is often dictated by the time of precursor

addition, has a profound influence on productivity. Enzyme induction may necessitate

addition of small amounts of precursor at the beginning. Addition of precursor in the late

exponential (state) growth phase usually avoids toxicity effects. In some cases, the options of

growing or resting cells should both be considered, because the metabolic activities (thus,

also biocatalytic activities) of each are not the same.

8

When growing or susceptible cells are employed, it is usually imperative to keep the process

under aseptic conditions, although it may be difficult or expensive to implement this

situation. However, after the cell growth stage, the bioconversion phase may be performed

under less strict hygienic conditions. In this case, cell growth and cultivation can be carried

out in one large fermenter and the harvested cells are stored properly for subsequent

bioconversion.

Mode of operation

When the biocatalytic process is a fed-batch or a continuous process operating at a constant

volumetric productivity, the process might reach a steady state with respect to precursor and

dissolved product concentrations in which the production rate in the bioreactor is equal to the

product crystallization rate. Product crystals may be kept in the crystallizer or continuously

removed. The continuation of the process will probably be dependent on sustained

biocatalytic activity. ISPR may promote sustained biocatalytic activity because product

toxicity and by-product formation are minimized. If not purged, however, by-products will

still accumulate and hence may achieve levels where interference with the biocatalyst or

crystallization becomes unacceptable.

Product crystal morphology and purity

When ISPR by external crystallization is implemented, product crystals are formed in the

presence of reaction medium containing other solutes and impurities. These can have varying

effects to the crystallization of the product (McPherson 1999; Mullin 2001; Tavare 1995), but

so far, this is still difficult to predict. In some cases, co-solutes can accelerate or inhibit

crystal formation and growth; sometimes, they have no effect at all. In other cases, the

solubility of the crystal in the reaction medium may be lowered by the presence of these co-

solutes, thus, rendering crystallization as a favorable recovery technique.

Crystallization involves nucleation and crystal growth. These two sub-processes must be

considered in order to describe the crystallization behavior of the product crystals and the

influence of impurities. While it is important to determine crystallization kinetics of a certain

product, this should not hinder the implementation of ISPR in the process. Supposing that

the desired crystal morphology and purity are affected by the presence of co-solutes or

impurities in the reaction medium, recrystallization is the common approach to solve the

problem (Mullin 2001; Tavare 1995).

Energy consumption

In some cases, it is easy to imagine that energy consumption (i.e. due to cooling or

evaporation) may be significant. Heat integration of the process may be required. Recycling

cold, product-depleted, mother liquor to the fermenter, can help in keeping the fermenter

temperature constant, thus, external cooling may not be necessary, resulting in less energy

9

costs and less investment. Over all, the choices to be made should be driven by and should

simply result in more efficient and cost-effective processes.

Case study: Synthesis of 6R-dihydro-oxoisophorone

To explore the potential of ISPR by external crystallization a conceptual process is designed

and compared with a known conventional process option. This involves the synthesis of 6R-

dihydro-oxoisophorone (DOIP), a white crystalline solid, with a solubility of about 10 g.L-1

in water at the fermentation temperature of 30oC (Buque-Taboada et al. 2004; Leuenberger et

al. 1976). DOIP is used as a key intermediate in the production of some carotenoids

(Leuenberger 1985; Leuenberger et al. 1976) and saffron flavors (Sode et al. 1987). The

synthesis reaction (Figure 4) involves the asymmetric reduction of 4-oxoisophorone (OIP)

using baker’s yeast (Saccharomyces cerevisiae) (Buque-Taboada et al. 2004; Buque-Taboada

et al. 2005a-b; Fukuoka et al. 2001; Leuenberger 1985; Leuenberger et al. 1976) or

Saccharomyces rouxii (Fukuoka et al. 2001) as biocatalyst. The main product, DOIP, is

degraded by the yeast mainly to (4S,6R)-actinol, an unwanted by-product in the process

(Buque-Taboada et al. 2004; Buque-Taboada et al. 2005a-b; Leuenberger 1985; Leuenberger

et al. 1976). These yeasts are capable of producing DOIP with an enantiomeric excess of

98% (Buque-Taboada et al. 2004; Buque-Taboada et al. 2005a-b; Fukuoka et al. 2001;

Leuenberger 1985; Leuenberger et al. 1976).


As product degradation occurs in the fermenter, it is imperative to keep the product

concentration low; this can be done by employing an ISPR technique where the product is

removed as soon as it is formed in the reactor.

Little is known about the current industrial production process for DOIP and its current

production levels, as it probably mainly is an in-house product. However, recent patent

literature by F. Hoffmann-La Roche AG (Fukuoka et al. 2001), one of the producers of

zeaxanthin and some flavors, describes a process for the production of DOIP (also known as

levodione) with immobilized yeast biomass. For the sake of comparison, this process is

regarded as the base case and the process description is changed slightly to fit the

specifications set in the case study, as will be discussed later on.

Considering its application, DOIP has a rather small production volume. The design

challenge is to produce 6 tons per year of DOIP crystals in the size range of 0.01 - 0.1 mm

(Buque-Taboada et al. 2005a,c) with a purity of 99.5 %w/w and an enantiomeric excess (e.e.)

of 98%. Inherent to this challenge is to compare the production costs per kg of DOIP of the

process employing ISPR by crystallization and the base case on the same basis.

10

The conceptual process design (CPD) approach used in this case study is described in detail

elsewhere (Coulson et al. 1998; Douglas 1988; Grievink et al. 2004; Lide 2004; Perry and

Green 1999; Reid et al. 1977; Sinnott 1999; van’t Riet and Tramper 1991). The procedures

indicated therein have been carefully followed and this work is the condensed and simplified

version of the detailed CPD. Economic evaluation is done on the same basis considering

current costs of raw materials, equipment and utilities.

Process Design

For the case with in situ product crystallization (ISPC) with external configuration the

biocatalyst used is a typical baker’s yeast (S. cerevisiae) with reaction kinetics and production

rates described previously (Buque-Taboada et al. 2004; Buque-Taboada et al. 2005b). For

the base case, immobilized S. rouxii is employed with reaction rates and productivities as

described (Fukuoka et al. 2001). An operating time of 6000 h/a is chosen for the production

of 6 ton DOIP.

Process description

The process flow diagram of the base case is shown in Figure 5A-5B. The base case employs

a bubble column fermenter with immobilized cells as biocatalyst as this gave good

volumetric productivity and low estimated costs (Fukuoka et al. 2001). For downstream

processing, extraction or adsorption are options prior to DOIP crystallization from organic

solvent. Extraction with ethyl acetate (Leuenberger et al. 1976) is chosen for the base case

process to avoid too much speculation in the design of adsorption columns where the required

parameters are not available. Crystallization occurs from the ethyl acetate extract upon

cooling and solvent evaporation (Fukuoka et al. 2001). A single multipurpose vessel is used

for extraction and crystallization as fermentation is the process rate-limiting step according to

the batch cycle diagram.

The process flow diagram for the ISPC case is shown in Figure 5C. For this process, a

stirred-tank fermenter coupled with an external crystallizer is considered where an

ultrafiltration unit is placed in between these units to separate the liquid from the cells and to

recycle the cells to the fermenter as described previously (Buque-Taboada et al. 2004).

Crystallization is done by cooling to 5oC and for process simulation, the crystallization

kinetics described previously (Buque-Taboada et al. 2006) is adapted. The filtration of the

product crystals is done using a Nutsche filter and nitrogen gas is employed for crystal drying

to prevent explosions.

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Basic Assumptions

Location and plant life

It is desired to perform the production in a multipurpose plant (MPP). At such a plant there is

a wide variety of process equipment available for processing fine chemicals. The equipment

is used for different production processes over the year.

At the MPP, infrastructure for communication and transportation of raw materials and

products are available. The facility is able to supply heat transfer oil at a wide temperature

range, approximately between –10 to 150oC and electricity, among others. A wastewater

treatment plant (WWTP) and waste handling facility is also installed in the MPP. It is

assumed that the MPP facility will be partly used for DOIP production when a smaller

production capacity is desired, but for the sake of economic analysis, a full-time usage of the

MPP is adapted.

The economical plant life for the design cases is assumed to be 15 years. Conventional

equipment, which is usable in the production of other fine-chemical products, has an

equipment life of 15 years while unconventional (specialized) equipment has an assumed life

of 10 years.

Battery limit

The biomass used as catalyst is cultivated outside the battery limit for both cases. However,

regeneration of the immobilized biomass is implemented within the battery limit for the base

case.

Further, the bioconversion and downstream processing equipment are within the battery limit.

For the base case, this includes units for reaction, filtration, extraction, evaporation,

crystallization and solid handling. For the ISPC case, devices are limited to reactors, filters,

crystallizers and the solid handling units. For both cases auxiliary equipment is desired such

as storage vessels, pumps, piping, valves, and process control devices.

The fermentation (feedstock) solution entering the battery limit is composed of many

different components usually needed in small quantities. For simplicity, this is condensed to

three types of solution: media solution, reactivation solution (for biomass regeneration in the

base case), and the biomass suspension (used in the ISPC case).

Wastewater treatment is done outside the battery limit. The immobilized biomass cannot be

processed in a WWTP and needs to be discarded in another way at the end of the process. In

the base case, the spent solvent needs to be discarded in an organic waste treatment facility.

Utilities such as electricity, heating and cooling are inside the battery limit.

Process block schemes

A simplified process flowsheet for the base case is shown in Figure 6 where the mass

streams, pressures and temperatures are indicated. The thicker lines show the flow of main

reactant (OIP) and product (DOIP). The values indicated (in ton per ton of product) in each

12

process stream have been used in the economic analysis. In this case, the overall process

yield of DOIP on OIP was 79 %w/w during the production of 6.2 tons of DOIP in

approximately 6000 hours. Details of the major equipment employed in the process are

summarized in Table 2.



A simplified flowsheet of the ISPC case is shown in Figure 7 where the yield per process step

is indicated. The overall process yield of DOIP on OIP is 91%w/w. A process cycle of one

batch is about 400 hours, which results in a total of 15 batches a year to produce the desired

amount. The ISPC process has characteristics of both a continuous and a batch process;

however, the streams for a continuous process indicated are employed in the economic

analysis. Details of the major equipment required in the process are summarized in Table 3.



It should be noted that only the final, optimized process designs are shown. These may seem

to contain differences, for example with respect to biocatalyst choice, that might have been

obviated for a fair comparison. However, to obtain a realistic comparison, we did not block

potential process improvements even if these were only favorable to only one of the cases.

Process analysis

Base case

The base case design applies proven technology. There is sufficient know-how of the sub-

processes, which are widely applied in industry. This makes the base case process a reliable

one with respect to the individual operations of the equipment. The process is flexible but the

overall operation is not straightforward, which can cause potential setbacks (i.e. disruption of

equipment scheduling) on the over-all production process.

Immobilized biomass as the biocatalyst is chosen so that volumetric productivity is high. The

short batch process time in a bubble-column fermenter is an attractive feature of the process,

as it reduces the required equipment size.

The base case has an over-all yield of 79% w/w, which is not attractive when a relatively

expensive substrate is employed, as is the case for OIP. Furthermore, a large amount of

process water (202 ton per ton product) is required, resulting in a dilute product stream and

large waste streams. The large amount of organic solvent required (18.4 ton per ton of

product) leads to considerable costs.

13

ISPC case

The ISPC case is based on new lab-scale technology, which still needs to be tested (e.g. on

pilot-scale) to improve process reliability. The ISPC case employs less equipment (thus, less

manpower) but its semi-continuous and integrated nature can be disastrous to the production

process itself. The process requires a relatively large fermenter with stirring equipment,

which is more costly to install, operate, and maintain than a bubble-column fermenter

employed in the base case.

The ISPC case has an over-all yield of 91% w/w, which is a substantial improvement of the

base case having a yield of 79% w/w. This is due to the fact that product degradation is

minimized, stream recycles are implemented, a long process time is implemented, and water

build-up in the system is prevented. Process water required is 14.4 ton per ton product,

which is 14 times smaller than in the base case. In addition, the ISPC case does not use

organic solvent, which leads to an even smaller and easy-to-treat waste stream.

Most equipment choices for the ISPC case are attractive options. The ultrafilter does not lead

to high costs, the stirred-tank fermenter has higher yield (although costly) and cooling

crystallization is efficient. The stream recycles are an important feature of the design,

however, careful implementation is needed to avoid potential process disturbances and

failures.

Economic Analysis

The cost of the main substrate OIP is estimated at 100 euro/kg and DOIP, being four times

more expensive in laboratory catalogues, at 400 euro/kg. All other costs were estimated

according to available sources. The total cost as well as the production cost per kg of DOIP

produced in each case is shown in Table 4. The ISPC case has lower annual production costs

and thus a higher net cash flow. It has a higher economic potential, scoring better on all

economic criteria (Table 5). This is attributed to the higher yield and the absence of organic

solvent usage. In the base case, the lower yield causes a lower net cash flow and the use of

organic solvent resulted in higher variable costs, which includes waste costs. Although the

purchased equipment costs (PEC) of the ISPC case are higher, the final total investment

required is the same for both processes. The higher PEC for the ISPC case is due mostly to

the high costs of the fermenter. The variable costs due to cooling in the crystallization loop

in the ISPC case turn out to be modest.



Possible economic improvements

Since substrate cost is a major cost factor as it is relatively expensive, improving the yield of

DOIP on OIP is recommended to increase the net cash flow especially in the base case. A

cheaper, more efficient, easy-to-dispose-of kind of organic solvent must be employed if it

14

cannot be avoided in the base case, so that the utilities and waste costs are lowered.

Reducing the consumption of process water and organic solvent will save on utilities and

waste treatment. The ISPC case is in many ways already an improvement of the base case,

but can still be enhanced. The yield and productivity can be improved further by optimizing

flow and recycle streams and by employing a biocatalyst with higher activity, amongst

others. Table 6 shows the DOIP productivity employing yeasts at different conditions. Free

cells of S. rouxxi show productivity, which is almost 3 times better than S. cerevisiae.

Growing cells of S. cerevisiae show the highest productivity but this might be economically

compensated by the costs of nutrients required in growing cells. Alternative fermenters

should be explored as the current stirred-tank fermenter (including its operation) is a major

cost factor in the process. Upscaling would significantly reduce investments for both options.

Although the acquisition of biomass feedstock is outside of the battery limit, it is

economically attractive to consider acquiring or producing it on a yearly basis by cultivating

the biomass in a large rented fermentation facility and storing it properly to maintain its

viability throughout the year.


Conclusions and outlook

Crystallization as an in situ product recovery technique might be an attractive option in the

production of crystalline compounds employing whole cells as biocatalyst. The case study

presented as an example, shows that employing in situ product crystallization in a typical

whole-cell catalyzed reaction can lead to increased productivity and yield with minimal waste

streams, at lower production costs than a conventional process. The ISPC process does not

require the use of auxiliary phases such as organic solvents. This process strategy may be

applicable to many biocatalytic (e.g. reduction) systems where crystalline products are

obtained, which include specialty biochemicals such as amino acids, steroids, antibiotics,

proteins, esters and ketones. Furthermore, this might be an attractive option for processes

employing relatively expensive substrates.

However, the design approaches and choices for such a process are crucial for its successful

implementation. As this study is the start of a systematic development for ISPC processes,

further research is required to unravel the different factors and issues concerning this process.

In the end, in situ product recovery by crystallization should be a common practice in the

design and implementation of robust integrated processes involving fermentation or

biotransformation.

More specifically, the following issues of the process must be addressed:

Extensive application in biocatalyzed processes, leading to a wider scope of product and

biocatalyst categories,

Strategies for ISPC implementation under aseptic conditions,

15

Approaches in particle-particle separation,

Crystallization aspects (e.g. generalized correlation on the influence of co-solutes on the

product crystal, improved (re-)crystallization techniques),

Improvement of process reliability and consistency (by pilot-scale studies),

Process optimization and comparison to alternative ISPR methods.

Acknowledgements

This study has partly been funded by the MHO-USC-DUT Project in Chemical Engineering.

The following persons are gratefully acknowledged for their important contribution to this

work: Sjoerd Blokker, Marcel Dabkowski, Willem Groendijk, Dirk Renckens, Jeroen de

Rond, and Prof.dr.ir. Johan Grievink of the Delft University of Technology.

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20

TABLES Table 1. Industrial biotechnology products formed as crystals by microbial action.

Crystalline products Examples References

Amino acids Glutamic acid Ueda et al. 2003

Carboxylic acids 6-hydroxynicotinic acid Hurh et al. 1994

Carboxylates Calcium lactate Kaščák et al. 1996

Chiral alcohols Many Blacker and Holt 1997; Chartrain et

al. 1999; Matsumae et al. 1999;

Leuenberger et al. 1976; Leuenberger

1985; Fukuoka et al. 2001

Steroids Many Crocq et al. 1997; Miller 1985

Antibiotics Tetracycline, oxytetracycline Cardoso 1993

Pigments -carotene Dufosse et al. 2005

Vitamins Riboflavin van Loon et al. 1996

Table 2. Specifications of major equipment used in the base case.

Specification Fermenter Extractor Crystallizer Filters

R101 S201 S201 S101 S202

Type Bubble

Column

MPVa MPVa Microfilter Nutsche

Volume [m3] 1.65 4.5 4.5

Area [m2] 0.44 1.33 1.33 0.07 0.8

Flux [m3.m-2.h-1] - - - 49.2 1

Number 1 1 1 1 1

- series 1 1 1 1 1

- parallel - - - - - a Multipurpose vessel with agitation accessories.

Table 3. Specifications of major equipment used in the ISPC case.

Specification Fermenter Crystallizerb Filters

R101 S102 S101 S102 S103

Type Stirred-tank

reactor

Scraped cooling

crystallizer

Ultrafilter Microfilterc Nutsche

Volume [m3] 6 (4.1a) 3.1 - - -

Area [m2] 2.7 1.5 6 0.07 1.8

Flux [m3.m-2.h-1] - - 0.07 0.61 0.22

Number 1 1 2 1 1

- series 1 1 1 1 1

- parallel - - 2

21

a Fermenter working volume. b With internal microfilter. c Installed inside the crystallizer.

Table 4. Final economic figures in both cases.

Cost Items Base case ISPC case

Purchased Equipment Costs (PEC) [k€] 661 907

Investment Costs [k€]

Direct Costs (DC) 1,740 1,620

Indirect Costs (IC) 1,040 970

Direct Fixed Capital (DFC) 3,160 2,970

Total Investment costs (IT) 4,000 3,720

Operating Costs [k€/a]

Fixed Costs (FC) 860 580

Variable Costs (VC) 1,080 761

Raw materials 798 709

Consumables 17 9

Utilities 146 28

Waste treatment and disposal 104 1

Miscellaneous expenses 15 14

Other costs 129 87

Annual Production Costs (APC) [k€/a] 2,070 1,428

Income [k€/a]

Revenues 2,490 2,560

Net Cash Flow (NCF) 420 1,130

DOIP produced per year [kg/a] 6222 6393

Production Costs [€/kg] 333 223

Margin (NCF/kg DOIP) [€/kg] 68 177

Table 5. Economic criteria in both cases.

Criteria Base case ISPC case

Rate of Return [%] 11 31

Pay-back time [years] 10.5 4.3

DCFRR* [%] 6.4 26.1

22

Return on Investment [%] 2 17

*Discounted cash flow rate of return

Table 6. DOIP productivity by two important yeasts.

Yeast strain Productivity [mg.(g cells)-1.h-1] References

Resting cellsa Growing cellsa

Saccharomyces cerevisiae 10 50 Buque-Taboada et al. 2004;

2005a

Free cells Immobilized cells

Saccharomyces cerevisiae 10 - Leuenberger et al. 1976

Saccharomyces cerevisiae 9 0.7 Fukuoka et al. 2001

Saccharomyces rouxii 26 4.9 Fukuoka et al. 2001 a The productivity is in terms of gram dry weight of cells.

23

FIGURE CAPTIONS

Figure 1. Different configurations for ISPR.

Figure 2. Conversion of substrate S to product P in the cell.

Figure 3. External configurations for ISPR with cooling crystallization. Cell retention is

done by (A) external filter, (B) spin filter inside the bioreactor, (C) cell

immobilization. In all cases, the filtrate is directly fed to the external

crystallization loop to crystallize the product. The product-depleted mother

liquor is recycled to the bioreactor.

Figure 4. Synthesis of 6R-dihydro-oxoisophorone via reduction of 4-oxoisophorone.

Figure 5. Process flow diagrams for (A and B) base case and (C) ISPC case.

Figure 6. Simplified block scheme diagram of DOIP production process (base case).

Figure 7. Simplified block scheme diagram of DOIP production process (ISPC case).

24

Figure 1 (Buque-Taboada et al., 2005)


Internal External

Directcontact

Indirectcontact

Sout PoutSin PinPS

conversion

CPin < CPout

CPin 0

Sout PoutSin PinPS

conversion

CPin > CPout

CPout 0

A B

Sout PoutSin PinPS

conversion

CPin < CPout

CPin 0

Sout PoutSin PinPS

conversion

CPin > CPout

CPout 0

A B

25

4

5

32

1 Feed

ProductCrystals

Liquid medium recycle

Fermentatio Crystallization Filtrationwith in-situ (cooling)spin filter

4

5

32

1 Feed

Productcrystals


Fermentation Crystallization Filtrationwith immobilized (cooling)cells

Feed

product crystals

1

2

4

5

7

3 6

Biomass


Fermentation Filtration Crystallization Filtration (cooling)

A

C

B

26


27


(OIP)

O

O

O

O

O

O H

Yeast Yeast

4-Oxoisophorone Levodione or

(6R)-dihydro-oxoisophorone (DOIP)

(4S,6R)- (ACT)

28


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ecyc

ling

R10

1: F

erm

ento

rE1

02:

Con

dens

orS1

01:

Ultr

afilt

er

P101

: Fe

ed O

IPS1

02:

Cry

stal

lizer

P1

02:

Feed

Glu

cose

S103

: N

utsc

hefil

ter

P103

: Bi

omas

s su

sp.

V101

: C

onde

nsor

P104

: Fe

ed B

ase

P105

: Fe

ed m

isc.

P106

: Pr

oduc

t re

mov

alP1

07:

Rec

ycle

+

empt

yP1

08:

Slur

ry t

o S1

03

Pro

cess

Eq

uip

men

t S

um

mar

y fo

r B

ase

Cas

e

E101

: H

eatin

gR

101:

Bub

ble

colu

mn

E102

: C

onde

nsor

E103

: C

oolin

gS1

01:

Dea

d En

d Fi

lter

E201

: C

onde

nsor

S201

: M

ulti-

Purp

ose

V.

S202

: Fi

lter

K10

1: V

acuu

m p

ump

V101

: C

onde

nsor

P101

: Fe

ed b

ase

V102

: St

orag

e P1

02:

Filli

ng R

101

V201

: EA

med

ium

P103

: Em

pty

R10

1V2

02:

EA lo

wP1

04:

Empt

y V1

02V2

03:

EA c

lean

P2

01:

Fill

S201

V204

: St

orag

eP2

02:

Empt

y S2

01

V205

: St

orag

e P2

03:

Empt

y V2

04V2

06:

Con

dens

orP2

04:

Empt

y V2

05P2

05:

Empt

y V2

06

P206

: Em

pty

S202

BA

SE

CA

SE

: F

ER

ME

NT

AT

ION

S101

Pur

ified

air

V10

2R

101

P10

3P

104

Aq

ueou

s so

lutio

n to

V20

4

Spen

t ai

rV

101

P10

2F

erm

enta

tion

sol

utio

n

T10

3

Wat

er

T10

4

Rea

ctiv

atio

n so

luti

on

T10

5

Aqu

eous

sod

ium

hy

drox

ide

T10

1

P10

1Bub

ble

Col

um

n R

eact

orB

uffe

r V

esse

l

Spen

t re

acti

vat

ion

solu

tion

A

Wa

ter

to W

WT

P

Sp

ent

nitr

oge

n

Su

ppo

rtC

he

mic

als

T10

1

P1

01O

IP

T10

2

P1

02G

luco

se

T10

3

P1

03B

iom

ass

sus

pen

sion

Fil

tere

d

Air

T10

4

P1

04B

ase

T10

6S

eed

cry

sta

ls

Sp

ent

Air

R1

01

DO

IP c

rys

tals

P1

08

S1

02

Was

hin

g w

ate

r

Nit

roge

n

P1

06

P1

07

T10

5

P1

05

S1

01

S1

03

Wa

ter

to

WW

TP

E1

02

E1

01

V1

01

B

C

Spe

nt d

ryin

g g

as

DO

IP c

ryst

als

Spe

nt w

ash

Was

te w

ater

to

WW

TP

Eth

yl A

ceta

te

pur

ge

BA

SE

CA

SE

: D

OW

NS

TR

EA

M P

RO

CE

SS

ING

S20

1

EA

m

ediu

m

T201

EA

fres

h

Aq

ueou

s so

lutio

n fr

om V

102

P20

1

Was

hing

wat

er

V2

01

EA

lo

w

V20

2

V20

3

P20

2

V2

04V

205

P20

4P

203

S202

Nit

roge

n

EA

cl

ean

G

K2

01E

201

V2

06

P20

5

P20

6

Pro

cess

Eq

uip

men

t S

um

mar

y fo

r IS

PC

C

ase

E101

: R

ecyc

ling

R10

1: F

erm

ento

rE1

02:

Con

dens

orS1

01:

Ultr

afilt

er

P101

: Fe

ed O

IPS1

02:

Cry

stal

lizer

P1

02:

Feed

Glu

cose

S103

: N

utsc

hefil

ter

P103

: Bi

omas

s su

sp.

V101

: C

onde

nsor

P104

: Fe

ed B

ase

P105

: Fe

ed m

isc.

P106

: Pr

oduc

t re

mov

alP1

07:

Rec

ycle

+

empt

yP1

08:

Slur

ry t

o S1

03

Pro

cess

Eq

uip

men

t S

um

mar

y fo

r B

ase

Cas

e

E101

: H

eatin

gR

101:

Bub

ble

colu

mn

E102

: C

onde

nsor

E103

: C

oolin

gS1

01:

Dea

d En

d Fi

lter

E201

: C

onde

nsor

S201

: M

ulti-

Purp

ose

V.

S202

: Fi

lter

K10

1: V

acuu

m p

ump

V101

: C

onde

nsor

P101

: Fe

ed b

ase

V102

: St

orag

e P1

02:

Filli

ng R

101

V201

: EA

med

ium

P103

: Em

pty

R10

1V2

02:

EA lo

wP1

04:

Empt

y V1

02V2

03:

EA c

lean

P2

01:

Fill

S201

V204

: St

orag

eP2

02:

Empt

y S2

01

V205

: St

orag

e P2

03:

Empt

y V2

04V2

06:

Con

dens

orP2

04:

Empt

y V2

05P2

05:

Empt

y V2

06

P206

: Em

pty

S202

BA

SE

CA

SE

: F

ER

ME

NT

AT

ION

S101

Pur

ified

air

V10

2R

101

P10

3P

104

Aq

ueou

s so

lutio

n to

V20

4

Spen

t ai

rV

101

P10

2F

erm

enta

tion

sol

utio

n

T10

3

Wat

er

T10

4

Rea

ctiv

atio

n so

luti

on

T10

5

Aqu

eous

sod

ium

hy

drox

ide

T10

1

P10

1Bub

ble

Col

um

n R

eact

orB

uffe

r V

esse

l

Spen

t re

acti

vat

ion

solu

tion

A

Wa

ter

to W

WT

P

Sp

ent

nitr

oge

n

Su

ppo

rtC

he

mic

als

T10

1

P1

01O

IP

T10

2

P1

02G

luco

se

T10

3

P1

03B

iom

ass

sus

pen

sion

Fil

tere

d

Air

T10

4

P1

04B

ase

T10

6S

eed

cry

sta

ls

Sp

ent

Air

R1

01

DO

IP c

rys

tals

P1

08

S1

02

Was

hin

g w

ate

r

Nit

roge

n

P1

06

P1

07

T10

5

P1

05

S1

01

S1

03

Wa

ter

to

WW

TP

E1

02

E1

01

V1

01

Spe

nt d

ryin

g g

as

DO

IP c

ryst

als

Spe

nt w

ash

Was

te w

ater

to

WW

TP

Eth

yl A

ceta

te

pur

ge

BA

SE

CA

SE

: D

OW

NS

TR

EA

M P

RO

CE

SS

ING

S20

1

EA

m

ediu

m

T201

EA

fres

h

Aq

ueou

s so

lutio

n fr

om V

102

P20

1

Was

hing

wat

er

V2

01

EA

lo

w

V20

2

V20

3

P20

2

V2

04V

205

P20

4P

203

S202

Nit

roge

n

EA

cl

ean

G

K2

01E

201

V2

06

P20

5

P20

6

Spe

nt d

ryin

g g

as

DO

IP c

ryst

als

Spe

nt w

ash

Was

te w

ater

to

WW

TP

Eth

yl A

ceta

te

pur

ge

BA

SE

CA

SE

: D

OW

NS

TR

EA

M P

RO

CE

SS

ING

S20

1

EA

m

ediu

m

T201

EA

fres

h

Aq

ueou

s so

lutio

n fr

om V

102

P20

1

Was

hing

wat

er

V2

01V

201

EA

lo

w

V20

2V

202

V20

3V

203

P20

2

V2

04V

205

P20

4P

203

S202

Nit

roge

n

EA

cl

ean

GG

K2

01E

201

V2

06V

206

P20

5

P20

6

Pro

cess

Eq

uip

men

t S

um

mar

y fo

r IS

PC

C

ase

E101

: R

ecyc

ling

R10

1: F

erm

ento

rE1

02:

Con

dens

orS1

01:

Ultr

afilt

er

P101

: Fe

ed O

IPS1

02:

Cry

stal

lizer

P1

02:

Feed

Glu

cose

S103

: N

utsc

hefil

ter

P103

: Bi

omas

s su

sp.

V101

: C

onde

nsor

P104

: Fe

ed B

ase

P105

: Fe

ed m

isc.

P106

: Pr

oduc

t re

mov

alP1

07:

Rec

ycle

+

empt

yP1

08:

Slur

ry t

o S1

03

Pro

cess

Eq

uip

men

t S

um

mar

y fo

r B

ase

Cas

e

E101

: H

eatin

gR

101:

Bub

ble

colu

mn

E102

: C

onde

nsor

E103

: C

oolin

gS1

01:

Dea

d En

d Fi

lter

E201

: C

onde

nsor

S201

: M

ulti-

Purp

ose

V.

S202

: Fi

lter

K10

1: V

acuu

m p

ump

V101

: C

onde

nsor

P101

: Fe

ed b

ase

V102

: St

orag

e P1

02:

Filli

ng R

101

V201

: EA

med

ium

P103

: Em

pty

R10

1V2

02:

EA lo

wP1

04:

Empt

y V1

02V2

03:

EA c

lean

P2

01:

Fill

S201

V204

: St

orag

eP2

02:

Empt

y S2

01

V205

: St

orag

e P2

03:

Empt

y V2

04V2

06:

Con

dens

orP2

04:

Empt

y V2

05P2

05:

Empt

y V2

06

P206

: Em

pty

S202

BA

SE

CA

SE

: F

ER

ME

NT

AT

ION

S101

Pur

ified

air

V10

2R

101

P10

3P

104

Aq

ueou

s so

lutio

n to

V20

4

Spen

t ai

rV

101

P10

2F

erm

enta

tion

sol

utio

n

T10

3

Wat

er

T10

4

Rea

ctiv

atio

n so

luti

on

T10

5

Aqu

eous

sod

ium

hy

drox

ide

T10

1

P10

1Bub

ble

Col

um

n R

eact

orB

uffe

r V

esse

l

Spen

t re

acti

vat

ion

solu

tion

A

BA

SE

CA

SE

: F

ER

ME

NT

AT

ION

S101

Pur

ified

air

V10

2V

102

R1

01R

101

P10

3P

103

P10

4P

104

Aq

ueou

s so

lutio

n to

V20

4

Spen

t ai

rV

101

V1

01

P10

2P

102

Fer

men

tati

on s

olu

tion

T10

3

Wat

er

T10

4T

104

Rea

ctiv

atio

n so

luti

on

T10

5T

105

Aqu

eous

sod

ium

hy

drox

ide

T10

1

P10

1P

101B

ubbl

e C

olu

mn

Rea

ctor

Buf

fer

Ves

sel

Spen

t re

acti

vat

ion

solu

tion

Spen

t re

acti

vat

ion

solu

tion

A

Wa

ter

to W

WT

P

Sp

ent

nitr

oge

n

Su

ppo

rtC

he

mic

als

T10

1

P1

01O

IP

T10

2

P1

02G

luco

se

T10

3

P1

03B

iom

ass

sus

pen

sion

Fil

tere

d

Air

T10

4

P1

04B

ase

T10

6S

eed

cry

sta

ls

Sp

ent

Air

R1

01

DO

IP c

rys

tals

P1

08

S1

02

Was

hin

g w

ate

r

Nit

roge

n

P1

06

P1

07

T10

5

P1

05

S1

01

S1

03

Wa

ter

to

WW

TP

E1

02

E1

01

V1

01

Wa

ter

to W

WT

P

Sp

ent

nitr

oge

n

Su

ppo

rtC

he

mic

als

T10

1

P1

01O

IP

T10

2

P1

02G

luco

se

T10

3

P1

03B

iom

ass

sus

pen

sion

Fil

tere

d

Air

T10

4

P1

04B

ase

T10

6S

eed

cry

sta

ls

Sp

ent

Air

R1

01

DO

IP c

rys

tals

P1

08

S1

02

Was

hin

g w

ate

r

Nit

roge

n

P1

06

P1

07

T10

5

P1

05

S1

01S

101

S1

03

Wa

ter

to

WW

TP

E1

02

E1

01

V1

01

29



OIP 1.24 t/t

Glucose 3.8 t/t

Water 200.8 t/t

Air 28.3 t/t

Solvent 18.4 t/t

Reactor

P = 1 bar

pH = 4.5Yield = 85.5%

Extractor

P = 1 bar

Yield = 99.5%

Crystallizer

P = 0.03

- 1 barYield = 92.5%

Reaction Separation Product recovery

Solvent 30.8 t/t

Air 30 t/t

31.1 t/t

To waste

To waste 0.19 t/t

DOIP crystals1.0 t/t204.1 t/t

222.2 t/t

Filter

T = 20oCP = 1 bar

Yield = 99.9%

6.33 t/t

Filtrate 6 t/t

Washing water 0.86 t/t

Purification

Total in 253.4 t/t Total out 253.4 t/t

Reactor

P = 1 bar

pH = 4.5Yield = 85.5%

Extractor

P = 1 bar

Yield = 99.5%

Crystallizer

P = 0.03



To waste

To waste

DOIP crystalsFilter

Purification

T = 5oCT = 20oCT = 30oC

OIP 1.24 t/t

Glucose 3.8 t/t

Water 200.8 t/t

Air 28.3 t/t

Solvent 18.4 t/t

Reactor

P = 1 bar

pH = 4.5Yield = 85.5%

Extractor

P = 1 bar

Yield = 99.5%

Crystallizer

P = 0.03



Solvent 30.8 t/t

Air 30 t/t

31.1 t/t

To waste

To waste 0.19 t/t

DOIP crystals1.0 t/t204.1 t/t

222.2 t/t

Filter

T = 20oCP = 1 bar

Yield = 99.9%

6.33 t/t

Filtrate 6 t/t

Washing water 0.86 t/t

Purification

Total in 253.4 t/t Total out 253.4 t/t

Reactor

P = 1 bar

pH = 4.5Yield = 85.5%

Extractor

P = 1 bar

Yield = 99.5%

Crystallizer

P = 0.03



To waste

To waste

DOIP crystalsFilter

Purification

T = 5oCT = 20oCT = 30oC

OIP 1.08 t/t

Glucose 4.32 t/t

Water 13.4 t/t

Air 53.7 t/t

Water 1.03 t/t

Biomass 0.32 t/t

Nitrogen 1.4 t/t

Reactor

T = 30oC

P = 1 bar

pH = 5.5

Yield = 97.5%

Crystallizer

T = 5oC

P= 1 bar

Yield = 93.6%

Filter

T = 5oC

P = 1 bar

Yield = 99.9%

Reaction Product recovery PurificationAir 55.5 t/t

Nitrogen 1.4 t/t

DOIP crystals (1 t/t)6.1 t/t

To Waste 11.2 t/t

Total in 75.2 t/t

6.1 t/t

Total out 75.2 t/t

Reactor

T = 30oC

P = 1 bar

pH = 5.5

Yield = 97.5%

Crystallizer

T = 5oC

P= 1 bar

Yield = 93.6%

Filter

T = 5oC

P = 1 bar

Yield = 99.9%

Reaction Product recovery Purification

To Waste 6.1 t/t

OIP 1.08 t/t

Glucose 4.32 t/t

Water 13.4 t/t

Air 53.7 t/t

Water 1.03 t/t

Biomass 0.32 t/t

Nitrogen 1.4 t/t

Reactor

T = 30oC

P = 1 bar

pH = 5.5

Yield = 97.5%

Crystallizer

T = 5oC

P= 1 bar

Yield = 93.6%

Filter

T = 5oC

P = 1 bar

Yield = 99.9%

Reaction Product recovery PurificationAir 55.5 t/t

Nitrogen 1.4 t/t

DOIP crystals (1 t/t)6.1 t/t

To Waste 11.2 t/t

Total in 75.2 t/t

6.1 t/t

Total out 75.2 t/t

Reactor

T = 30oC

P = 1 bar

pH = 5.5

Yield = 97.5%

Crystallizer

T = 5oC

P= 1 bar

Yield = 93.6%

Filter

T = 5oC

P = 1 bar

Yield = 99.9%

Reaction Product recovery Purification

To Waste 6.1 t/t

In situ product recovery (ISPR) by crystallization: basic principles, design, and potential...

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Transcript of In situ product recovery (ISPR) by crystallization: basic principles, design, and potential...