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An alternative approach to traditional Protein A schemes is comparable in overall efficiency, product recovery, and quality.
In human monoclonal antibody (HuMAb) processes, nonaffinity purification schemes have been developed as viable alternatives to affinity schemes, especially when cell culture expression levels have been increased significantly. Primary recovery by tangential flow filtration (TFF) for concentration and diafiltration plays a crucial role in a nonaffinity purification platform with cation exchange as a capture step. A nonaffinity purification scheme, including an optimized primary recovery TFF step, is comparable to Protein A-based purification processes in overall process efficiency, product recovery, and quality.
Tangential flow filtration is a common unit operation either for cell culture clarification and concentration or just for concentration of the cell culture supernatant in biopharmaceutical manufacturing where affinity or nonaffinity capture is used. As a platform approach, exploring nonaffinity capture as a viable alternative to Protein A has yielded efficient purification schemes. These nonaffinity schemes with three, two, or one chromatography steps have been successfully developed for HuMAbs.1,2 Cell culture harvest needs to be preconditioned by concentration and buffer exchange to the appropriate pH and conductivity to have a uniform and maximum load on CEX column. For this primary recovery step to fulfill both concentration and diafiltration, a TFF step is generally optimized for different process parameters. During the diafiltration step, significant slow precipitation from the cell culture harvest is formed as the pH and conductivity are decreased. Development and optimization of TFF is crucial for the nonaffinity purification platform not only for the optimal loading condition on CEX, but also for determining the cost efficiency of the purification schemes.
Binding to CEX resin requires below-neutral pH and low conductivity for HuMAbs. The lower the pH, the higher the binding capacity, thereby, reducing capture chromatography cycles. However, the low pH requirement can lead to challenges because of gradual precipitation during diafiltration. Therefore, assessing the influence of several process parameters is essential to optimize the TFF step during primary recovery.
Primary recovery TFF plays a more significant role in the non-affinity scheme than in the affinity scheme not only because it is a batch-volume reduction and feed conditioning step for the capture, but it is also a partial purification step.
During the primary recovery TFF, the clarified harvest bulk is concentrated several fold, which will reduce batch processing time on capture chromatography especially for high-binding CEX resins. With more high-binding CEX resins competing in the market, development scientists have the luxury of evaluating and choosing the highest binding resins with high processing flow rates.3 Even for the resins that can be operated at very high flow rate, the load time would be significant without volume reduction at primary recovery step.
During TFF, the diafiltration endpoint for pH and conductivity are typically kept equivalent to those at the CEX load to avoid further dilution or titration between the primary recovery operations and capture chromatography step. Usually, the diafiltration end-point is when the conductivity and pH of the retentate reach those of the diafiltration buffer and it typically takes about five diafiltration volumes (DV), which involves most of the processing time in this unit operation. Instead, the diafiltration time can be reduced if the diafiltration buffer with lower pH and conductivity than the end-point requirements is used. For example, if the load for CEX is set at pH 5.9 and 4.3 mS/cm, it takes only 3 DVs by using a diafiltration buffer of pH 5.7 and 3.5 mS/cm (Figure 1) instead of 5 DVs with a diafiltration buffer of pH 5.9 and 4.3 mS/cm. Because it takes a longer time for 1 DV as flux decreases towards the end of diafiltration, buffer with lower pH and conductivity significantly shortens the diafiltration time (~114 minutes in Figure 1). Buffer costs can be reduced by 40% when diafiltration is decreased from 5 to 3 DVs which translates into savings of process time as well as raw material.
Figure 1. Reduction of diafiltration time by using a lower pH and conductivity buffer
Primary recovery TFF is also a partial purification step. As the pH moves lower during diafiltration, Chinese hamster ovary (CHO) host cell contaminants, especially DNA, gradually precipitate out. As the pH passes through the wide pI range of host cell proteins, it causes gradual precipitation of these contaminants. CHO DNA is significantly reduced as the diafiltration progresses, from starting ~106 pg/mg to ~103 pg/mg at the end of the primary recovery step, while CHO host cell proteins (HCP) are reduced at a modest level (Figure 2). Figure 3 shows a robust purification step for CHO DNA reduction for different Hu-MAbs. DNA partial purification at TFF generally benefits non-Protein A purification schemes, especially for high cell density and high titer cell cultures. This also prolongs capture cycle lifetime because of cleaner and uniform feed.
Figure 3. Reduction of CHO host cell DNA during primary recovery (PR) of different HuMAbs
HuMAbs in the diafiltered buffer condition at relatively lower pH also exhibit better stability than the clarified cell culture supernatant (Figure 4). TFF bulk typically has a three- to four-week storage time, while clarified bulk can last only several days. The prolonged stability could be because of the removal of proteases from the cell culture supernatant during buffer exchange. This could be beneficial because the cell densities tend to increase for hightiter cell culture processes and more proteases are produced during the cell culture and harvest operations. Also, the secondary modifications of proteins, such as deamidation and isomerization of antibodies, are better controlled at the lower pH as a result of diafiltration compared to the high pH of cell culture supernatant.4 This provides a convenient in-process hold point, which can potentially benefit flexible manufacturing schedules in commercial scales.
Figure 3. Reduction of CHO host cell DNA during primary recovery (PR) of different HuMAbs
The primary recovery TFF processes complex, clarified cell culture bulk, and diafiltration can result in enhanced precipitation and gel layer formation, thereby influencing the flux rate. Optimizing the operating conditions to maximize the flux and minimize the operating time should consider several process parameters. Choosing a suitable diafiltration buffer and a TFF membrane is essential to decrease operating time. Testing load protein amount and concentration fold of the feed, as well as optimal combination of inlet and retentate pressures, is important to define the maximized flux rate and optimized throughput.
Figure 4. Monoclonal antibody stability in clarified cell culture bulk versus primary recovery (PR) tangential flow filtration bulk
The foremost step in optimization is to identify a suitable buffer system for TFF because the chemical composition of the buffer plays a key role in product recovery. Several aspects should be considered when choosing an optimal buffer system. First, the diafiltration buffer should be compatible with the product during exchange without affecting product recovery. Second, the buffer needs to be compatible with the capture load conditions. For instance, in the case of a HuMAb process, the best binding condition for CEX was found to be sodium phosphate buffer with sodium chloride. However, the process performance at TFF stage was compromised by using this CEX binding buffer as diafiltration buffer, which resulted in poor product recovery. To accommodate best conditions for both TFF and CEX, diafiltration can be performed in a phosphate buffer without NaCl to prevent product loss, and added before binding to CEX.
Figure 5. Diafiltration flux curves from different tangential flow filtration membranes
Performance differences could exist in membranes from various manufacturers probably because of variation in pore size distribution even when they are made up of the same material and with similar molecular weight cut-off (MWCO). To identify the most suitable membrane for HuMAbs, evaluation of three membrane sources and two different pore sizes, 30 kD versus 50 kD, with similar process streams were performed (Figure 5). Better recyclability and consistent flux were noted with 50 kD membranes, whereas with 30 kD, the second cycle performance was compromised by decreased flux rate. However, similar product recoveries were obtained by both 50 kD and 30 kD membranes. Among three different membranes from different suppliers, one of them outperformed for process consistency during two cycles in terms of process flux rate.
Figure 6. Concentration and diafiltration time profiles versus load volume
A series of studies on the load volume were performed and the results are plotted in Figure 6. At constant concentration factor, as the load volume increases from 40 L/M2 to 200 L/M2 , the concentration and diafiltration times increase. Though the overall processing time is increased as load volume is changed from 40 L/M2 and 200 L/M2 , even at the highest load amount, TFF operation could be completed by 6.5 hours. This translates into a five-fold lower membrane cost for processing the same amount of material.
Figure 7. Concentration and diafiltration time profiles versus concentration fold
As concentration fold increases from 5X to 15X, there is a moderate increase in concentration time. However, the major impact is seen on diafiltration step, which occupies maximum TFF operation time (Figure 7). In fact, the diafiltration time at 10X to 15X is less than half of 5X concentration fold. This results in decreasing membrane and buffer usage and processing time, affecting overall raw material and labor costs.
The permeate flux rate can decrease significantly, especially during the diafiltration step because of gradual precipitation of host cell contaminants from concentrated cell culture supernatant. Therefore, range studies of feed and retentate pressures are essential for smooth TFF operation. The experiments examined the effects from changes of transmembrane pressure (TMP) as well as ΔP, which is the pressure difference between feed and retentate (Figures 8 and 9).
Figure 8. Concentration and diafiltration time profiles versus ÎP
When TMP was kept constant (15 psi), the increase of ΔP from 20 psi to 30 psi prolonged the concentration time very moderately, while the diafiltration time was reduced significantly from 4 to 2.5 hours.
Figure 9. Concentration and diafiltration time profiles versus transmembrane pressure
However, under constant ΔP (20 psi), elevating TMP reduced the concentration time and had very little impact on diafiltration time.
Figure 10. Cost comparison of disposable versus reusable tangential flow filtration membrane
By optimizing ΔP and TMP ranges, an additive effect on overall TFF process time reduction can be obtained.
Fouling of TFF membrane by gel layer formation from complex composition of cell culture is common during the diafiltration stage. Normalized water permeability (NWP) is monitored at the beginning and end of each cycle to evaluate the cleanliness of the membrane and performance consistency. Typically, NWP decreases sharply during the first couple of cycles (10–20%) and stays in the range of 60–80% in subsequent runs. Using NWP as the sole evaluation criteria can terminate membrane lifetime earlier than necessary. NWP can be used as an early indication of membrane modification, and is not necessarily directly related to process performance.5 Further, membrane performance can be measured by monitoring fluxversus-time curve, process recovery, process time, and product purity.6 The TFF membrane life can be at least 10 cycles in a typical HuMAb process as assessed by the flux curve, process time, and product recovery in a scale-down TFF system. In this context, an alternate approach can be to use a single-use membrane, which can reduce or eliminate cleaning and cleaning validation.7,8 Fully disposable crossflow systems are available in the market, which require extra costs for disposable adapter plates, pressure gauges, a flow meter, and a conductivity meter.7 But savings are more from eliminating cleaning validation and reductions in cleaning solutions, labor cost, membrane storage, process waste disposal, and analytical requirements. Disposable units will have significant cost advantage, especially when the membrane cycle usage is limited (Figure 10).
Even though primary recovery TFF adds a unit operation to the whole process, this step results in significant batch volume and overall capture cycletime reduction. Loading time reduction is advantageous for high-binding resins where bed volumes can be relatively smaller. For example, for a 5,000-L batch, the unit operation includes one TFF and two cycles of cation exchange versus six cycles of Protein A (Table 1).
Table 1. Comparison of operation cycles for nonaffinity and affinity schemes
As primary recovery TFF is a partial purification step, it aids in reducing the demand of host cell contaminant removal during CEX chromatography. These two-step operations provide efficient purification fold. Primary recovery TFF with ~90% or above process step recovery, along with CEX capture, were comparable to Protein A schemes in overall process recovery.
Assuming the polishing steps are similar for both CEX and affinity schemes, which is the case for HuMAbs, there are more cost advantages for processing the same amount of HuMAbs for a nonaffinity scheme over a protein A scheme. By developing optimal process conditions, the binding capacity of current leading cation exchange resins can reach several-fold over Protein A resins (Table 1). With a much cleaner load on the cation exchange, the resin lifetime is well maintained to easily reach more than 100 cycles. Cation exchange capture also provides hidden advantages if there is a need to change the resin because of longer schedule interruption during product campaigns or an unforeseen new resin need during storage in large-scale manufacturing.
An ion exchange purification process for HuMAbs provides an alternative approach to traditional Protein A schemes, and is comparable in terms of process consistency, time, product recovery, and quality.
Capture load conditioning for cation exchange is performed by a TFF operation, which can provide several advantages, such as batch volume reduction and partial purification. Furthermore, the TFF bulk provides a better environment and builds flexibility as a process storage intermediate during manufacturing. Different process parameters that seem to influence the concentration and diafiltraion stages of primary recovery are discussed to design efficient and scalable nonaffinity purification scheme for HuMAbs.
JUE (MICHELLE) WANG is senior manager, TIMOTHY DIEHL is a scientist, MARK WATKINS-FISCHL is a scientist, DEBORAH PERKINS is a scientist, and DEENA AGUIAR is a scientist, all of Purification Process Development, and ALAHARI ARUNAKUMARI is the senior director of Process Development, all at Medarex, Bloomsbury, NJ, 908.479.2451, email@example.com
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