ATPE, aqueous two-phase extraction; ATPS, aqueous two-phase system; BE, back-extraction; BP, bottom phase; BPE, concentrated bottom phase solution (extraction step); CHO, Chinese hamster ovary; CS, cell supernatant; HCP, host cell proteins; IgG, immunoglobulin G; IR, removal of impurities; LLE, liquid–liquid extraction; MSB, mixer-settler battery; MW, molecular weight; PF, purification factor; PHCP, purity in terms of host cell proteins; PProtein, purity in terms of total protein; PTotal, purity determined by size-exclusion chromatography; SEC, size-exclusion chromatography; TP, top phase; TP1, PEG-rich top phase concentrated solution (extraction step); TP2, salt-rich top phase concentrated solution (extraction step); YBE, yield of the back-extraction step; YWashing, yield of the washing step
An aqueous two-phase extraction (ATPE) process based on a PEG/phosphate system was developed for the capture of human immunoglobulin G and successfully applied to a Chinese hamster ovary and a PER.C6® cell supernatant. A continuous ATPE process incorporating three different steps (extraction, back-extraction, and washing) was set up and validated in a pump mixer-settler battery. Most of the higher molecular weight cell supernatant impurities were removed during the extraction step, while most of the lower molecular weight impurities were removed during the subsequent steps. A global recovery yield of 80% and a final protein purity of more than 99% were obtained for the IgG purification from a CHO cell supernatant, representing a 155-fold reduction in the protein/IgG ratio. For the purification of IgG from a PER.C6® cell supernatant, a global recovery yield of 100%, and a host cell protein purity were attained, representing a 22-fold reduction in the host cell protein/IgG ratio. These results, thus, open promising perspectives for the application of the developed ATPE process as a platform for the capture of antibodies. In fact, this new process has shown the ability to successfully recover and purify different antibodies from distinct cell culture supernatants. This technology can also overcome some of the limitations encountered using the typical chromatographic processes, besides inherent advantages of scalability, process integration, capability of continuous operation, and economic feasibility.
The high cost of the currently used downstream technologies is a key problem in many present day bioengineering processes. Recent advances in biotechnology have led to an increase in productivity in the upstream processing of many biopharmaceuticals. As a result, the process development efforts have shifted toward the development of efficient product recovery and purification steps [1, 2]. The typical platform approach used in the downstream processing of mAbs includes clarification, concentration, selective purification steps, and virus clearance . The selective purification steps usually comprise the mAbs adsorption to a protein A resin, followed by two further chromatography steps which remove host cell proteins (HCP), DNA, aggregates, any leached protein A and provide an adequate level of overall viral removal [3, 4]. Although chromatographic separations have been the workhorse of most purification processes, several limitations have been pointed out, such as (i) batch operation, (ii) low capacity, (iii) complex scale-up, (iv) time-consuming and high pressure packing processing, (v) slow intraparticle diffusion, (vi) low chemical and proteolytic stability and consequent contamination of the final product, and (vii) the high cost of the resins [5–9]. The replacement of some of these chromatographic steps by non-chromatographic alternatives, with high capacity and throughput, has been then suggested [7, 10–12].
Aqueous two-phase extraction (ATPE) is a potential and promising non-chromatographic alternative, which can combine a high recovery, selectivity and biocompatibility with an easy scale-up and continuous operation mode [10, 13, 14]. Recently, the economic and environmental sustainability of an ATPE-based capture process for antibodies purification has been evaluated and compared to the currently established platform using protein A chromatography. Aqueous two-phase system (ATPS) has shown to be considerably advantageous in terms of process economy and operation, especially when processing high titer cell culture supernatants . This liquid–liquid extraction (LLE) technology has been successfully reported in recent times as a primary stage unit operation in the downstream processing of several biological products including therapeutic proteins, such as antibodies [4, 10, 11, 15–24]. Its technical feasibility at large scale has been, however, demonstrated for the downstream processing of just few biological products [25–27]. This may be attributed, on one hand, to the limited knowledge of the mechanism of solute partitioning in ATPS, and in another hand, to the disadvantages of batch operations as well as to the difficulties associated with the implementation of ATPE processes in a continuous mode of operation . Indeed, in most of the few ATPE processes described in the literature, the equipment assembly (agitated vessel + centrifuge) used provides only one theoretical stage and is expensive at large scale .
The technical feasibility of a multi-stage equilibrium ATPE of human antibodies from a Chinese hamster ovary (CHO) cells supernatant has been reported before by our group . Significant improvements in both recovery yield and purity were observed when compared to a single-stage extraction step and it was possible to completely purify immunoglobulin G (IgG) from the higher molecular weight (MW) impurities and partially from the lower MW ones . In the present research work, a continuous ATPE process including extraction, back-extraction (BE) and washing steps has been developed and validated for the purification of human antibodies from CHO and PER.C6® cell supernatants (CSs). Different strategies for the removal of the higher (extraction steps) and the lower (BE and washing steps) MW CS impurities is firstly investigated as a single step and later as a multi-stage procedure. A continuous ATPE process is, thus, developed enabling the integration of this unit operation as a capture step in the downstream processing of antibodies.
2 Materials and methods
PEG with MW of 3350 Da, potassium phosphate dibasic anhydrous (K2HPO4), sodium phosphate monobasic anhydrous (NaH2PO4), and sodium chloride (NaCl) were purchased from Sigma (St. Louis, MO, USA). All polymers were used without further purification. Human IgG for therapeutic administration (product name: Gammanorm) was obtained from Octapharma (Lachen, Switzerland). All other chemicals were of analytical grade.
Different cell culture supernatants have been used in this research work. The CS 1 consisted of a CHO CS containing immunoglobulin G1 (IgG1) directed against a human surface antigen and was produced and delivered by Excellgene (Monthey, Switzerland). An Excellgene proprietary serum-free medium containing only one protein was used for production, and phenol red was added to the medium as a pH indicator. According to the ELISA determinations performed by Excellgene, the IgG titer was very low and was spiked with IgG from Gammanorm (pI of about 9 ) to final titers of 0.56 and 1.58 g/L. The CS 2 consisted of a PER.C6® CS containing IgG1 (pI of about 8.5) and was provided by DSM Biologics (Groningen, The Netherlands). The IgG titer determined by protein A affinity chromatography was about 0.5 g/L.
2.2 Aqueous two-phase extraction
2.2.1 Single-stage extraction
The bottom (BP) and top phase (TP) of single-stage ATPSs were prepared separately by weighing the appropriated amounts of components from stock solutions of 50% w/w PEG 3350 Da, 40% w/w phosphate buffer and solid NaCl. Phosphate buffer at pH 6 was prepared using a NaH2PO4/K2HPO4 mass ratio of 1.27. The BP (BP extraction) was composed of 14.6% w/w phosphate pH 6, 0.01% w/w PEG, 10.6% NaCl and 25% w/w CS, and was prepared by mixing a BP concentrated solution (named concentrated bottom phase solution (extraction step) (BPE)) with the cell culture supernatant (CS) in a ratio of 3:1, respectively. The TP (TP extraction) was composed of 29.1% w/w PEG, 2.7% w/w phosphate pH 6 and 8.5% NaCl, and was added to the ATPS in the form of two solutions, a 50% w/w PEG (PEG-rich top phase concentrated solution (extraction step) (TP1)) and a salt solution containing 6.75% phosphate and 21.25% NaCl (salt-rich top phase concentrated solution (extraction step) (TP2)).
Single-stage assays were carried out at bench scale (10 mL) and set up in 15 mL graduated centrifuge tubes. The different working solutions (BPE, TP1, TP2, and CS) were weighted into the test tube and mixed in a vortex (Ika, Staufen, Germany). The systems were, afterwards, incubated at room temperature for about 2 h and centrifuged at 3000 rpm for 10 min to allow a complete phase separation (Eppendorf, Hamburg, Germany). Phase volumes were determined and samples from the top and bottom were taken for phase composition analysis. All samples were diluted at least six times and analyzed against blank systems containing the same phase composition, but without the supernatant feed stock.
2.2.2 Multi-stage extraction
The multi-stage equilibrium extraction trials were carried out in a pilot scale cascade of pump mixer-settlers operating counter-currently as depicted in Fig. 1A. Each pump mixer-settler was composed of a 15 mL pump mixer with a centrally located mixing impeller and a glass settler (48 mm × 171 mm; liquid height inside the settler was 43 mm and the length of the settling part was 110 mm). In order to enhance phase separation, the dynamic part of the settlers (length equal to 61 mm) contained a metal mesh coalescer. The BP and TP were prepared separately according to the procedure described for the single-stage extraction experiments. However, the loading of cell culture supernatants represent 30% v/v of the BP and approximately 20% v/v overall. The flow rates (F) used are given in Table 1. The stirring was performed with the same global electric motor for all the pump mixers, and was usually set at about 250 rpm.
The settlers were first half filled with TP. The solution BPE and the CS were then fed into the pump mixer-settler battery (MSB). When the BP was already in the middle of the cascade, the TP solutions (TP1 and TP2) were started to be fed counter-currently into the cascade. The pump mixer-settler cascade was run continuously until steady state was reached, i.e. until the liquid hold-up in the MSB had been replaced at least four to six times and the outlet flow rates and solute concentration were constant after successive measurements (circa 10 h). Phase volumes were determined, and samples from the TP and BP of each stage were taken for composition analysis.
|Trial||Cell supernatant||Extraction||Back-extraction (N = 1)||Washing (N = 3)|
|N||FBPE inb)||FCS in||FTP1 inc)||FTP2 ind)||FBP BE ine)||FTP Washing inf)|
|1||CHO CS (0.6 g/L IgG)||5||445||192||160||103|
|2||CHO CS (0.6 g/L IgG)||6||450||195||202||131||1011||1490|
|3||CHO CS (1.5 g/L IgG)||6||410||185||205||130||485||560|
|4||PER.C6® CS (0.5 g/L IgG)||6||405||171||172||116||462||665|
2.2.3 Back-extraction (BE)
IgG from the extraction PEG-rich phase was back-extracted into a new salt-phase (BP BE) using different (i) phosphate BE solutions, (ii) mass ratios between the extraction TP and the phosphate BE solution, and (iii) pH values. Batch and continuous BE steps where set-up in 15 mL graduated centrifuge tubes and in a pump mixer-settler, respectively. The mixing of the phases, respective separation and analysis was performed according to the procedure described above for both single and multi-stage extraction steps.
IgG from the BE phosphate-rich phases was washed by adding different PEG 3350 washing solutions (TP washing) at different concentrations and mass ratios between the washing solution and the BP outlet of the BE step. Batch and continuous washing steps were set-up in 15 mL graduated centrifuge tubes and in a cascade of pump mixer-settlers, respectively. The mixing of the phases, respective separation and analysis was performed according to the procedure described above for both single and multi-stage extraction steps.
2.3 Analytical methods
2.3.1 Phase characterization
The BP and TP from the selected stages were characterized in terms of composition according to the procedure described by Rosa et al. .
2.3.2 Protein A affinity chromatography
The concentration of IgG in both TP and BP was determined by protein A affinity chromatography (Uppsala, Sweden) using a Poros protein A affinity column from Applied Biosystems (Foster City, CA, USA). Samples from both phases were diluted at least six times in a sample buffer containing 0.05% w/v Tween 80 and 150 mM NaCl in 10 mM sodium phosphate buffer at pH 8.5. Adsorption of IgG to the column was performed at 1 mL/min for 0.9 min with 10 mM sodium phosphate buffer at pH 8.5 containing 150 mM NaCl. Elution was, afterwards, performed for 1.6 min in 12 mM HCl with 150 mM NaCl. Absorbance was monitored at 280 nm.
2.3.3 Size-exclusion chromatography (SEC)
The total purity of both TP and BP from the ATPSs containing the CHO CS was evaluated by SEC using a TSK-GEL Super SW3000 column (30 cm × 4.6 mm id, 4 μm) and a TSK-GEL super SW guard column (3.5 cm × 4.6 mm id) from Tosoh Bioscience (Stuttgart, Germany). The column was equilibrated with 50 mM phosphate buffer, pH 7.0 with 300 mM NaCl and run in isocratic mode at 0.35 mL/min. Absorbance was monitored at 215 nm.
The total purity of the samples containing the PER.C6® CS was assessed using a TSK-GEL G3000SWXL column (30 cm × 7.8 mm id, 5 μm) and a TSK-GEL super SW guard column (4 cm × 6 mm id) from Tosoh Bioscience (Stuttgart, Germany). The column was equilibrated with 100 mM sodium phosphate buffer, pH 7.0 containing 100 mM sodium sulfate and run at 1 mL/min. Absorbance was monitored at 280 nm.
2.3.4 Total protein quantification
The total protein content in both TP and BP from the ATPSs containing the CHO CS was determined by the Bradford assay  using a Coomassie reagent supplied by Pierce (Rockford, IL, USA). To avoid interference from phase components, samples were diluted at least six times and analyzed against blanks containing the same phase composition but without proteins. IgG from Octapharma was used as a standard for protein calibration.
2.3.5 Host cell protein (HCP) quantification
A sandwich ELISA, an in-house assay from DSM Biologics, was used to determine the concentration of HCP in both TP and BP from the ATPSs containing the PER.C6® CS, using microplates coated by adding rabbit anti-HCP antibodies (1000-folded diluted in PBS). The coating occurred overnight at 2–8°C, after which the wells were washed four times with washing buffer (0.1% Tween 20 in PBS). The plate was subsequently block for 1 h in a shaker incubator at 37°C with blocking buffer (0.4% BSA in PBS). The detection was based on reaction of biotinylated anti-HCP-antibodies with streptavidin horseradish peroxidase (HRP) conjugate. The HRP enzymatic activity was quantified using the substrate tetramethylbenzidine (TMB). The enzymatic reaction was stopped with 2 N sulfuric acid an absorbance was measured at 450 nm.
2.4 Extraction performance parameters
In order to evaluate the extraction performance of the extraction, BE and washing steps, different parameters were used. These include (i) the volume ratio, VR, defined as the volume (or flow rate) ratio between the TP and BP; (ii) the recovery yield of IgG in the extraction step (yield of the extraction step (YExtraction)) defined as the ratio between the mass of IgG in the TP from extraction and the total mass of IgG added to the extraction system multiplied by 100; (iii) the recovery yield of IgG in the BE step (yield of the back-extraction step (YBE)) defined as the ratio between the mass of IgG in the BP from BE and the total mass of IgG added to the BE system multiplied by 100; (iv) the recovery yield of IgG in the washing step (yield of the washing step (YWashing)), defined as the ratio between the mass of IgG in the BP from washing and the total mass of IgG added to the washing system multiplied by 100; (v) the protein and HCP purity, purity in terms of total protein (PProtein) and purity in terms of host cell proteins (PHCP), calculated by the ratio of IgG to total protein concentration, determined by Bradford or HCP analysis, respectively; (vi) the total purity, purity determined by size-exclusion chromatography (PTotal), determined by the ratio of the area of the IgG peak to the total area of the chromatogram subtracted by the total area of the blank phase chromatogram; (vii) the purification factor (PF), calculated by the ratio of the final purity in the phase to the initial purity in the feed stock; (viii) the percentage of impurities removal, removal of impurities (IR), determined based on the SE chromatograms according to Eq. (1), where Aimpurities is the peak area of the impurities, V is the volume of the corresponding phase and DF is the dilution factor of the sample loaded into the column. The Aimpurities was determined by subtracting the area of the IgG peak and the total area of the corresponding blank phase to the total area of the sample phase chromatogram:
3 Results and discussion
In order to design and optimize an ATPS continuous process in a pump MSB, several extraction and BE single stage studies were performed, with the CHO and PER.C6® CSs.
3.1 Design studies
3.1.1 Extraction studies
A cross-current multi-stage extraction of IgG with ATPS has been previously simulated in test tubes in order to show the technical feasibility of a multi-stage equilibrium ATPE for the recovery of IgG . It was observed that using a PEG/phosphate ATPS containing 10% w/w NaCl, five stages and a volume ratio of 0.4, 89% of total IgG could be recovered in the PEG-rich phase highly purified from the high MW compounds and partially purified from the lower MW compounds .
Before designing a counter-current multi-stage extraction system, the performance of two single-stage LLE equipments, namely a mixer-settler and pump mixer-settler, was evaluated for the recovery of IgG and compared to a test tube trial under the same experimental conditions. The total volume of the batch bench scale trial was about 10 mL, while in the continuous mixer-settler and pump mixer-settler trials, the total flow rates were 565 and 748 mL/h, respectively, and the stirring speed was 250 rpm. No major differences were observed in the performance of all single-stage equipments tested. Thus, the recovery of IgG was neither influenced by the scale-up of the system, factor of about 75, nor by the decrease of the residence time in the mixer from about 2 h (test tube) to 1.2 min (pump mixer-settler). Moreover, given that a good phase separation was achieved using a gravity settler containing a metal mesh as a coalescer, it is also evident that is possible to have a continuous process without the use of any centrifugation step. These results, hence, open the possibility to use typical LLE contactors (e.g. extraction columns), usually characterized by very short residence times, for this type of extraction systems without losses in efficiency.
3.1.2 Back-extraction studies
Previous research studies have revealed that most of the low MW CHO CS impurities partitioned mainly to the PEG-rich TP, independently of the NaCl concentration, while IgG partitioned mainly to the BP under low NaCl concentrations . Small scale single-stage BE studies were thus conducted in order to evaluate the possibility of separating IgG from the lower MW impurities present in the CS. In this step, a fresh phosphate phase is added to the TP from the extraction step in order to back-extract IgG to a phosphate-rich BP. Several volume ratios (0.4–10.5), concentrations of phosphate (10, 12, and 14%) and pH values (6 and 8) have been tested (Table 2).
|Back extraction systems||Phosphate buffer (%) (w/w)||MR TP E/buffer||VR||YBE (%)||PTotal (%)||PProtein/PHCP (%)||IR (%)|
|CHO cells supernatant ([IgG] = 0.56 g/L; PTotal = 18%; PProtein = 38%)|
|BE1||10% (pH 8)||1:1||1.8||99||64||68||92|
|BE3||1:3||No phase separation|
|BE4||10% (pH 6)||1:1||No phase separation|
|BE6||12% (pH 6)||1:1||2.4||83||61||66||91|
|BE8||1:3||No phase separation|
|BE9||14% (pH 6)||1:1||1.4||95||58||71||90|
|PER.C6® cells supernatant ([IgG] = 0.43 g/L; PTotal = 4%; PHCP = 42%)|
|BE12||14% (pH 6)||1:3||0.4||96||19||nd||85|
The best performing BE system in terms of IgG purification from the CHO CS was the system with the highest volume ratio (10.5), obtained using 10% phosphate buffer at pH 6.0 (BE5). More than 95% of all impurities were removed, however, at the expense of a very low BE recovery yield (34%). The high NaCl concentrations determined for this BE system (7.4 and 6.5% in the BP and TP, respectively) together with the high volume ratio used may explain why such a low recovery yield was obtained for this step, as previous studies have shown that high NaCl concentration was one of the driving forces for the IgG extraction into the PEG-rich TP [4, 11, 15, 19, 26].
Higher BE recovery yields for IgG purification from the CHO CS were obtained for lower volume ratios, which led, however, to lower purities. In fact the best performing BE system in terms of IgG recovery yield (BE11) had a much lower volume ratio (0.4) and thus a considerably lower NaCl concentration (2.15% in both phases). This system was obtained using 14% phosphate buffer at pH 6.0 and allowed a recovery of all IgG, with 90% of the total impurities being removed. Similar results have also been obtained for the IgG recovery from the PER.C6® CS.
The selected BE conditions should aim at a higher recovery yield than higher purity. Thus 14% phosphate buffer at pH 6.0 was selected for further use but in order to achieve a higher IgG enrichment and IR, a further extraction step should be considered (Section 3.1.3).
3.1.3 Washing studies
In order to achieve a higher IgG enrichment and IR, an additional extraction step was developed. The BE phosphate-rich phase was then further concentrated and washed out from the remaining impurities by adding different PEG 3350 stock solutions at different concentrations (30 and 35%) and volume ratios (0.7–14.3), as described in Table 3. Although the recovery yield did not seem to be dependent on the volume ratio (ranging between 84 and 100%), lower purities were obtained when lower volume ratios were used. The best performing washing systems for the IgG purification from a CHO CS was obtained using 35% of PEG with a volume ratio of 7.9, which allowed the recovery of 92% of IgG and removal of 98% of total impurities.
|Washing systems||PEG 3350 (%) (w/w)||MR PEG sol./BP BE out||VR||YWashing (%)||PTotal (%)||IR (%)|
|CHO cells supernatant ([IgG] = 0.56 g/L; PTotal = 18%; PProtein = 38%)|
|W2a)||1.5:1||No phase separation|
|W6a)||3:1||No phase separation|
|PER.C6® cells supernatant ([IgG] = 0.43 g/L; PTotal = 4%; PHCP = 42%)|
Even better results were achieved for the IgG purification from the PER.C6® CS. The final total purity obtained after the washing step was increased by a factor of about five-fold when compared to the total purity obtained after the BE step. All IgG could be recovered in the BP from the washing step, and more than 99% of total impurities could be removed.
According to the IgG recovery yields obtained and also to the low NaCl concentration in this step (<1% w/w), it is, thus, predicted that two to three stages will be enough for the complete recovery of IgG in BP.
3.2 Process implementation
Based on the previously optimized conditions for the three different process steps: extraction, BE and washing (Section 3.1), a continuous counter-current ATPE process for the recovery and purification of antibodies was then designed and validated in a pump MSB, which was built according to the scheme illustrated in Fig. 1B. Different trials have been performed, where both PER.C6® and CHO CSs with different IgG titers have been processed (Table 1).
3.2.1 CHO cell supernatant
Before implementing the complete counter-current ATPE of IgG from the CHO CS under the previously predicted experimental conditions, the extraction step was optimized running two trials with different number of stages (N) and volume ratios, as summarized in Table 1 (Trials 1 and 2). After 10 h from the beginning of the trials, both outlet flow rates and solute concentration were fairly stable indicating that steady state had been reached and the first samples began to be collected. Considerable improvements in the extraction performance parameters were observed in contrast to a batch extraction in a test tube. About 25% higher recovery yields (86 and 90% obtained in Trials 1 and 2, respectively, against 62% in the test tube) and 15% higher protein purities (71 and 68% obtained in Trials 1 and 2, respectively, against 56% in the test tube) were attained when compared to a single-stage test tube trial, hence, stressing the advantages of using a multi-stage procedure. The results obtained were in agreement to the ATPE process scheme predicted previously by the authors . As higher recovery yields were obtained using six stages, further processing (BE and washing) was performed using the conditions described for the Trial 2.
Figure 2A and B depict the results obtained for the extraction performance parameters of the Trials 2 and 3 regarding the purification of a CHO CS performed under the experimental conditions described in Table 1. Global recovery yields of 80% were achieved and more than 98% of the total impurities have been removed. Final PFs, determined by SEC, of 5.6 and 2.6 (out of maximums of 5.7 and 2.7) were attained for Trials 2 and 3, respectively. Most of the higher MW compounds were removed during the first extraction step, while the lower MW compounds were mainly removed during the subsequent BE step and completely washed out in the last washing step. After the washing step, all impurities were removed, with the exception of the impurity with a retention time of 9.8 min for which a percentage of removal of about 85% was obtained in both trials. A significant reduction in the percentage of aggregates from 20% in the CHO CS to less than 4% in the BP from the washing step was also attained. A 155-fold reduction in the contaminant proteins/IgG ratio has been achieved using the described ATPE process and starting with an initial feed containing 3.0 × 105 ng contaminant proteins/mg IgG. This clearance is illustrated in Fig. 3 where the SEC chromatogram from the initial feed stock solution (Fig. 3A) and from the final BP from the washing step (Fig. 3B) are compared. The peak with a retention time of 11.08 min that remains in Fig. 3B corresponds to the blank phase peak (i.e. to the phase-forming compounds).
3.2.2 PER.C6® cell supernatant
A PER.C6® CS was also processed through counter-current ATPE under the experimental conditions summarized in Table 1 (Trial 4). Figure 2C depicts the results obtained for the extraction performance parameters. All IgG was recovered and purified from the PER.C6® CS with a percentage of total impurities removal of more than 99%. Final PFs of 2.1 (determined by HCP analysis) and 20.4 (determined by SEC) out of maximums of 2.2 and 20.9, respectively, were obtained. As observed for the purification of IgG from a CHO CS, a fractionation of higher and lower MW impurities was also observed during the PER.C6® CS aqueous two-phase processing. Most of the higher MW impurities, including the IgG aggregates, were removed during the first extraction step, while the lower MW impurities were mostly removed during the BE step and completely washed out during the washing step. A significant 22-fold reduction of the HCP/IgG ratio was also obtained after the ATPE process, when compared to the initial feed which contained 1.0 × 106 ng HCP/mg IgG. The SEC chromatograms of the initial feed stock solution (Fig. 4A) and of the final BP from the washing step (Fig. 4B) are compared in Fig. 4. The peak with a reten-tion time of 11.9 min that remains in Fig. 4B corresponds to the blank phase peak (i.e. to the phase-forming compounds).
Most of the HCP were removed during the BE step due to their preferential partitioning to the PEG-rich phase. In contrast, the contaminant proteins present in the CHO CS have been mostly removed during the extraction step due to their preferential partitioning to the phosphate-rich phase. This shows that the process is robust can handle changes in the initial feed stock impurities without compromising the purity of the final product.
An alternative LLE-based process for the capture of human antibodies from complex cell culture media has been successfully developed and validated in a typical LLE equipment (MSB). This alternative capture step is based on the ATPE principles and incorporates three different steps: (i) extraction, where most of the higher MW contaminants are removed, and (ii) BE, and (iii) washing, which allow not only the purification and separation of IgG from the lower MW contaminants and polymer-rich phase, but it will also enable the recycling of the polymer for future uses.
This ATPE process allowed the recovery of IgG from a CHO CS with a global recovery yield of 80% and a final total purity of 97% and proteins purity of more than 99%, corresponding to a total contaminants removal of more than 98%. A PER.C6® CS has also been processed using the developed alternative capture process. All IgG could be recovered with a final total purity of 97% and an HCP purity of 95%, leading to more than 99% of total contaminants removal. The obtained results, hence, open promising perspectives for the application of the developed ATPE process as a platform for the capture of antibodies. In fact, this new process has shown the ability to successfully recover and purify different antibodies from distinct cells culture supernatants, being robust enough to accommodate changes in the properties of the starting feed stock impurities. Besides inherent advantages of scalability, this technology can be held in continuous mode and can thus be used to develop a continuous downstream process. In this case ATPS could be combined with other continuous unit operations, namely with a stack-disk continuous centrifuge upstream and with the multicolumn countercurrent solvent gradient purification (MCSGP) process , downstream. In addition, the possibility and ability of ATPE to integrate clarification and partial purification in just one step, without the need of initial recovery steps for cells removal, should also be considered as a major advantage.
The authors acknowledge Dr. Olaf Mol from DSM Biologics (Groningen, The Netherlands) for allowing the execution of the experimental work regarding the PER.C6® cell supernatant at DSM Biologics facilities. This work was carried out within the European project AIMs (contract no. NMP3-CT-2004-500160) supported by funding under the Sixth Research Framework Programme of the European Union. P. A. J. Rosa acknowledges “Fundação para a Ciência e Tecnologia” for a PhD fellowship (BD 25040/2005). A. M. Azevedo acknowledges the programme “Ciência 2007” from the Portuguese Ministry of Education and Science.
The authors declare no conflict of interest.