Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology

CHAPTER 4

Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology L. Peeva, A. Livingston Department of Chemical Engineering, Imperial College, South Kensington Campus, London, UK

Abbreviations API API-INT CVD DMAP DMF EtTS FFA GTI MI MW NF OSN RBO Roxi TOABr

active pharmaceutical ingredient active pharmaceutical ingredient intermediate constant volume diafiltration 4-dimethylaminopyridine N,N-dimethylformamide ethyl tosylate free fatty acids genotoxic impurity mass intensity molecular weight nanofiltration organic solvent nanofiltration rice bran oil Roxithromycin tetraoctylammonium bromide

List of Symbols C Cin Cout CR,i CRio Fin Fsolvent i N Reji

concentration (could be expressed in any units for concentration g L
1; mol L
1 etc.) product concentration in the feed stream (g L
1) product concentration in the outlet (product) stream (g L
1) retentate concentration of component i at any given time (any units for concentration) initial concentration of component i (any units for concentration) product feed stream flow rate (L h
1) swap solvent stream flow rate (L h
1) component i number of solvent volumes per volume of feed used to purify the mixture (
) i component rejection by the membrane defined as 1
Cpermeate,i/Cretentate,i (
)

Current Trends and Future Developments on (Bio-) Membranes. https://doi.org/10.1016/B978-0-12-813606-5.00004-X # 2019 Elsevier Inc. All rights reserved.

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1 Introduction Separation processes account for up to 40%–70% of both capital and operating costs (Marchetti et al., 2014; Adler et al., 2000) in many industries and, in particular, in the fine chemical, pharmaceutical, and biopharmaceutical industries, where efficient and economical purification is a continuing challenge. Membrane separations have been investigated already for many years as alternative to distillation, evaporation, adsorption, extraction, and chromatography, with the main hurdle facing successful application being the organic solvent environment in most chemical and pharmaceutical production. Over the last two decades, Organic Solvent Nanofiltration (OSN) emerged as a new technology paradigm for molecular separations in organic solvents, and suitable solvent resistant membranes were developed. OSN has been shown to be a feasible alternative to, and has been combined with, existing separation processes, such as distillation, evaporation, adsorption, extraction, and chromatography (Marchetti et al., 2014; Vandezande et al., 2008). OSN operations have been classified into three conceptual operating modes: concentration, solvent exchange, and purification (see Fig. 1) (Marchetti et al., 2014). Concentration processes recover high-value solutes from a dilute solution (solute enrichment) or recover solvents by removing dissolved impurities (solvent recovery). Solvent exchange transforms a solution of a solute in solvent A to a solution of a solute in solvent B. Both concentration and solvent exchange processes require tight membranes, which retain the solute while permeating solvent(s). In purification processes, the goal is to separate two or more solutes present in a solution, for example, the main product of a reaction from an impurity. Interestingly, the highest number of publications on OSN applications is dedicated to purification applications (see Fig. 1D), where the selectivity between the two solutes is the key factor determining process feasibility (Marchetti et al., 2017). Comprehensive reviews of each type of the above applications in the pharmaceutical and biopharmaceutical industries have been presented in our earlier works (Marchetti et al., 2014, 2017; Peeva et al., 2017; Szekely et al., 2014). In this chapter, we will focus our attention on more specific cases of applications of multistage processes.

2 Multistage Membrane Processes with Potential to Broaden Applications in Industry Despite the rapidly growing interest in OSN, the number of industrial applications is still rather limited. Apart from the MAX-DEWAX process at ExxonMobil Corporation 1998 which has now been discontinued (Gould et al., 2001; White and Nitsch, 2000), a more recent industrial OSN application developed by Evonik-MET is an API recovery concentrating a 1 wt% API waste stream to 10 wt%, which is then fed to existing downstream processing units for further purification. The API has a molecular weight of 420 g mol
1 and is highly sensitive to higher temperatures as it rapidly decomposes at temperatures above ambient,

Fig. 1 Membrane filtration processes for liquid applications: (A) concentration, (B) solvent exchange, and (C) purification processes. (D) OSN publications for each application, over the period 1990–2016. Adapted from Marchetti, P., Peeva, L., Livingston, A., 2017. The selectivity challenge in organic solvent Nanofiltration: membrane and process solutions. Annu. Rev. Chem. Biomol. Eng. 8, 473–497.

100 Chapter 4 and therefore, evaporation or distillation of this stream is not possible. The unit is small in size with a membrane area of only 15 m2(Dura-Mem 300 as 4”  40” spiral wound modules), but generates a profit of 1 million euros per year (<1 year payback period) (Priske et al., 2016). From this example of the process, economic and environmental impact of the membrane technology is obvious; however, several challenges still remain for the full potential of this technology to be achieved. One of the major hurdles faced by membrane technology is the incomplete separation resulting in high product losses and large solvent consumption (Kim et al., 2014b). In spite of extensive research in membrane development and various proposed membrane modifications, the separation performance of membranes hasn’t improved much over the years. The limited separation performance may largely be due to the transport mechanism of the separated species through a membrane. Although this mechanism is still not fully elucidated and understood, it is widely accepted that, for most membranes, it is governed by the so-called pore flow transport (Bowen and Welfoot, 2002b, 2002a), which is somewhat a prerequisite for imperfect separation. Predictions according to the pore flow model of membrane rejection profiles for uniform pore size membranes in the nanofiltration range are presented in Fig. 2. An “ideal” rejection curve is also shown for comparison. It can be seen that even with perfectly controlled, uniform pore size, it is difficult to get a clear separation of, for example, a product (P) from impurity (I) even though there is significant

Fig. 2 Simulated membrane rejection plotted against molecular size and weight for membranes with uniform pore sizes (rp). Even with perfectly uniform pore size of 1.5 nm, it is difficult to separate the product from impurity efficiently. Adapted from Kim, J.F., Freitas Da Silva, A.M., Valtcheva, I.B., Livingston, A.G., 2013. When the membrane is not enough: a simplified membrane cascade using organic solvent Nanofiltration (OSN). Sep. Purif. Technol. 116, 277–286.

Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology 101 difference in their molecular weights. As the pore size of the membrane increases, the rejection curves do not converge toward the ideal separation curve, but “smear” toward higher molecular weights due to hindered transport of molecules “traveling” into liquid-filled pores of the same size as molecular dimensions (Deen, 1987; Kim et al., 2014b). In order to illustrate how important the difference in rejection (retention) between two species is, let’s look at an example where two solutes, A (product) and B (impurity), have to be separated via a very popular membrane separation process called constant volume diafiltration (CVD). A schematic representation of CVD is shown in Fig. 3A, and more details

100

Yieldi (%)

80

Pure solvent (F)

60

1 2

40

3

20

F=P B A AB BA A

B

(A)

10

0

20

40 60 Rejectioni (%)

(B)

Membrane

Permeate (P)

80

100

10 5 4 3 2

100 80

B B B

5

0

RejA>RejB

PurityA (%)

A

4

1

60 40 20 0 –100

(C)

–50

0 RejA–RejB (%)

50

100

Fig. 3 (A) schematic representation of constant volume diafiltration process for separating solute A from solute B; (B) Effect of rejection of solute i on the yield of the same solute; (C) Effect of rejection difference on the purity of solute A at different diafiltration volumes (numbers on top of the lines indicate diafiltration volumes). Adapted from Marchetti, P., Peeva, L., Livingston, A., 2017. The selectivity challenge in organic solvent Nanofiltration: membrane and process solutions. Annu. Rev. Chem. Biomol. Eng. 8, 473–497.

102 Chapter 4 are given in the next section. For the yield of any solute subjected to a CVD process, the following equation could be derived: Yield ¼

CR, i ¼ e
Nð1
Reji Þ CR, i0

(1)

where CR,i is the retentate concentration of component i at any given time, CR,io is the initial concentration of i, N is the number of solvent volumes per volume of feed used to purify the mixture, and Reji is the i component rejection by the membrane defined as 1
Cpermeate,i/Cretentate,i. A typical yield curve as a function of component rejection is presented in Fig. 3B, for different number of solvent volumes N. As could be expected, the higher the rejection, the higher the yield is; however, in CVD configuration, even a marginal deviation from 100% in rejection could result in considerable product losses. Now let’s look at the possibility for separating solute A from solute B using CVD. From Eq. (1) for solutes A and B for a starting solution with equal concentrations (CA0 ¼ CB0), it is easy to derive an equation describing purity of A as a function of rejection differences (Eq. 2). Note yield, rejection, and purity could be expressed in % by multiplying by 100. PurityA ¼

CA 1 ¼
N ðRejA
RejB Þ CA + CB 1 + e

(2)

As could be seen from Fig. 3C, purity >99% could be obtained after 10 solvent volumes only if the rejection difference is 50%; however, for the same solvent volumes, rejection of product >99% is required for higher than 90% yield. These requirements severely constrain the applicability of a simple CVD process. From the above example, it is obvious that achieving perfect separation with a single membrane stage would be extremely difficult (if not impossible) for purely physicochemical reasons. Another approach to overcome this problem is to apply an engineering solution by implementing a multistage process, known as a membrane cascade. This approach is not new and various membrane cascade configurations have been proposed in the literature (predominantly for aqueous applications) and analyzed, mainly theoretically and/or by extrapolating from a single-stage results (Lightfoot, 2005; McCandless, 1990; Vorotyntsev et al., 2001; Lightfoot et al., 2008; Vanneste et al., 2011; Maskan et al., 2000; Avgidou et al., 2004; Keurentjes et al., 1992; Ghosh, 2003; Noda and Gryte, 1981; Caus et al., 2009; Abejo´n et al., 2012). Membrane cascades have even been considered as an alternative process to chromatography (Lightfoot, 2005). In the area of solvent nanofiltration, there are still a very limited number of publications on implementation of membrane cascades. Probably, one of the first analyses on the potential of membrane cascade in solvent nanofiltration (Kale et al., 1999) investigated a process for

Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology 103 deacidifying rice bran oil by solvent extraction and membrane technology. Based on the laboratory results from a dead end filtration cell, the authors presented a preliminary design of a system to recover FFA from methanol extracts by multistage NF for a plant capacity of 10,000 kg h
1 of crude rice bran oil. The design goal was the production of a retentate stream containing 20% FFA that would go to the methanol evaporator and a permeate stream with a low level of FFA that could be directly recycled to the extractor. They evaluated case studies for a single membrane stage, two, and three membrane stages, where permeate from each stage is fed into the next stage. Their estimations suggested that the capital cost of a one-stage membrane plant is about $30 per kg h
1 of oil processed, while a three-stage plant will cost $64 per kg h
1 of crude oil. Operating cost also increases from $9.8 to $19.7 per ton of FFA recovered, depending on the number of stages. However, FFA recovery increases from 93% with one stage to 99% with three stages, and thus the value of the FFA recovered compensates for the additional stages. With a value of $56 per $ of operating cost gained for a single stage, $36.9 per $ of operating cost and $28 per $ of operating cost for two and three stages, respectively, the membrane cascade still remains an attractive process.

3 Membrane Cascades Operating in Batch Diafiltration Mode There are various examples of how the membrane cascade can improve the separation/ fractionation of species in the pharmaceutical industry. Most studies utilize a cascade configuration operating in diafiltration mode. In such a mode, a high MW product retained by the membrane is gradually purified by washes with a pure solvent which carries away low MW impurities that are able to permeate through the membrane (see Figs. 1C and 3A). The purity gradually increases with the permeate volume, but so do the product losses through the membrane due to imperfect separation (product is not 100% retained by the membrane). Also, due to the imperfect separation, the impurities retention is also relatively high (see Figs. 2 and 3B,C), resulting in excessive solvent usage. Thus, the membrane cascade becomes essential, in order to minimize solvent usage and product losses. Alternatively, the purification is also possible when a high MW impurity is retained by the membrane while the low MW product carried by the solvent permeates through, but this case study suffers from similar disadvantages as listed above. One of the first works using the concept of a membrane cascade (Sereewatthanawut et al., 2010) investigated the application of solvent nanofiltration for active pharmaceutical ingredient purification in diafiltration mode. The authors investigated the purification of a macrocyclic intermediate of a new drug candidate at Janssen Pharmaceutica (API-INT, MW 675 g mol
1), from a series of oligomeric impurities based on API-INT with MW > 1000 g mol
1 (i.e., dimers, trimers, tetramers, pentamers, etc.). The main objective of the case study was to remove >99% of the tetramers and higher oligomeric impurities from

104 Chapter 4 API-INT, while recovering >95.0% of API-INT. Due to the high rejection of API-INT >60% by the membranes under study, an excessive amount of pure solvent was required to achieve the target separation and yield. In order to solve this problem, they introduced a second membrane stage utilizing a tighter membrane. This second stage was able to retain and concentrate the API, while the recovered solvent was reused in the first stage to carry on the diafiltration, thus massively reducing fresh solvent consumption. The process was dubbed “dual mode diafiltration” and was able to reduce the overall content of oligomeric impurities in an organic solution synthesized at Janssen Pharmaceutica from 6.8% to 2.4%, which is below the target limit of 3.0%, as well as removing 99% of the particularly challenging higher oligomeric impurities (i.e., tetramer and higher of API-INT) with an API yield of 99.2%, thus proving highly superior to crystallization and charcoal treatment. Although further chromatographic separation was still required, this subsequently made the chromatography process more efficient, in a successful synergy of OSN and chromatography. With dual membrane diafiltration, solvent use was minimized, i.e., reduced 10 times from the original consumption. After initial proof of concept, the OSN process was introduced at pilot scale. Later on, the same author proposed a process for refining and γ-oryzanol enrichment of rice bran oil (RBO) in a diafiltration membrane cascade (Sereewatthanawut et al., 2011). The crude RBO is typically produced by extraction of crude rice bran with hexane. From the extraction process, the crude RBO contains approximately 1 wt% γ-oryzanol (MW 603 g mol
1). The remaining mass of oil consists of glycerides (>90 wt%) and FFA (4–5 wt% (determined by titration); MW 280 g mol
1). The majority of glycerides found in most edible oils are triglycerides (MW 800–900 g mol1) and a small fraction of mono- and diglycerides are present (MW 300 and 500–600 g mol
1, respectively). The proposed multistage OSN process consists of (i) glycerides separation from γ-oryzanol in the first cascade stage, and (ii) FFA removal in the second stage, with a potential solvent recovery as a third stage (to minimize the large solvent usage). The process demonstration results showed that enrichment of γ-oryzanol from 0.95% in the feed oil to 4.1% in the product oil could be achieved. The RBO could be refined to acceptable levels (FFA < 0.2 wt%) with minimal γ-oryzanol losses, resulting in a twofold increase in the oil antioxidant capacity of the RBO. However, this work is more conceptual since actual cascade operation has not been demonstrated and the reported results are obtained in separate membrane units. Another study (Vanneste et al., 2013) investigated the possibility of using membrane cascades to carry out difficult pharmaceutical separations. It targeted the separation of an intermediate 1-(2-bromoethyl)-4-ethyl-1,4-dihydro-5H-tetrazol-5-one (MW 221 g mol
1) from an impurity, ethylene bromide (MW 188 g mol
1). The molecular weight of the intermediate is only 17.6% higher than the molecular weight of the impurity and the purity of the starting mixture is low (26%) and has to be increased to 90%. The results from the cascade modeling (based on flux and rejection data obtained in a cross-flow unit) showed that the product yield in the retentate increased from 35.5% to 84.3% by adding two stages;

Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology 105 however, the increase in yield from two to three stages was three times lower, suggesting an optimum number of stages exists. An extensive cost analysis showed that not only the product yield (product losses) is important for the process, but also the number of stages and the module size (determining the process time) and both have to be optimized. Also, it is desirable that multiple separations are carried out with the same setup to reduce the cost of separation. For the selected case study, it appeared that a two-stage cascade is the best option; however, no actual cascade experiments were presented. One of the main hurdles toward membrane cascade implementation is automatization and control and the usage of multiple high pressure pumps and reservoirs. As has been pointed out by Lightfoot, even very simple counterflow cascades have not been widely used due to control problems and lack of operating experience (Lightfoot, 2005). Recently, Kim et al. (2013) a simplified-control cascade process is demonstrated that operates using a single high pressure pump as the primary pressure source and has no need for a buffer tank between membrane stages. The cascade performance has been validated in diafiltration mode on purification of active pharmaceutical ingredient (API, macrolide antibiotic Roxithromycin (Roxi)) from two model genotoxic impurities of different chemical classes (GTI, 4-dimethylaminopyridine (DMAP) and ethyl tosylate (EtTS)) (Kim et al., 2014b). By implementing a two-stage cascade configuration, the process yield was increased from 58% (for a single-stage diafiltration) to 95% (for two-stage diafiltration), while maintaining <5 ppm GTI in the final solution. A predicted yield improvement between single- and dual-stage diafiltrations as a function of product rejection is presented in Fig. 4. Due to the relatively high rejection of the GTIs, large solvent volumes were required to achieve desired API purity. A green metric analysis suggested that, if this solvent is disposed, the mass intensity of the process is inacceptable high (MI > 1200). A solvent recovery step has been assessed using charcoal as a nonselective adsorbent, and it has been shown that pure solvent can be recovered from the permeate and reused in the diafiltration. Considering the costs of solvent, charcoal, and waste disposal, it was concluded that 70% solvent recovery is the cost-optimum point. Depending on the batch scale, the dual-stage cascade with solvent recovery can achieve up to 92% cost saving, while reducing the mass and solvent intensity up to 73% with energy savings of up to 96% compared to distillation and a 70% reduction of CO2 footprint. In a further study by the same author, the solvent recovery via activated charcoal adsorption was replaced by solvent recovery in two-stage membrane unit (Kim et al., 2014a), where the second membrane stage was equipped with tighter membrane capable of retaining the impurity. Again Roxithromycin was used as an example of API, while triphenylmethanol as exemplary impurity. The closed loop process was able to operate without any solvent addition, achieving 99% yield of the API and 97% impurity removal. It has to be pointed out that the applicability of this process is limited by the retention of the impurity by the membrane. For very small MW impurities (typically poorly retained by the membrane),

Two stage cascade Single stage

Adsorbent n

Stirring Plate

(Vi)

(CP1)

Solvent

Permeate 1

(F1)

Permeate 2

(F2)

P Membrane Cell 1

(Fi)

P

Membrane Cell 2

(V2)

Feed

(F3)

Stirring Plate

Stirring Plate Feed Pump Retentate 1

(CRi)

Back Pressure Regulator - 1 Retentate 2

(CR2)

Back Pressure Regulator - 2

(A) 100

Single stage diafiltration Two stage diafiltration

80

Yield (%)

60

40

20

0 85

(B)

86

87

88

89

90

91 92 93 94 Product rejection (%)

95

96

97

98

99

100

Fig. 4 (A) Schematic representation of a two-stage membrane cascade with recycle of retentate and an adsorption solvent recovery unit. (B) Predicted yield improvement after 10 diavolumes for different product rejection. The dual stage configuration achieves significant yield improvement, overcoming the membrane limitation due to imperfect separation (rejection less than 99%). As long as the product rejection is above 90%, the process promises high product yield, making the process competitive to other traditional unit operations. Adapted from Kim, J.F., Sz
ekely, G., Valtcheva, I.B., Livingston, A.G., 2014b. Increasing the sustainability of membrane processes through cascade approach and solvent recovery - pharmaceutical purification case study. Green Chem. 16, 133–145 with permission from The Royal Society of Chemistry.

Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology 107 efficient solvent purification is not possible. Nevertheless, implementing a solvent recovery unit in diafiltration process generally brings significant CO2 footprint reduction, mainly from avoiding solvent incineration. In addition, membranes generate less solid waste than adsorption where adsorbent waste can become a significant factor. Further improvement of the solvent recovery concept in the diafiltration mode membrane cascade has been proposed (Schaepertoens et al., 2016). They developed a three-stage membrane cascade diafiltration process comprised of two membrane stages for separation with a third membrane stage added for integrated solvent recovery and recycle (closed loop configuration). They targeted the model separation of catechol (Mw 110.11 g mol
1, reagent of the classic synthesis of dibenzo-18-crown-6) from the product crown ether (Mw 360.40 g mol
1). Similarly to Kim’s work (Kim et al., 2014b), the first two stages allow for increased separation selectivity and higher yield, while the integrated solvent recovery stage mitigates the otherwise large solvent consumption of the diafiltration process. Because of the imperfect separation in the solvent recovery stage (impurity is retained <100%), the impurity is removed in a semicontinuous fashion when it reaches certain concentration in the solvent recovery stage. This is achieved via washing the solvent recovery unit at intervals in order to attain high product purities. The suggested process achieved a purity of 98.7% and 98.2% yield through semicontinuous operation with two washes of the solvent recovery stage and consumed 85% less solvent as compared to a system without solvent recovery stage with similar yield and purity. On the negative side, the duration of the process with solvent recovery is nearly double as compared to the process without solvent recovery stage for achieving the same purity.

4 Membrane Cascades Operating in Continuous Mode All examples reported above are describing batch diafiltration processes. However, many pharmaceutical manufacturers are actively investigating the benefits of converting their processes to continuous production. Companies implementing continuous manufacturing are anticipating a fast return on their investment and cost savings of 10%–20% as compared to batch manufacturing (www.siemens.com/pharma), reduced energy and carbon footprints, and improved overall safety (Gutmann et al., 2015). Membrane unit operations are well-suited for continuous processes due to their ease of operation in flow, scalability, and the absence of phase transitions or biphasic systems. Thus, several research groups have focused their investigations on implementation of membrane cascades in continuous operations. A continuous solute fractionating membrane cascade configuration has been proposed (Siew et al., 2013a). A McCabe–Thiele approach was used to develop the process models for the membrane cascades. Schematic representation of the proposed configuration is shown in Fig. 5. The cascade could be divided into two sections. The feed stream, supplied into the feed stage, is channeled progressively through the retentate chambers of a series of

108 Chapter 4 Purifying section

Feed Stripping section

Stripping solvent

Fig. 5 Schematic of continuous solute fractionation membrane cascade. Adapted from Siew, W.E., Livingston, A.G., Ates, C., Merschaert, A., 2013a. Continuous solute fractionation with membrane cascades—a high productivity alternative to diafiltration. Sep. Purif. Technol. 102, 1–14.

membrane units at the right side of the feed, stripping section of the cascade, where it is stripped of the more permeable solute with a counter-current flow of pure solvent (stripping solvent). The stripping section allows for maximal utilization of the stripping solvent stream as it is reused over a number of stages. A stream enriched with the less permeable solute exits as the retentate of the first stage (product-enriched stream). The depleted solvent stream exiting as the permeate of the feed stage is then sent through a purifying section, which further retains the less permeable species relative to the more permeable solute. This is achieved through a series of membrane stages and appropriate recycling of the permeate stream from the final stage back into the purifying section. The fractionating cascade can be simplified when the rejection of one species approaches unity, since the use of a purifying section is unnecessary as very little of the less permeable species will be lost from the stripping section. The most apparent effect of using the cascade instead of batch diafiltration is the considerable decrease (about 30% decrease for three-stage cascade) in solvent usage. Mathematical simulations showed that an increase in the number of stages decreases the amount of stripping solvent required for the purification; however, beyond four stages, this reduction becomes marginal. A three-stage stripping membrane cascade operation was verified for the separation of a developmental API from an excess reagent. A binary feed solution, containing the API and the excess reagent, was stripped of the reagent to produce a product stream enriched in the API that could be further polished using a single crystallization step, which was otherwise impossible with the original feed solution. As mentioned before, control over the flow rates in the cascade proved to be challenging. Further investigation of membrane cascades for simultaneous API enrichment and organic solvent recovery was performed (Siew et al., 2013b). Again, API enrichment/concentration was performed into the striping section of the cascade, while solvent purification took place in the purifying section. The stripping solvent stream was removed. The process was evaluated theoretically and demonstrated in a three-stage membrane cascade where the feed was fed into the middle stage. The target was concentration of a dilute API product solution and solvent recovery downstream of a chromatographic process. An API, from

Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology 109 UCB Pharma S.A., was used as the solute, while solvent mixture recovery was a mixture of 90% ethyl acetate and 10% methanol. The three-stage cascade achieved an effective rejection of 80% compared to a single pass rejection of 55%, and the concentration of API in the product stream increased 2 times compared to the feed stream. Due to the limited number of stages and relatively low API rejection, proper purification of the solvent was not possible. However, theoretical analysis suggests that this could be considerably improved if membranes with better rejections could be used. Recent work (Peeva et al., 2014) demonstrated continuous purification of API (Roxithromycin) from potential GTI (DMAP) in a simple and efficient two-stage membrane cascade. The cascade configuration was an improved version of the one proposed by Kim et al. (2013); however, some modifications were made in order to improve mass transfer in the membrane cells and increase flexibility of the cascade. The magnetic bar stirring (see Fig. 4A) was replaced by recirculation of the fluid through the cells via a powerful gear pump. This modification not only improves the mixing, but provides flexibility in the number of cells that could be operated at each stage. Cascade performance was initially evaluated via mathematical simulations, and then validated experimentally. The authors demonstrated that by careful selection of operating parameters, high purity of the API >99% could be achieved from feed stream of almost any initial purity (experimental results are shown for initial API purity of 78% and 55%). Reasonable purification of heavily contaminated production streams (initial purity <20%) is still possible, but in expense of a low yield. Purity and yield in the cascade configuration (97.9% and 97% respectively) were considerably higher than a single-stage unit (96.9% and 68.4%) under the same operating conditions. The continuous cascade could be easily coupled with an adsorption unit (e.g., two alternating fixed bed columns utilizing an inexpensive nonselective adsorbent such as charcoal and operated in a cyclic adsorption-regeneration batch mode) or another continuous purification step (e.g., extraction, evaporation) in order to purify the solvent and recycle it back to the process. The continuous process has a superior MI index (MI21) when compared to a batch process with 100% solvent recycle via adsorption (MI200) and is within the lowest range for pharmaceutical production process where the lowest reported values are of 23 (Fig. 6). Interesting studies (Abejo´n et al., 2014, 2015) evaluated multiple configurations of continuous organic solvent nanofiltration membrane cascades for the separation of an intermediate precursor (1-(2-bromoethyl)-4-ethyl-1,4-dihydro-5H-tetrazol-5-one) of an anesthetic compound (alfentanil) from an impurity (ethylene bromide). This is a very challenging separation due to the similar size of both solutes (MW product 221 g mol
1 and MW impurity 188 g mol
1). All data on membrane performance were taken from the literature and the analysis is only theoretical. Extensive cost analysis suggests that the main cost of the process is the treatment of the residual stream leaving the system (>85% for some cascades). They concluded that this process can only be considered as a viable option when the solvent

110 Chapter 4 1600 Batch purification

Continuous purification

1400

Mass intensity (–)

1200 1000 800 600 400 200 0 Single stage batch solvent discharge

Two stages batch solvent discharge

Two stages batch 100% solvent recovery

Two stages continuous solvent discharge

Two stages continuous 100% solvent recovery

Fig. 6 Comparison of the MI index for batch and continuous purification of API—Roxithromycin using a single stage batch diafiltration, two stages batch diafiltration and two stages continuous membrane cascade. Solvent recovery is performed via adsorption with activated charcoal. Adapted from Peeva, L., Burgal, J.D.S., Valtcheva, I., Livingston, A.G., 2014. Continuous purification of active pharmaceutical ingredients using multistage organic solvent nanofiltration membrane cascade. Chem. Eng. Sci. 116, 183–194.

recovery from the residual stream leaving the system is incorporated. By incorporating additional solvent recovery stages in the cascade configuration, the total cost of the process could be reduced considerably, by up to 64%–77% depending on the required solvent quality. A four-step design method for combination of OSN and distillation in a hybrid separation of wide boiling mixtures has been suggested (Micovic et al., 2014). This design method was applied to investigate separation of heavy boiler (hexacosane, 5%) from low and middle boiler (decane, 70% and dodecanal 25%) in a wide boiling mixture from hydroformylation. In step 1, different process alternatives were generated, which include stand-alone operations and hybrid separations, and two most promising process options were chosen. In step 2, a process analysis is performed to investigate what values of permeability and permselectivity are needed for the OSN-assisted process to be economically better than other alternatives. The membrane stage design included optimization of a multistage membrane cascade. The process analysis showed that apart from the separation properties of the membrane, other important parameters to be considered are driving force reducing effects (e.g., concentration polarization), as an idealized approach may overestimate the advantages of OSN-assisted process. In step 3, membrane screening is performed and GMT ONF1 was identified as the

Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology 111 most suitable membrane, with which detailed experiments were performed in the operating range determined in process analysis. In the final step 4, further process optimization performed using the experimental results showed that, at high temperatures, the OSN-process may be more economical than stand-alone distillation, under conditions where the membrane material shows enough long-term stability and costs less than125 Euro per m2 per year. Another work (Adi et al., 2016) presented the application of a superstructure optimization approach to determine the best possible configuration of OSN membrane cascades for the separation of organic mixtures. The effectiveness of this approach has been demonstrated experimentally for a binary system of heptane and hexadecane, in a three-stage membrane cascade, thus providing a proof-of-concept for hydrocarbon mixture separation using membrane cascades. Experimental results were within good agreement with the predicted values. Superstructure approaches have also been proposed by other researchers (Schmidt et al., 2014) in a four-step design workflow for OSN processes utilizing conceptual process design tools and OSN membrane cascade-based optimization. In the first step, the minimum separation targets (minimum rejections, selectivities, or permeabilities) are identified. In the second step, promising OSN membranes and solvents are preselected from a database. In step 3, the performance of the promising membranes from step 2 is validated experimentally within the recommended boundaries. In the final step, process modeling and optimization based on detailed economic data are performed. In contrast to other approaches, the design is based on a detailed model of a flexible OSN membrane cascade superstructure which can be used to determine optimized process configurations and process conditions simultaneously. The design tool is validated for a case study of recycling of a homogeneous Rhodiumtriphenylphosphine-based catalyst during hydroformylation. The authors apply interesting and unusual approach in this study by introducing targeted solvent addition to the separation mixture. For example, for Puramem™280 membrane, a large increase in triphenylphosphine rejection from 87% in pure n-hexanal (original reaction solvent) to approximately 98% in a solvent mixture with 50 wt% toluene was observed associated also with an overall flux increase. It was concluded that an integration of solvent additions in the process optimization is very promising. Thus, the process model also incorporates addition of solvents before or during catalyst recycling using OSN as well as final product purification by distillation. It was estimated that the most economic process includes a five-stage OSN membrane cascade without toluene additions (16.76 € t
1 product). The best process with toluene additions (three-stage) has 14% higher production cost than the most economic one, but a two-stage process with toluene addition also has a reasonable price of 22.99 € t
1 product. The results showed that targeted solvent additions can be utilized to enhance process economics and attain processes with fewer stages improving process operability and control and reducing overall membrane area demand.

112 Chapter 4 Another interesting application of membrane cascade has been proposed (Lin and Livingston, 2007). The authors proposed a counter-current membrane cascade for solvent exchange, key unit operation in pharmaceutical production, and one of the major solvent consuming processes. Solvent exchange from low boiling point to high boiling point solvent in flow is relatively easily achievable by distillation; however, a solvent exchange in the opposite direction (reverse boiling point order) is typically difficult and is associated with significant energy consumption and large quantities of intermediate solvent mixtures. Aside from economic effects, thermal operations may not be ideal for APIs and/or catalysts as many of these are thermally labile. Solvent exchange performed via batch diafiltration is possible; however, even if the target product is completely retained by the membrane and is not lost through the permeate, large solvent volumes are required (according to our estimations 7 solvent volumes per volume of product solution to obtain 99.9% solvent swap, assuming the membrane does not discriminate between the two solvents). The proposed multistage cascade was operated in a counter-current mode, with the initial solvent stream containing API and the replacing solvent stream being simultaneously fed into the first and the last stages, respectively (configuration similar to the stripping section in Fig. 5). Two solvents commonly used in organic synthesis, toluene and methanol, were chosen as a case study while a model compound (tetraoctylammonium bromide (TOABr), MW 547 g mol
1) was representing the API. Mathematical modeling was performed to evaluate the effect of number of membrane stages and swap solvent to initial solvent ratio on cascade performance. As could be expected with the increase of number of membrane stages and amount of swap solvent, the solvent exchange is improved; however, the improvement between the three- and four-stage cascades was marginal. The results were validated experimentally for one-, two- and three-stage cascades for swap solvent to initial solvent ratio of 1. The results were in agreement with the theoretical ones and for a three-stage cascade 75% solvent exchange was obtained, while using batch diafiltration after 1 swap solvent volume the solvent is exchanged 63% showing the benefits of introducing membrane cascade. The authors pointed out that complete solvent exchange with the proposed membrane configuration was not possible and further cascade optimization studies are required. Here, we evaluate whether through further process optimization it is possible to improve the above-described solvent exchange process. After selecting an appropriate membrane similar to the one used in Lin’s study, the flux/rejection properties were first determined (99.7% for TOABr by Puramem 280, no significant change was observed related to the solvent composition). The experimental results for flux and rejection were used to evaluate whether it is possible to minimize solvent usage by reconfiguring the membrane cascade. An optimization procedure was performed to select the best membrane cascade configuration allowing for good yield and minimum amount of swap solvent. The effect of number of stages was also evaluated. The optimization procedure was performed using gPROMs dynamic simulator (gOPT) with a target function for minimum solvent volumes usage

Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology 113

Number of solvent volumes (–)

10

8

6

4

2

0 1

2

3 4 Number of stages (–)

5

Fig. 7 Number of solvent volumes per volume of feed as a function of number of cascade stages. (No solvent volumes ¼ Swap solvent stream flow rate/Feed stream flow rate). The solvent is exchanged 99.9%.

(Fsolvent/Fin ¼ minimum). The product feed stream Fin was set at 0.1 L h
1. To obtain experimentally meaningful results, the operating pressures were constrained between 10 and 30 bar the outlet product stream to 0.01 L h
1 and the swap solvent content to 99.9%. Results of solvent usage and product yield (as recovery) are presented in Figs. 7 and 8.

Product recovery (%)

100 90 80 70 Without product recovery stage With product recovery stage

60 50 1

2

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Fig. 8 Product recovery (yield) as a function of number of cascade stages. (Product recovery ¼ (Product stream flow rate  Cout/Feed stream flow rate  Cin)  100). The solvent is exchanged 99.9%.

114 Chapter 4 As can be seen from Fig. 7, the minimum solvent usage 1 is achievable in a four-stage cascade. Further increase of the cascade stages is not beneficial for the process. A similar result is found for the product recovery (yield ¼ mass product out/mass product in) where the maximum recovery of 92.8% again is achieved in the four-stage cascade. Introducing a product recovery stage (see Figs. 8 and 9) increases product recovery to 99.9% in both three- and four-stage cascades. Again, adding more than four stages is not beneficial to the process in terms of product recovery. The final challenge was to actually set up and implement the optimized membrane cascade. The initial cascade configuration used (Fig. 9) was utilizing two mass flow controllers. However, it was still difficult to maintain precise flows with the back pressure regulators and there were deviations from the theoretically optimized flows. For the actual experimental flow rate values obtained, the theoretical estimation suggested that the cascade will reach only 95.5% solvent swap with 99.3% product recovery, better than the results previously obtained in a cascade (Fig. 10A). As can be seen from the figure, the model prediction is reasonably close to the experimental values. In fact, the experimental value for the swap solvent in the product stream was higher than estimated (95.5 vs. 97%), which could be due to an analytical error or slight inaccuracy of the model (small flow/permeance variations with temperature and pressure which the model cannot account for). The TOABr concentration fluctuations in all the three stages were more significant and are also due to flow instability during the operations. The membrane permeance was changing (most likely due to adaptation to the new solvent composition). Several readjustments of the flow rates were performed (after 42 h and 64 h) in an attempt to improve swap solvent content to 99%. Results for the solvents and solute concentrations as well as the product (TOABr) recovery (yield) are presented in Fig. 10 A,B,C. The cascade was operated for an additional 110 h under readjusted conditions (from 64 h onward). In terms of solvent exchange, its performance was excellent achieving swap solvent (methanol) concentration 99% that remained stable during the run. The case was different for the product concentration with significant variations in the cascade stages over time (Fig. 10B) and steady state concentration could not be reached. Consequently, the recovery (Fig. 10C) had significant fluctuations, but 93% as an average (99.8% theoretical). It was noticed that the product concentration being relatively low as compared to the solvents is particularly sensitive to fluctuations of the permeate flow in the recovery stage. In a trial to improve flow stability, two additional mass flow controllers were installed on the permeate line from the recovery stage and the retentate line of stage three (indicated with red arrows and red dotted circles on Fig. 9) and flow rates were readjusted in order to maximize the recovery. The cascade was operated for additional 85 h (data points indicated with a red circle). As could be seen from Fig. 10 again, the solvents’ concentrations were stable overtime in all the three stages and there was an

Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology 115

Fig. 9 Optimised configuration for three stage membrane cascade with product recovery stage. The values in green rectangulars are the average actual operating values during the solvent exchange run from 174 h onward. The places of the new flow controllers are indicated with red arrows and dotted red circles.

116 Chapter 4 10

80 60 Stage 1—model Stage 2—model Stage 3—model Stage 1—experimental Stage 2—experimental Stage 3—experimental

40 20

TOABr concentration (g L–1)

Methanol concentration (%)

100

0 0

50

(A)

100 150 Time (h)

200

8 6 4 2 0

250

0

(B)

50

100 150 Time (h)

200

250

Product recovery (%)

120 100 80 60 40

Experimental Model

20 0 0

(C)

50

100 150 Time (h)

200

250

(D) Fig. 10 Methanol (A) and TOABr (B) concentrations in the different cascade stages over time during the run. (C) Product yield over time during the run. (D) Simulated solvent exchange results in OSN membrane cascades with varying numbers of stages at varying flow rate ratios of replacing solvent to initial solvent (QPn + 1/QF1) (D) Adapted from Lin, J.C.-T., Livingston, A.G., 2007. Nanofiltration membrane cascade for continuous solvent exchange. Chem. Eng. Sci. 62, 2728–2736.

improvement of the product concentration stability. The product recovery was on average 99.9% in 98% swap solvent as it was predicted by the model. This result demonstrates that, by careful control of the cascade streams, stable operation is achievable. It should be noted that, at large scale applications, the operational flow range would be orders of magnitude higher, more accurate flow control would be possible, and the effect of small fluctuations will be diminished. Overall, the new cascade configuration shows a reasonable improvement compared to the previously applied by Lin, where with the same solvent/feed ratio of 2 for a three-stage configuration only 85% solvent swap could be obtained (Fig. 10D).

Nanofiltration in the Pharmaceutical and Biopharmaceutical Technology 117 Additional information about the membrane cascade setup utilized can be found elsewhere (Peeva et al., 2014, 2016). We recently demonstrated the solvent exchange strategy in a multistage membrane process for continuous pharmaceutical drugs manufacturing (Peeva et al., 2016). We presented the first example of continuous consecutive reactions where the catalyst recovery and the solvent exchange are achieved in continuous membrane units. As a case study, two consecutive reaction steps were selected from the synthesis of the API [6-chloro-2(4-chlorobenzoyl)-1H-indol-3-yl]-acetic acid, a selective cyclooxygenase 2 (COX-2) inhibitor developed by Pfizer Global Research. The two steps require a reverse boiling-point-order solvent exchange from N,N-dimethylformamide (DMF) to ethanol. A Heck coupling reaction is performed in DMF in a continuous membrane reactor which retains the catalyst. The Heck reaction product undergoes solvent exchange in a counter-current membrane cascade where DMF is continuously replaced by ethanol. After exchange, the product dissolved in ethanol passes through a column packed with an iron catalyst and undergoes reduction (>99% yield). Schematic representation of the proposed reaction-separation process is presented in Fig. 11A–C. The membrane unit operations showed excellent operational stability over a prolonged period of >2 months. The gridlock of imperfect membrane separation could be also resolved by combining solvent nanofiltration with other processes, retrofitting the existing technological processes in order to reduce the energy and carbon foot prints. Hybrid processes have been proposed combining OSN with reforming and distillation units (White and Wildemuth, 2006; Micovic et al., 2014; Lutze and Gorak, 2013), continuous crystallization unit (Ferguson et al., 2014; Rundquist et al., 2012b), column chromatography (Nimmig and Kaspereit, 2013; Rundquist et al., 2012a), adsorption (Pink et al., 2008; Sz
ekely et al., 2012), solvent extraction (Boam et al., 2014), and others.

5 Conclusions and Future Trends From the examples above, it is evident that the process engineering approach is able to overcome the main weakness of membrane separation technology—the imperfect separation. By careful selection of operating parameters and process optimization, membranes can find their places in the production lines. Improved design automatization and control is desirable in order to make membrane cascades acceptable in large production. Development of accessible databanks with membrane properties and appropriate design tools will boost the interest of industry toward this technology.

Fig. 11 (A) Reaction scheme for API synthesis; (B) Schematic representation of the alternative routes for solvent exchange; (C) New concept for a continuous process where the Heck reaction and the solvent exchange are performed in continuous membrane units. Adapted from Peeva, L., Da Silva Burgal, J., Heckenast, Z., Brazy, F., Cazenave, F., Livingston, A., 2016. Continuous consecutive reactions with inter-reaction solvent exchange by membrane separation. Angew. Chem. Int. Ed. 55, 13576–13579.

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