The Search for Process Intensification and Simplification ...download.xuebalib.com/ods8Kc5ovjN.pdfHigh density, large vial High density, large cell bag FIG. 34.1 Typical Upstream Platform

701Biopharmaceutical Processing. https://doi.org/10.1016/B978-0-08-100623-8.00034-7© 2018 Elsevier Ltd. All rights reserved.

The Search for Process Intensification and Simplification: Alternative Approaches versus Current Platform Processes for Monoclonal AntibodiesRobert S. Gronke, Alan GilbertBiogen Inc. Cambridge, MA, United States

34.1 OVERVIEW

Focused process development for monoclonal antibodies (mAb) began in earnest in the 1990s with positive clinical trial outcomes from a number of mAbs (e.g., Rituxan [1–3], Remicade [4], and Humira [5,6]). The main emphasis was placed on meeting regulatory agency requirements and ensuring commercial market supply with development emphasis placed on achieving titer of about 100 mg/L, high yield, product quality, process consistency, and ease of manufacture.

As the number of clinical indications for mAb treatment grew in the 2000s (e.g., rheumatoid arthritis, cancers of the colon and blood, psoriasis, psoriatic arthritis, cryopyrin-associated periodic syndrome, and osteoporosis [7]) into larger patient populations that required higher doses, the biotechnology industry responded by delivering processes with a 10-fold-increased titer (0.5–2 g/L), improving yield (~50%–70%), and purity. Simultaneously, development efforts needed to be minimized for each individual mAb process (i.e., platform approach), where possible, to save development time and labor while maximizing facility capacity and flexibility [8,9]. This platform approach yielded numerous innova-tions in process development including animal component-free media, fed-batch bioreactors, higher density cell cultures (10–20 × 106 cells/mL), higher protein A capture capacities (20–30 mg/mL), weak partitioning anion exchange, improved 15 nm virus filtration, and liquid drug substance [10–12].

New challenges arose in the late 2000s and continue today as the number and type of mAbs became more varied (bispecific antibodies, antibody-drug conjugates, antibody fragments, biosimilars, etc.), new innovations by commer-cial vendors such as rocking bag technology, disposal processing, higher capacity membranes and resins, and improved virus filtration (VF) technology became commercially available. Furthermore, longer fed-batch cultures with high cell densities have resulted in higher levels of product-related variants such as aggregates, partially reduced antibodies, acidic variants, or clipped forms that further challenge the downstream platform processing, as these typically are all captured on a Protein A resin.

Today, processes used to make mAbs continue to evolve as cost of goods, better facility utilization, and high concen-tration formulations become key drivers to deliver patients high-quality therapeutics with enhanced convenience. So how should a company best respond to the ever changing mAb marketplace with their processes? The answer lies in developing a mAb platform process strategy that is flexible enough to allow innovation.

34.1.1 Platform Approach for mAbs

Manufacture of mAbs using a platform approach has resulted in numerous advantages for biopharmaceutical companies, the broader biopharmaceutical industry and their regulatory agencies alike. These advantages include at least three busi-ness needs and drivers, such as: (1) experience gained from one mAb process can be applied to the next process such that

Chapter 34

702 SECTION | VI Industrial Process Design

high process efficiency and understanding within the organization are maintained, (2) raw materials, batch records, SOPs, etc., are similar thereby minimizing inventory variations, tech transfer costs, training time between campaigns, and cost of goods, and (3) facility fit and equipment including bioreactor design, tank sizes, and skid configurations are harmonized to make multiple product drug substances enabling a multi-use facility.

A typical industrial cell culture platform process used over the last 10–15 years for a mammalian cell line-derived mAb is described in Fig. 34.1 (Cell Culture) along with common variations on the platform (Fig. 34.2). Likewise, the downstream platform process is described in Fig. 34.3 (Harvest/Purification) along with common variations in Fig. 34.4.

Traditional vial

Perfusion

Fed batch

High density,large vial

High density,large cell bag

FIG. 34.1 Typical Upstream Platform Processing for Monoclonal Antibodies

Upstream variations Process

- High density cell bank

- Shake flasks, Wave bioreactor

- N-1 Perfusion

- Batch, Perfusion

N-1 = Seed train bioreactor immediately prior to production bioreactor

Working Cell Bank

Inoculum preparation

Seed train bioreactors

Fed-batch production bioreactor

FIG. 34.2 Common variations on the upstream platform process for mAbs.

The Search for Process Intensification and Simplification Chapter | 34 703

Harvestcentrifugation/depth filtration

Chromatography stepsProtein-A capture + 1-2 polishing steps

Virus filtration

UF (1-2) + DF

Detergent viral inactivation

Low pH viral inactivation

FIG. 34.3 Typical downstream platform process for monoclonal antibodies. UF, Ultrafiltration; DF, Diafiltration.

Downstream variations

-Pre-treatment with flocculants or low pH

-Microfiltration; for disposable facilities, depth filtration only

-Addition of Solvent and/or detergents prior to column

-Placement after 2nd column

-AEX in B/E mode -CEX in weak partitioning mode

-3rd column-molecule dependent-HIC in B/E or F/T mode

-15 or 20 nm pore size

-1 or 2 UF systems -SP-TFF in 1 system

-Excipient(s) added-Liquid storage

Cell cultureunclarified

harvest

Centrifugationand depth filtration

Protein Achromatography

(B/E)

Low pH viralinactivation

Anion exchangechromatography

(weak partitioning)

Mixed-modechromatography

(B/E)

Cation-exchangechromatography

(B/E)

Hydrophobic interaction

chromatography

Anion exchangechromatography

(weak partitioning)

Viral filtration

UF/DF

High concentration DS (Frozen Liquid)

Process

FIG. 34.4 Common variations on the downstream platform process for mAbs. B/E, bind and elute chromatography; F/T, flow through chromatography; SP-TFF, single-pass tangential flow filtration; UF/DF, ultrafiltration/diafiltration; DS, drug substance.


34.1.2 Changing Needs of the Platform: Higher Demand Requirements and Balancing Between Maintaining a Platform Versus Innovation

Over time, a number of issues with a platform have developed that can largely be grouped into a few areas. The first is an overall need for higher productivity and/or throughput in the mAb platform process than the platform currently delivers. This driver may come from commercial demand and/or facility usage optimization. Achieving this goal ensures uninter-rupted patient supply, lower cost of goods, and manufacturing flexibility. The next challenge would be to develop additional control requirements to successfully process batches of consistent and high quality where all critical product attributes and raw materials are well controlled with high yield and productivity. Realizing this objective may be measured by a reduction in failed batches and quicker batch release.

Conversely, there may be additional concerns to not change the platform operations as the operations may already work well in many respects. By committing to a platform process, some companies, however, take this approach to an extreme and thus risk discounting emerging methodologies and new technologies. Furthermore, by structuring the mindset of de-velopment and manufacturing to avoid change, innovation itself may be stifled. Consequently, a balance between a fixed, optimized platform process and new innovations to the process should be struck. Likewise, care should be taken to not allow one-off variations that are needed or unique to one mAb to become immediately incorporated into the mAb platform process. Table 34.1 summarizes the major pros and cons for maintaining a platform approach for a biological process along with suggestions on how to balance these often-opposing forces.

Platform mAb processes tend to be very streamlined for reasons described above. However, in doing so, they may not be developed well enough to handle specific product quality attributes beyond elimination of known safety issues

TABLE 34.1 Monoclonal Antibody Platform Assessment (Some Examples)

Platform Approach Pros Cons Comment/Advice

Single cell culture basal medium, and feed or perfusion medium

– Ease of sourcing components

– Consistency in preparation– Extensive knowledge base

developed with respect to the platform solutions

– Variations may be necessary to achieve desired product quality

– Unknown medium impurities are common to all processes using the platform

Companies need to decide how long to maintain a fixed cell culture medium platform.The platform could be a finite set of raw materials that could be combined to make unique solutions

Single cell culture production process

– Eases challenge of scale-down modeling

– Consistent deliverable for downstream

– May limit cell culture productivity A single approach is acceptable as long as all manufacturing facilities are identical. Variations may be required to accommodate facility tank size differences

Single host cell line – Simple cell line development workflow

– Less process development required to understand variety of host cell lines

– Requires development of additional process levers in order to reach wider product quality space

– Harder to deliver on a specific product quality profile

Companies need to develop alternative cell lines as not every mAb is best expressed in one cell line and alternative hosts can also produce products with varying attributes more easily

Use disk-stack centrifugation and depth filtration for clarification

– Can handle a wide variety of culture seed densities

– Works well for multi-product stainless facility

– Constraints observed with high VCD or low viability

– Challenges in scale-down modeling– May not be amenable for a single

use/disposable approach

A single approach may be acceptable as long as it aligns with the intended use of the manufacturing facility

Use a specific Protein A resin for every mAb purification

– High degree of consistency, yield and performance

– Single inventory item– Can buy in bulk quantity– Resin packing, storage and

cleaning parameters don’t need to be re-done each development cycle

– No secondary sourcing limits competitive price bidding

– Industry evolves such that resins with higher capacity and/or flow properties become available but not used

Companies need to decide how long they can maintain a fixed protein A resin before its advantages start to become more of a liability over time


(e.g., virus, aggregates, and microbial contaminants). Thus control of some product quality attributes generated during upstream processing beyond aggregation (e.g., levels of fucosylation, glycation, antennary structure, or acidic variants) is required, knowing that the downstream process can’t effectively handle most of these. One way to control these product quality attributes is through consistent upstream process implementation and delivering the same product quality profile generated during cell line selection. Downstream processing is challenged to handle high product mass and concentrations while achieving resolution of impurities. Processes can also drift due to raw material variations that result in out of trend/specification incidences. Thus a process developer must be careful to explore raw material variations while maximizing efficiency and process control simultaneously.

It is recognized that there are many variations on what each company decides is a fixed aspect of the platform versus an aspect that cannot be locked as part of the platform. One example for upstream processing where harmonizing a platform process could be difficult would be the seeding density for all stages in scale-up. An individual cell line may vary in growth rate sufficiently that fixing the seeding density for all cell lines could easily lead to overgrowth for the case of a high grow-ing cell line or insufficient cells for transfer to the next stage in the case of a low growing cell line. Thus it is better not to include a fixed seed density for all stages as part of the platform process. Instead, each cell line could have an independent seed density target on a stage-by-stage basis as the seed density can be quickly calculated based on the known growth rate of that cell line and the required cell numbers.

One example that is difficult for purification to establish as a fixed part of a platform is a common processing tempera-ture. In general, processing at ambient temperature is easier and thus preferred by manufacturing and process development. However, many mAbs are not stable at ambient temperature and undergo deamidation, clipping, acidic isoform formation, and/or aggregation. Recent new examples of mAbs instability have been reported at ambient temperature, including en-hanced trisulfide formation [13–16] or thioredoxin reductase (i.e., in the harvest) [17–20]). Running all platform processes at 2–8°C could be an option to enhance product stability, but there is a cost in terms of dynamic binding capacity loss (espe-cially on protein A or HIC), equipment expense (e.g., jacketed chromatography columns/tanks and/or use of cold rooms for process buffers), higher viscosities (i.e., higher inlet pressure), or longer processing times (especially VF and UF/DF due to flux decline). Thus, it is better not to include a fixed processing temperature as part of the platform process as it could be mAb dependent.

TABLE 34.1 Monoclonal Antibody Platform Assessment (Some Examples)—cont’d

Platform Approach Pros Cons Comment/Advice

Low pH viral inactivation

– Post protein A position is most convenient in process

– Generally accepted as a robust step for virus inactivation

– Many mAbs are low pH sensitive so alternative approaches are frequently needed

Innovation alternatives include use of solvent/detergent, detergent alone, high concentration arginine, and/or UV light, etc.

Always run AEX in a weak partitioning modea

– Best purity attained in most cases

– Fast and simple approach to developing a process step

– May remove/control a particular glycoform in the feed (e.g., acidic species)

– Always some yield loss– mAb variation with a very high pI

can achieve same high purity without yield loss (i.e., flow through mode)

– mAb variation with a low pI will take yield loss and even a larger yield loss still might not achieve high enough purity (e.g., low virus clearance, poor HCP removal)

The best strategy is to recognize and determine upfront the pI variations of incoming mAbs and apply the appropriate downstream chromatography approach

Single UF/DF for final formulation

– Applicable for a wide range of mAbs and formulations

– Methodology enables most final formulation requirements

– Limitations to meet high dose SC demands

– Over-concentration/back dilution needed to meet high step yield (e.g., aggregate formation)

– Concentration factors >50× will require two UF systems

SP-TFF is better suited in many circumstances (high yield, smaller footprint), no overconcentration, but diafiltration is still challenging

a Weak partitioning chromatography is an isocratic chromatographic protein separation method performed under mobile phase conditions where a significant amount of the product protein binds to the resin, well in excess of typical flow through operationsVCD, viable cell density; UV, ultraviolet; AEX, anion exchange chromatography; pI, Isoelectric point; HCP, host cell proteins; SC, subcutaneous; UF/DF, ultrafiltration/diafiltration; SP-TFF, single pass tangential flow filtration.


34.1.3 Process Intensification

Process intensification, as defined here, is the desire to improve productivity (grams/L/h or grams/gram resin-hr or plant productivity in grams/plant-h). Intensification of a production process is often necessary for a variety of reasons, including the need to mass produce mAbs (e.g., multi-tons/year) to meet patient demand without having to expand existing facili-ties, or development of high concentration formulations to enable subcutaneous dosing. These can be achieved through the use of higher titer in shorter production time (e.g., use of perfusion system at the N-1 stage), highly concentrated reactants (e.g., buffer concentrates) and intermediates/drug substance (e.g., use of SP-TFF), more compact operating conditions (e.g., higher loading ratios on columns), continuous processing (e.g., Simulated Moving Bed chromatography), and/or combin-ing unit operations into single units (pool-less processing, resin blending). Implementation of these technologies and meth-odologies can result in a dramatic reduction in the footprint of a process on the manufacturing floor, thereby minimizing investment and resources and potentially improving the speed to market for patients [21–24]. Although cost reduction could be considered part of process intensification, we will not address this topic in this chapter.

If one looks at facility utilization for mammalian capacity, it has been remained relatively constant despite the addition of new pipeline products over the period of 2009 to 2013 due to the substantial titer increases realized across the industry. Looking at current bioreactor capacity and company’s phase 2 and 3 pipelines (including biosimilars), facility utilization is predicted to average about 80% out to 2020 [25]. A new 160,000-square-foot biological facility with 15,000 L bioreactors requires over 500 employees and was estimated to cost approximately $1 billion USD [26], forcing companies to continue to innovate to maximize the use of existing facility space.

For those companies using batch or fed-batch bioreactors, improving titers from 5 g/L to 10 + g/L and downstream process throughput for protein A from 500 mg/L/day to 1200–1500 mg/L/day are required to avoid having to build new facilities for large mass requirement products. For companies using perfusion technology, continuous improvement in titer and longer run times are putting pressures on cell line development to ensure that genomic stability and product quality can be maintained throughout the culture duration [27]. If successful, these innovations will allow a company to move from making 3.3 metric tons of mAb per year to >10 metric tons per year. Alternatively, one could intensify a process by mak-ing it fully or partially continuous downstream. This topic however will be addressed in other chapters of this book (see Chapters 28 and 36).

34.1.4 Process Simplification

Process simplification, as defined here, is a technique designed to eliminate wasteful or non-value added actions, reduce process cycle time, and remove disconnections between unit operations. For a process developer, process simplification translates to designing the same process with less effort which can include less manipulations or time within a unit opera-tion, between units, or even eliminate unit operations. Examples include use of a high density cell bank (eliminates manipu-lations during inoculum preparation), combining two downstream unit operations into one (e.g., substitution of a two-step chromatography with a mixed mode resin), on-column chemical treatment vs. a two-step operation (trisulfide reduction or on-column viral inactivation), or seamlessly blending one unit operation into the next to avoid extra processing (e.g., elimi-nation of in-between step UF/DF, avoid pH/dilution adjustments, intermediate hold tanks). Examples of both intensification and simplification are described in more detail below.

34.1.5 Examples of Intensification and Simplification: Cell Culture

Generation and Cryopreservation of Cell BanksAs each cell culture process begins with the thaw of working cell bank vial, the cell bank generation process is the first step suitable for intensification. A traditional cell culture process would begin with thaw of a vial and sufficient cells to inoculate a small flask of approximately 25–100 mL working volume. Compared to the scale of an industrial production bioreactor, a large scale-up factor remains from this initial thaw, which requires time in the facility and resources to main-tain and monitor the culture. A higher cell density cell banking process that ultimately permits a larger working volume at thaw is therefore becoming the standard operating procedure. Tao and coworkers demonstrated a perfusion-based cell banking process resulting in a high density 4.5 mL cell bank containing 100 × 106 viable cells (vc)/mL that permit-ted direct inoculation of a 20-L rocking bioreactor [28]. Heidemann and colleagues also utilized a perfusion-based cell banking process but created a larger volume cell bank in a 100-mL cell bag at 20 × 106 vc/mL [29]. More recently Seth and coworkers created a frozen seed train intermediate that consisted of a 150-mL cell bag at 70 × 106 vc/mL [30]. The resulting combination of large volume at high cell density allows the cell culture process to begin in an 80 L bioreactor.


The stated intention of these seed train intermediates is simply to supply an individual manufacturing campaign. Based on the volume of cell culture required per bag and the bioreactor size used in this process, a cell bank of over 100 cell bags could theoretically be created.

The latter case of a high density cell bank in a 150-mL cell bag is atypical, and this individual manufacturing step is an extreme case of high cell density. The resulting cell bags require at least two orders of magnitude more cells when com-pared to a traditional cell banking process. However both the 4.5-mL vial and large volume cell bag are described to save multiple stages and operations within the manufacturing scale up process. Compared to a traditional vialing process, one could intellectually assign the complexity associated with manufacturing a higher cell density cell bank to the first batch. Starting with the second batch, the manufacturing process then uses a simplified manufacturing process, and the payback in reduced labor for a process with fewer stages effectively more than pays for the complexity in making the cell bank. With a recent report achieving over 200 × 106 vc/mL via perfusion, even larger bioreactors could theoretically be inoculated directly if the cells remain suitable for an industrial-scale cell banking process [31]. Overall the trend in cell banks is clearly in the direction of higher cell densities and higher volumes.

Establishing Platform Cell Culture SolutionsA process with a single basal medium simplifies the entire cell culture process. The medium preparation process is rep-licated with only gravimetric requirement changes as the scale increases. Maintaining consistency from preparation to preparation should lead to reliable medium with the intended properties for cell culture. Consistent preparations cannot be guaranteed due to the potential for raw material impurities to impact outcomes. The medium must also be optimized for the specific situation as a feed medium in fed batch production is not interchangeable with a perfusion medium. However, by maintaining a single basal, perfusion, and/or feed medium for an entire cell culture platform, a company can establish increased experience and data pertaining to the raw materials required and any potential impurities.

Even without considering raw material impurities, cell culture medium is a complex mixture of approximately 30–50 components including sugars, amino acids, buffers, salts, vitamins, and trace metals. To simplify the preparation, a platform cell culture medium will be created mostly by dissolving a single bulk powder containing accurately proportioned compo-nents in water. This step guarantees the precise delivery of a large number of components on the manufacturing floor in a simple way. Additives incompatible with the manufacturing process of this powder may be added after the powder is dis-solved. There are a variety of technologies available at cell culture medium manufacturers to create this powder, including pin milling, granulation, and compaction [12,32]. The benefits of pursuing these cell culture media alternative technologies need to be balanced with the potential for simplification by outsourcing the medium generation and sourcing directly liquid medium [33].

Defining a single cell culture medium is a critical step to simplify the platform. Recently Roche and Genentech elimi-nated the differences between the cell culture medium platforms and developed a single cell culture platform [34]. The resulting optimized process delivered an average 30% increase in titer. A process which easily delivers consistently high titers is critical to the business drivers behind development of a platform for monoclonal antibodies. Medium development work may still be required in case of a difficult-to-express protein [35], or when product quality drivers exist such as the need to alter glycosylation [36–40].

Optimizing the Production BioreactorA major limitation of production in a facility is the production bioreactor. The production bioreactor process duration (commonly about two weeks for a fed-batch process) is frequently a rate limiting step in an industrial facility, given that all other upstream and downstream process steps require only up to a few days. Consequently, several groups have investi-gated utilizing perfusion in the bioreactor immediately preceding production [21,41–43]. The common net result is higher densities at inoculation of the production stage, shorter process durations, and similar titer and product quality. Thus the production process is even more productive with higher overall facility volumetric productivity. The additional challenge lies in optimizing perfusion medium to generate the necessary cells for the production bioreactor without creating new burdens in frequent, large volume medium preparation in a facility. However, by shifting from a standard vial to a high density vial or high density cell bag and to a perfusion seed bioreactor, many manufacturing steps and days in the process can be eliminated as seen in Fig. 34.5. The data in Fig. 34.5 demonstrates the possibility of combining all of the previ-ously described high productivity or high cell density steps into a single optimized process. In totality, approximately two weeks in upstream manufacturing could be eliminated per batch by implementing a high cell density bag. From a differ-ent perspective, over 20% increase in overall mAb mass produced could be realized in the same manufacturing facility by implementing the perfusion seed bioreactor.


Production Platform DecisionIn a production platform, the key parameters are the cell densities achieved over time and the cell-specific productivity as these values determine the overall cell and product mass generated in the cell culture process. The titer in a fed batch process can be directly described by a two-dimensional contour plot comparing the titer achieved for various integral viable cell concentrations (the area under the viable cell density curve) and the cell specific productivity. On the contour plot, one can overlay lines of constant titer or isotiter curves (Fig. 34.6). The same titer can be achieved by a high cell density and low specific productivity cell line or a low cell density and high specific productivity cell line. However, a titer approaching 20 g/L will likely require both a high cell density and high specific productivity. In a perfusion platform, the same type of contour map can be constructed but the perfusion rate becomes a critical variable. Fig. 34.7 demonstrates the titers obtainable assuming a fixed specific productivity of 50 picograms per cell day. In this case, high cell densities may be readily achievable due to the removal of waste, but titers may remain low due to high perfusion rates. While there are various reasons a scientific organization might choose a fed-batch or a perfusion process for a production mAb, one consideration is the antibody mass needed. A hypothetical perfusion and fed-batch process each producing antibody at similar rates in cell culture is depicted in Table 34.2. Of course there are some underlying assumptions in the total antibody mass calculations made per day. For example, a required cell bleed rate that potentially loses product and the required time to achieve the density described in perfusion are ignored in these calculations as is the turnaround time for the production bioreactor in the fed batch over the course of 60 days. In addition, it is conceivable that a cell line may have a different (potentially higher) cell specific productivity in a perfusion system. The intention of this table is simply to spell out the calculation that should be considered before choosing a production platform. When considering the differences between perfusion and fed batch, it is important to remember that the target output is mass, and one of the key inputs is medium. Consequently, titer is diluted in a perfusion system as the perfusion volumes increase, and the key parameter to calculate is the mass produced per liter of medium utilized, rather than the mass produced per bioreactor hold-up volume.

00

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(10

6 ce

lls/m

l)

12

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* (4) (3) (2) (1)

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40 50

FIG. 34.5 Comparison of theoretical viable cell density and process duration after inclusion of process intensification steps. (1) standard process, (2) high density vial, (3) high density cell bag, (4) high density cell bag combined with perfusion seed bioreactor (*).

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Inte

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40 50

FIG. 34.6 Isotiter plots for fed batch production. Black lines reflect constant titer combinations of integral viable cell concentration and specific productivity, and the resultant titer is listed inside the square.


34.1.6 Examples of Intensification or Simplification: Purification

HarvestCentrifugation, in combination with depth filtration, has been the workhorse for large-scale harvest clarification since the late 1990s. Over the past 10 years, advances in bioreactor development have achieved strikingly high cell densities of over 150*106 cells/mL while maintaining reasonably high harvest viabilities of >70% [31,44]. This combination of high cell mass and associated cellular debris however has required renewed efforts in harvest technology and processing just to maintain yield, product quality, and process throughput (see Chapter 9).

Pre-treatment clarification technologies, such as acid treatment, addition of flocculants, or precipitation continue to im-prove such that one can now target not only the cellular debris but also process-related impurities (e.g., DNA, host cell pro-teins, endotoxin) or even the product itself [45]. Without pre-treatment, the centrifuge packed-cell volumes can approach 15%–20%, which approaches the limits of the centrifuge’s capabilities. Shot intervals (i.e., the length of time between discharges of the concentrated cell mass) and their accompanying turbidity perturbations may be impacted. Nozzles in the centrifuge restrict the flow of the solids/concentrate stream exiting the centrifuge. In a typical disk stack centrifuge, one can achieve a flow rate of about 0.07 L/min per nozzle. Fig. 34.8 indicates that by increasing the number of nozzles from 4 to 8, the saturation flow rate is increased proportionally thereby achieving processing at packed cell volumes (PCV) up to about 16% (Panel A) within a desired flow rate range. Maximum PCVs can then be calculated for each feed flow rate, which would represent when the nozzles are fully saturated (i.e., 100% solids in the solids stream). Panel B compares the

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Steady state cell density(106 viable cells/mL)

Per

fusi

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ol/v

ol/d

ay)

80 100 120 140

FIG. 34.7 Isotiter plots for perfusion production. Black lines reflect constant titer combinations of perfusion rate and steady state cell density, and the relevant titer is listed inside the square. Titer was calculated assuming a fixed cell specific productivity of 50 picograms/(cell*day).

TABLE 34.2 Comparison of Upstream Yield Based on Upstream Process Mode of Operation

Attribute Units Perfusion Fed Batch

Cell Specific Productivity picograms/(cell*day) 50 50

Steady State Cell Density 106 cells/mL 100 N/A

Bioreactor perfusion rate Volume/Volume/Day 2 N/A

Integral Viable Cell Concentration 106 cell*day/mL N/A 250

Upstream Duration days 60 14

Titer g/L 2.5 12.5

Bioreactor Volume L 2000 15,000

Upstream Yield Per Batch kg 600 187.5

Upstream Yield Per Day kg/day 10.0 13.4


centrate turbidity profile for a traditional nozzle centrifuge (blue trace) vs. a newer continuous discharge centrifuge (pink trace). A continuous discharge centrifuge would operate at a flow rate slightly below the point of saturation but would maintain a constant baseline due to continuous removal of solids. Thus, improvements in the centrifuge itself can alleviate the throughput constraints a high-density cell culture can have on the harvest.

A more recent development in harvest technology is acoustic wave separation (AWS), developed by FloDesign Sonics and licensed to Pall Corporation, for cell clarification and perfusion in the production of immunoglobulins (glycoproteins). This technology, which employs three-dimensional standing acoustic waves to trap cells and cellular debris, forces particles to cluster and settle out of solution. This technology has the potential to reduce particulates in high density cell culture to levels that are more easily manageable for simpler depth filtration. It also would allow for continuous clarification and single-use formats without the challenge of fouling physical filter membranes. Clarification efficiencies of greater than 95% with yields of greater than 85% have been reported [46,47].

With respect to product quality, a recent example of increased mAb reduction occurring during harvest intermediate hold has occurred due to increased cell culture densities. This has required developers and manufacturers to implement oxidizing strategies [18], re-introduce cold temperature downstream processing and/or implementation of charged depth filters to prevent disulfide reduction during harvest [48,49].

Protein A CaptureProduct capture of mAbs onto Protein A affinity chromatography has been the hallmark choice of purification for decades due to its selectivity for a wide variety of mAbs and Fc-fusion proteins, its robust removal of process-related impurities,

Panel A

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Bufferdisplacement

Typical processing range

FIG. 34.8 Impact of nozzle number and discharge interval on centrifugation flow rate and turbidity profile. (A) Flow rates and packed solids volume required to saturate a Westfalia HFC-15 equipped with 4 or 8, 0.2 μm nozzles. (B) Impact of intermittent vs. continuous discharge type centrifuge on centrate turbidity profile.


high yield (e.g., >90%) and tolerance/interface with a variety of upstream clarified media feed streams [50]. Resin dynamic binding capacity has progressively increased from the 1990s to today from about 20 to 70 g/L with the advent of new resin matrices, base stable ligands, higher ligand densities, and less compressible resins. Unfortunately, the cost per liter of resin, however, has not declined over this same period of time, and in many cases has increased. If one examines the changes in binding capacity as a function of cost per L of resin over time, the price has essentially been fairly stable or slightly declin-ing over time. Nevertheless, maximizing the capacity of protein A has generated some intensive development efforts. One example has been to simply implement a dual flow rate loading strategy, whereby a faster flowrate (i.e., shorter residence time) is used in the initial stages of loading (when all the binding sites are available), followed by a reduced flowrate (i.e., residence time increased). This processing method enables additional antibody to diffuse into all the pores and bind to the less readily accessible sites, helping to achieve high capacity while maintaining acceptable processing times (Fig. 34.9) [51]. Another example is to increase the number of re-usable cycles of the resin via improved cleaning strategies [52].

A third approach would be to implement a continuous capture approach via either simulated moving bed technolo-gies (sometimes referred to as periodic counter-current chromatography), multi-column solvent gradient purification, or continuous countercurrent tangential chromatography. This later approach, which not only maximizes resin capacity but productivity as well, is getting closer to implementation into commercial processes [53].

Improving the Downstream Process Productivity (Post-Protein A Capture)An often-overlooked step in the downstream process are those efforts required to manipulate the feed stream between unit operations. Traditional processing between downstream chromatography steps of an optimized process (i.e., no UF/DF needed to enable further processing) typically consists of collection of the previous eluate pool (or effluent pool if the step is run in flow-through or weak partitioning mode), mixing, load adjustment as needed (e.g., pH adjustment, dilution, and/or salt addition), and additional mixing prior to further processing. Besides those aspects, sampling from each of these steps may also be performed.

One advantage of the traditional processing approach is that it allows some control over the process if decisions are needed between columns to determine how the next step needs to be run (e.g., loading ratio for the next column) along with some modest flexibility in scheduling, assuming stability is established on the intermediate. But how often is this control option (i.e., information) actually utilized in a commercial process?

Additionally, this level of process control can be limited by the process intermediate stability and the overall desire to increase downstream productivity. Taken in its entirety, the between-step processing can be quite cumbersome and labor in-tensive in that it requires 1–2 collection tanks/bags, 1–2 mixers, 1–2 mixing validations, titration curves (if pH adjustment), stability data, plus manufacturing and/or QC analysis time and resources. Furthermore, the more manipulations there are required between steps, the more chances for even a robust process to undergo a deviation during routine manufacturing. Ideally, if one could develop a process that eliminated at least some of the between-step processing, the benefits could be significant.

Considering the typical downstream platform process for mAbs (Fig. 34.3), one opportunity is between columns 2 and 3, a place where a viral inactivation treatment step typically is not inserted. For example, where AEX and HIC are run in series, the between processing includes both a pH and salt adjustment. Development work using strongly hydrophobic resins such as Hexyl or Phenyl have shown that the need for lyotropic salt addition to the feed of HIC columns can be

Dual flowSingle flow50

DB

C @

10%

BT

(mg

pro

tein

/mL

res

in)

55

60

65

70MabSelect SuRe LX

FIG. 34.9 Comparison of dynamic binding capacity (DBC) using single flowrate vs. dual flowrate loading strategy on MabSelect SuRe LX at 10% break through (BT).


eliminated whereby the HIC column is now run in the flow-through mode with high molecular weight aggregates retained on the resin [54]. More recent development has focused on enabling in-line pH adjustment to the HIC which can eliminate the intermediate pool.

In collaboration with GE Heathcare, Biogen has been developing a platform on ATKA systems called “straight through processing (STP).” This concept includes a multi-valve system that enables between column in-line adjustment (pH and/or salt) and mixing. Effluent from the AEX column is passed through an in-line mixer and monitored for UV. A UV signal triggers titration of a load adjustment feed to meet a consistent output pH. Recent laboratory results using this step show that straight-through processing was achievable for several mAbs when the HIC column is used in a binary mode to remove higher molecular weight aggregates and host cell proteins Table 34.3, with or without salt addition. The results were not as favorable for a fusion protein where the subsequent HIC column was run in a bind and elute mode, but this was attributed more to a process that was sensitive to variation.

Overall, the goal of eliminating an intermediate hold tank, improving process throughput and its associated labor was accomplished via straight through processing without sacrificing process control, product yield, or purity [55].

Polishing Step ImprovementsThe strategy behind mAb downstream development has traditionally been designed to let the Protein A capture step do the job of removing host cell proteins, DNA, and media components while concentrating the feed stream, to let the downstream polishing steps focus on removal of high and low molecular weight impurities and protein A leachate, and to provide virus clearance. While true in theory, carryover of HCPs and DNA into the Protein A eluate often occur at higher than desired levels, primarily due to complex formation with the target mAb [56]. Thus 2–3 polishing chromatography columns run in a bind and elute or flow through mode comprised of ion exchange and HIC have been the choices for polishing steps (see Fig. 34.4). Overall downstream yields typically reach 50% but rarely above 70% in part due to the polishing steps.

Mixed Mode Resins

While resin manufacturers continue to deliver polishing resins of higher capacities and a wider variety of bead and pore sizes, the advent of newer mixed-mode resins from the manufacturers can have the biggest opportunity to streamline the downstream process. Their implementation however has been met with varying degrees of success. On one hand, mixed-mode resins are very salt tolerant, can achieve removal of product-related impurities (e.g., aggregates/dimer) with high yield and can provide good virus clearance in a single downstream unit operation. On the other hand, clearance of HCPs has been less robust. Thus, the strategy of replacing two single mode chromatography resins with a mixed mode resin will be case by case depending based on the types of impurities in the post protein A feed stream. Table 34.4 outlines the pros and cons of using a mixed mode resin vs. a traditional two-column polishing process. Table 34.5 highlights some recent data on the CaptoAdhere resin from GE Healthcare for two different mAbs when run under high salt, neutral pH loading conditions.

Resin Blending

As one can discern from the above, the key to enabling downstream pool-less processing is to implement in-line mixing that requires only simple single solution adjustment between steps. However, if a process developer can find conditions where no such adjustments are needed between columns (i.e., eluate from one column can be applied directly to the next column),

TABLE 34.3 Comparison of Batch Versus Straight-Through Processing

Operation Mode (AEX to HIC) Molecule Process

Salt addition Between Columns?

Cumulative % Yield (AEX-HIC)

% HMW in HIC pool

HCP in HIC pool (ppm)

FT to FT mAb1 BatchSTP

No 9190

0.660.60

2.82.2


No 8282

0.540.52

0.80.8


Yes 7878

0.740.71


other options beyond pool-less processing open up. One such option is a concept called resin blending, where two (or more) single mode chromatography modalities are blended together and packed into a standard column format. This approach dif-fers from a traditional mixed-mode resin where either two or more different ligands or functional groups are incorporated into a single resin bead. In addition to elimination of intermediate holding tanks and mixing, combining resins into a single column would also eliminate multiple column packing and qualification, multiple skids and filtration, and simplify batch records to further increase productivity and simplicity. Combinations of resins that could be blended together, in theory are limitless, but must function such that parameter variations in one resin don’t affect the other resin’s performance. Often the order of contaminant removal is not sequential, and thus the resins don’t need to be stacked one on top of the other to achieve the same result.

A resin blending experiment was carried out for a high titer mAb using protein A eluate that was neutralized to pH 6.0 or 7.0. One strong AEX resin was pre-blended with a HIC resin and packed into a 5-mL column (0.66 × 15 cm) at a resin blend ratio of 3:4 to match target loading ratios for each resin, respectively. In the control process, both resins were designed to run in the flow-through mode whereby the AEX resin was run at pH 7.0 and the HIC resin was run at pH 6.0. Results in

TABLE 34.4 Comparison of Single Mode vs. Mixed Mode Chromatography

Two-Column Polishing (Post Protein A) (e.g., AEX + HIC) Single Mixed Mode Resin (Post Protein A) (e.g., AEX/HIC)

Advantages

Both separation mechanisms are straight-forward and thus can be quicker to understand, develop, and validate; current platform approach

Easier to integrate into a process (e.g., more salt & pH tolerant) without need for upfront dilutions or insertion of UF/DF steps prior to the column

Since clearance mechanisms are orthogonal, clearance across both columns can be additive, resulting in higher claimed clearance values

In B/E mode: One can apply a variety of aggressive washes to improve purity (i.e., trade-off single modes such as salt and pH) without yield loss

AEX provides excellent clearance of DNA, HCP and viruses as long as pH is >7; HIC provides clearance of aggregates/dimers

Can minimize the number of process steps/columns and thus have the potential for higher overall yield

Two-step polishing process has more potential to control impurities vs. a one-step downstream process

Recent work has demonstrated that robust viral clearance can be achieved

Eliminates need for in-between step processing including pH and salt adjustment, hold time validation

Disadvantages

Clearance of product-related impurities (e.g., dimer) is spotty on AEX, thus requiring the second HIC column that can clear dimer but often not much other process-related impurities, especially in when run in flow-through mode

More non-specific binding of impurities as these resins are comprised of closely spaced ionic, hydrophobic and sometimes hydrogen bonding moieties (e.g., Capto Adhere from GE Healthcare)

Combined yield is often lower than a single column Clearance of process-related impurities (e.g., HCPs) is spotty

For low pI mAbs, AEX is poor at clearing HCP, virus and thus it’s overall usefulness is questionable

If needed, insertion of a 2nd downstream polishing step may inhibit claiming combined virus clearance on both steps due to non-orthogonal approaches

TABLE 34.5 CaptoAdhere Yield and Impurity Clearance Assessment

AntibodyLoading Ratio (G/L resin) Product Yield (%)

% Dimer Removal

HCP LRVa (Log10)

Model Virus LRV (Log10)

X-MLV MMV

mAb1 275375

8488

5248

0.880.75

4.43.9

>4.52.7

mAb2 350 92 260 0.3 3.1 1.8

a LRV = Log10 reduction value of virus.Load Conditions Include pH Ranges Between 6.5 and 7.7 and conductivities between 20 and 30 mS/cm.


Table 34.6 show that equivalent yields, purities (HCP clearance, aggregate removal), and MMV virus clearance could be attained at pH 7.0, regardless of whether the resins were run in (1) traditional step-wise two-column fashion, (2) stacked one (AEX) on top of the other (HIC), or pre-blended together into a single packed column (Fig. 34.10).

Product Concentration/Final FormulationConcentration of the product feed stream, either during or at the end of downstream processing, is often required to allevi-ate facility tank constraints or reduce final formulation volume, respectively. Traditionally, tangential flow filtration (e.g., UF/DF), which employs membrane modules used in parallel for product concentration and/or diafiltration (Fig. 34.11, panel A) has been the preferred approach. During the UF/DF operation, the product is often over-concentrated beyond the intended target such that a post-rinse of the system enables high yield (i.e., 85%–95%) and is recirculated requiring sev-eral pump passes that could result in product denaturation (e.g., aggregate formation). Due to ever increasing demands for higher concentration formulations to enable subcutaneous dosing (e.g., ≥200 mg/mL), other process limitations of using UF/DF have surfaced including a significant increase in viscosity (e.g., >50 cP) and its accompanying decrease in feed flow rate (i.e. longer processing time). Increasing temperature can decrease viscosity and improve feed flow rates [57].

More recently, the use of a single pass tangential flow filtration (SP-TFF, Pall Corporation) system has been designed and evaluated, whereby the product is passed through membranes in series (i.e., no recirculation, minimal overconcentration) (Fig. 34.11, panel B). This could enable a high titer mAb to fit into an existing 15 K liter stainless-steel facility via insertion of an SP-TFF post- harvest, post Protein A column and/or during final drugs substance formulation (Fig. 34.12). Results in Table 34.7 show comparable or an improved recovery with SP-TFF vs. UF, however there was an increase in processing time.

Case 1Platform Process

Load Adjustment

Load Adjustment

Load Adjustment

No Load AdjustmentSame conditions

Load Adjustment Load Adjustment

Protein A Protein A Protein A

AEX AEX AEX

HIC HIC

HIC

Protein A

Case 2Pool-less processing

Case 3Stacking resins

Stacking resins:one resin bed on

top of other bed inthe same column

Resin blending:Resins pre-mixedand packed in one

column

Case 4Resin blending

FIG. 34.10 Progression to further simplify two downstream chromatography steps in series (AEX and HIC).

TABLE 34.6 Resin Combination Study: Yield and Impurity Summary

Column Configuration % Step Yield Overall % Yield % Aggregate in Eluate % HCP (ppm) MMV LRV (log10)

Protein A eluate feed N/A N/A 1.4% 40 N/A

Control processa N/A 89% 0.9% 5 4.3b

AEX (pH 7.0)c 96% N/A 1.5% 17 3.2

HIC (pH 6.0)c 85% 82% 0.6% 6 1.1

Resin stacking (pH 6.0) N/A 95% 0.8% 9 N.D.

Resin stacking (pH 7.0) N/A 90% 0.8% 4 N.D.

Resin blending (pH 6.0) N/A 92% 0.8% 11 2.7

Resin blending (pH 7.0) N/A 88% 0.6% 4 >4.4

a Two-column process, AEX run at pH 7.0, HIC run at pH 6.0.b Value is the sum of AEX at pH 7.0 and HIC at pH 6.0 in this data set.c Control conditions via two separately packed columns with pH adjustment between AEX and HIC columns.N.D. = Not determined; N/A = Not applicable.


Filter ModuleFeed Pump

Feed Pump orPressure Source

Filter Module

Feed source

Permeate Permeate

Permeate

Retentate

Buffer

Panel A Panel B

RecirculationTank

Retentate

FIG. 34.11 Schematic diagrams for UF/DF (A) and SP-TFF (B).

15K MFG purification process volumes

10X Concentration

5X Conc.

UF/DF

2g/L

25000

20000

15000

Pro

cess

Vo

lum

e (L

)

10000

5000

0

5g/L

10g/L

VFHIC FTAEX FTPro AHarvest

Tank constraint

FIG. 34.12 Examples where SP-TFF could alleviate volumetric issues on the manufacturing floor.

TABLE 34.7 Comparison of UF/DF vs. SP-TFF (Final Formulation Example)

Parameter UF/DF Final UF Results SP-TFFa SP-TFF Results

Process Time (min)

Higher productivity, especially at higher concentrations, due to higher crossflow rate with same membrane area

74 Lower feed flow rate required to achieve concentration factor in single pass leads to longer process times, especially at high concentration

134

%Yield Typically lower due to volumetric constraints. Often requires significant overconcentration to perform effective rinse

82% Similar or higher than UF and less dependent on mass loading. Single pass rinse enables effective rinse with minimal dilution.

88%

Product Impact Product recycling leads to numerous pump passes at high concentration which can result in soluble and insoluble aggregates for sensitive proteins

% aggregation increase = 0.34%

Eliminates repeated pump passes for high concentration protein solutions. Results in improved filterability

% aggregation increase = 0.20%

Facility Fit Significant volumetric constraints due to size of retentate vessel and holdup volume of skid

Required 2 UF/DF systems

Increased flexibility due to elimination of recirculation tank. Same system can be used for multiple applications

Single UF/DF plus SPTFF

Final Concentration Achieved (mg/ml)

Typically requires ≥20% overconcentration

232 Can achieve higher final product concentrations due to reduced rinse volume

277

a SP-TFF uses multiple membranes connected in series using a single pass of the material through the system.


One constraint using SP-TFF is its inability to be used for diafiltration. However, recent work has shown that under certain conditions, continuous diafiltration can be carried out using SP-TFF and still obtain high yield and throughput [58].

34.1.7 Solving Problems at the Interface Between Upstream and Downstream

In addition to the examples provided above, there are several other approaches to increase the productivity of or simplify the drug substance process. A partial list is shown below in Table 34.8.

Despite the rigor involved in developing a process platform, new discoveries in process or product quality as platforms are increased in productivity may require thorough scientific investigations and novel solutions. In the case of unexpected antibody reduction at harvest, aeration was required to protect the molecule [18] and further examination identified thio-redoxin 1 activity as the root cause [19]. In the case of the unexpected trisulfide bond antibody formation (CH2SSS CH2) within the interchain, hinge region bonds of a human IgG2 mAb [16] or in the HH and HL interchain bonds of all IgG subclasses, the problem was tackled from both an upstream and downstream perspective.

Once cysteine was identified as the source of the problem in cell culture with respect to trisulfide formation, cysteine could be limited through manipulation of the nutrient feed composition [15]. Decreasing the cysteine concentration in nutrient feed led to less hydrogen sulfide release and directly to decreased trisulfide content. Once the mechanism was understood, inhibitors of hydrogen sulfide release were screened directly. Several inhibitors, including pyruvate, were identified. An example of inhibited hydrogen sulfide release from nutrient feed medium can be seen in Fig. 34.13. Higher concentrations of pyruvate supplemented to nutrient feed decrease hydrogen sulfide formation and consequently prevent trisulfide formation.

Alternatively, a novel downstream approach was developed to eliminate trisulfide bonds in the antibody using an on-column cysteine wash step that could effectively reduce the trisulfide bonds back to intact disulfide bonds without disulfide

TABLE 34.8 Additional Recent Efforts for Intensification or Simplification of a Monoclonal Antibody Process

Process Section Intensification Simplification

Cell Culture N-1 perfusion, PAT to control product quality, monitor cell growth, and/or product titer

Single use technologies, ecofriendly reagents

Harvest/Purification Buffer concentrates, feed-forward/feed-back process control, real-time release, continuous downstream processing

Pre-packed columns, intermediate collection in disposable bags, membrane purification, purification by precipitation

DS/Formulation >200 mg/mL DS, viscosity reduction, novel excipients, crystallization

Liquid storage, ready-to-fill

0

0.5

1

1.5

2

Hyd

rog

en S

ulf

ide

(pp

m)

10x Pyru

vate

5x Pyru

vate

1x Pyru

vate

0.5x Pyru

vate

0.1x Pyru

vate

Feed M

ediu

m

Cystein

eFree M

ediu

m

FIG. 34.13 Release of hydrogen sulfide from feed medium is inhibited by pyruvate supplementation: Sodium pyruvate was supplemented to feed me-dium at the listed ratio relative to the cysteine concentration in feed medium. Cysteine-free feed medium and cysteine-containing feed medium lacking sodium pyruvate supplementations are presented as controls.


scrambling or over reduction, as monitored by an increase in free sulfhydryls [13]. A cysteine wash of sufficient concentra-tion was applied to the IgG bound to Protein A or ion exchange prior to its elution. Results in Table 34.9 show that a Protein A wash of 1 mM cysteine has been shown to be effective in reducing trisulfides from a range of starting trisulfide concen-trations (8%–13%) down to


ACKNOWLEDGMENTSWe kindly would like to thank the following Biogen contributors to this chapter for their input and/or data: James Lambropoulos, Christina Alves, Rashmi Kshirsagar, Alex Brinkman, John Armando, Jennifer Zhang, Ratnesh Joshi, Sanchayita Ghose, Hiro Aono, Susanne Alexander, John Pieracci, and Matt Westoby.

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TABLE 34.10 Impact of Improvements on Manufacturing Network Productivity

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Titer increase from 3 g/L to 10 g/L 2X

N-1 Perfusion + Buffer concentrates 3X

Improved downstream platform 4X

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The Search for Process Intensification and Simplification: Alternative Approaches versus Current Platform ProcesseOverviewPlatform Approach for mAbsChanging Needs of the Platform: Higher Demand Requirements and Balancing Between Maintaining a Platform Versus ...Process IntensificationProcess SimplificationExamples of Intensification and Simplification: Cell CultureGeneration and Cryopreservation of Cell BanksEstablishing Platform Cell Culture SolutionsOptimizing the Production BioreactorProduction Platform Decision

Examples of Intensification or Simplification: PurificationHarvestProtein A CaptureImproving the Downstream Process Productivity (Post-Protein A Capture)Polishing Step ImprovementsMixed Mode ResinsResin Blending

Product Concentration/Final Formulation

Solving Problems at the Interface Between Upstream and Downstream

Discussion/ConclusionsAcknowledgmentsReferences

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The Search for Process Intensification and Simplification ...download.xuebalib.com/ods8Kc5ovjN.pdfHigh density, large vial High density, large cell bag FIG. 34.1 Typical Upstream Platform