METHOD AND SYSTEM OF PROTEIN EXTRACTION

Abstract

The present invention relates to a method and a system of extracting a protein with high yield from a protein-comprising precipitate, in particular immunoglobulin, from human or non-human origins, such as blood plasma.

Claims

1-32. (canceled)

33. A closed system for extracting a protein of interest from a precipitate, comprising: (a) a first tank configured to contain the precipitate in the form of a suspension having a first dilution factor; (b) a first filtration unit connected with the first tank and comprising a dynamic filter element, wherein the first filtration unit is configured to: receive the suspension from the first tank, produce, using the first filtration unit, a first permeate enriched with the protein of interest and a first retentate depleted of the protein of interest, and return the first retentate to the first tank; (c) a second tank connected with the first filtration unit and configured to receive the first permeate enriched with the protein of interest; and (d) a second filtration unit configured to: concetrate the first permeate in the second tank so as to produce a second retentate enriched with the protein of interest and a second permeate depleted in the protein of interest, and one or both of (i) return the second retentate to the second tank and (ii) return the second permeate to the first tank.

34. The closed system according to claim 33, wherein the precipitate is an intermediate product of an alcohol fractionation process.

35. The closed system according to claim 34, wherein the intermediate product is selected from Cohn Fraction I (Fr I), Cohn Fraction II+III (Fr II+III), Cohn Fraction I+II+III (Fr I+II+III), Kistler/Nitschmann Precipitate A (KN A), and combinations of KN A and one or more of Fr I, Fr II+III and Fr I+II+III.

36. The closed system according to claim 33, wherein the precipitate is a culture supernatant or a fermentation product.

37. The closed system according to claim 33, wherein the protein of interest is an immunoglobulin (Ig).

38. The closed system according to claim 33, wherein the dynamic filter element is a dynamic cross flow filter element.

39. The closed system according to claim 38, wherein the dynamic filter element is a rotational cross-flow filter element.

40. The closed system according to claim 39, wherein the rotational cross-flow filter element is configured to rotate at a rotating speed from about 600 rpm to about 1200 rpm.

41. The closed system according to claim 33, wherein the dynamic filter element is a rotating filter element comprising one or more filter discs.

42. The closed system according to claim 33, wherein the first filtration unit is equipped with rotating filter discs comprising ceramic membranes.

43. The closed system according to claim 33, wherein the first filtration unit is equipped with baffles configured for turbulence mixing of the content of the first filtration unit.

44. The closed system according to claim 41, wherein the filter disks are configured to have a tangential speed of about 1 to about 7 m/sec.

45. The closed system according to claim 33, wherein the first filtration unit comprises a pressure vessel.

46. The closed system according to claim 33, wherein the first filtration unit is configured to maintain a temperature in the first filtration unit of 2° C. to 25° C.

47. The closed system according to claim 33, wherein the closed system is configured to control flow velocities of the first permeate and the second permeate such that a constant product volume is maintained in the second tank.

48. The closed system according to claim 42, wherein a transmembrane pressure across the ceramic membranes is from 0.1 bar to 2.5 bar.

49. The closed system according to claim 33, wherein the dynamic filter element comprises a filtration membrane having an average pore size of 5 nm to 5000 nm.

50. The closed system according to claim 33, wherein the precipitate has a total protein concentration of about 0.5% w/v to 6.5% w/v.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

[0139] The following drawings are not necessarily drawn to scale, emphasis instead is generally being placed upon illustrating the principles of various embodiments. In the following description, various embodiments of the invention are described with reference to the following drawing:

[0140] FIG. 1 is a schematic flow chart overview of the system of the present invention, and is described in more detail below.

[0141] FIG. 2 shows the flow rate over time through the first filtration unit in the presence (squares) and absence (circles) of filter aid.

[0142] FIG. 1 illustrates a schematic flow chart overview of the system 100 and the method according to one preferred embodiment of the present invention. Protein-comprising precipitate e.g. in form of a suspension, or in form of a paste or precipitate is suspended with liquid e.g. buffer. The compositions and concentration of the buffer are in accordance with the above described method in order to generate a starting composition such as a suspension having a first dilution factor e.g. between 3 to 10 (1:3 to 1:10). The suspension is housed in a first tank 1. The suspension can be fed to a first filtration unit 5, through the pump 2, several type of pumps can be used (e.g. piston-; rotary-;centrifugal- and membrane pump) and flow-regulated valve 3 of a pipe 12. The first filtration unit 5 is equipped with a rotating hollow shaft to which the filter discs are mounted (the filtrate flows from the outside to the inside of hollow shaft). The first filtration unit 5 is further set up with height adjustable scrapers to keep the filter cake thickness constant and thus achieve constant filtrate flow. The desired filtration pressure is controlled and regulated by overflow valve (unfiltered suspension outlet). The filter discs used can be a ceramic membrane, depth filter layers and sintered porous metal filter discs. Once the vessel of the first filtration unit 5 is filled with the suspension, a continuous pressure extraction and separation can be started. The first filtration unit 5, which can comprise a pressure unit/vessel, is provided with suitable internal settings and conditions to simultaneously increase the extraction efficiency and filtration process. The extraction efficiency is increased through turbulence mixing in the unit 5 without having to involve a mixer. Nevertheless, it can be foreseen that an additional mixer may be provided to assist the extraction process by creating turbulences. Moreover, higher final dilution factor e.g. 40 or 70 disclosed in the present invention also increases the extraction efficiency, leading to high protein (e.g. IgG) yield. Of course, any other higher final dilution factor (higher than 70) can also be envisaged.

[0143] The filtrate flows through a flowmeter 6 installed on pipe (or channel) 14 and is collected in the second tank 7. The unfiltered suspension flows back through the regulated outlet 3 installed on pipe 13 in tank 1. When a defined volume in the second tank 7 is reached, the UF 8 concentration process can be started in the second filtration unit. The filtrate in the second tank 7 flows through pipe 15 into the ultrafiltration (UF) system 8. The transmembrane pressure is set such that the permeate flow rate 17 is identical or almost identical to that with the first filtrate flow rate in pipe 14. The permeate of the UF system 8 flows through pipe (or line or channel) 17 back to the first tank 1, whereas the retentate of the UF system (=concentrated protein) flows through pipe 16 back to the second tank 7.

[0144] In accordance with the invention, the first process unit 5 is provided with one or more rotating filter discs comprising one or more of the first filter element for turbulence mixing of the content of the first process unit 5 for producing the first retentate and the first permeate. The first retentate can be fed back to the first tank 1 through a channel 13 via a control valve 3 whereas the first permeate can be fed to a second tank 7 via another channel 14. The first filter element can be a filtration membrane which is based on a ceramic material, having a pore diameter of between about 5 nm to 5000 nm, preferably between 20 nm to 100 nm or more preferably between 30 nm to 80 nm. It can also be foreseen that inorganic membranes or any other suitable membranes could also provide a similar effect as the ceramic based membrane. The first filtration unit 5 may be supplied with a pressure control device 4 such as a manometer in order to regulate the pressure within. Similarly, a flowmeter 6 can be installed in the system of the present invention for measuring the flow rate of the suspension or solution.

[0145] Feedstream from the second tank 7 can then be fed to a second filtration unit 8 through a channel 15 for a second separation process to be carried out. The second separation process can be a continuous concentration process (e.g. UF). The second filtration unit 8 is provided with one or more second cross flow filter element/s, wherein the second cross flow filter element can comprise an ultrafiltration membrane having an average molecular weight cutoff value of less than 50 kD. However, the membrane can also be less than 10 kD or more preferably less than 5 kD. The ultrafiltration membrane therefore produces a second retentate which is channeled back to the second tank 7 through a channel 16 whereas the second permeate is fed to the first tank 1 via a channel 17. To this end, it is noted that the pressure of the second filtration unit 10 can be regulated during concentration step (ultrafiltration) such that the flow velocity of channels 14 and 17 are substantially equal.

[0146] In the following description, a detailed description of the methods according to the present invention are outlined in several experimental examples.

EXAMPLE 1

[0147] According to the present invention, immunoglobulin G was extracted through a continuous extraction and separation process in a first process unit.

[0148] An amount of 1 kg protein-comprising precipitate (Precipitate A) was dissolved in 10 mM sodium acetate, 10 mM phosphate and 2 mM citric acid buffer for 30 minutes to give a first dilution factor of 10 (i.e. 1 kg of the precipitate dissolved in 9 kg of buffer), wherein the pH of the suspension was about pH 4.6. The suspension was prepared in the first tank and was fed to the first process unit for a continuous extraction and separation process. The first process unit was provided with a rotation filtration element comprising a ceramic-based membrane disc, having a filtration membrane with an average pore size of 80 nm. The ceramic filter used was a Ceramic Filter Disc 152 which had a diameter Øo 152 mm/Øi 25.5 mm; thickness d=4.5 mm; and membrane surface area 360 cm.sup.2. The tangential velocity of the disc was approximately 7 m/s at 60 Hz (800 rpm). The average filtration rate was about 200 ml/min. During the continuous extraction and separation process, for each 200 ml filtrate (first permeate) collected in a second tank, 200 ml of buffer was returned to the first tank from a second permeate obtained from a second filtration unit (ultrafiltration). After 4 hours the filtration was stopped, wherein the predetermined protein concentration in the first tank was less than 0.1 g/L (equates to a final dilution factor of 31). A total amount of 3 kg of the filtrate was further concentrated (10 kD ultrafilter membrane) in the second filtration unit to 20 g/L.

[0149] A comparative experiment was performed according to a prior art method using depth flow filtration. The same lot of Precipitate A was used in this experiment, wherein the precipitate was suspended in 0.22 M sodium acetate buffer. A final dilution factor of 6 (i.e. 1 kg of the precipitate dissolved in 5 kg of buffer) was used for this experiment. The suspension was mixed for 4 hours prior to depth filtration. Finally, the filtrate was concentrated to 20 g/L, using an ultrafiltration membrane having an average molecular weight cutoff of 10 kD.

[0150] The results of the immunoglobulin G yield are shown in Table 3. The IgG yield obtained using the continuous extraction method was higher than the prior art method. As explained above, the extraction method ensures the precipitate is exposed to an increased volume of liquid (or final dilution factor). This is thought to shift the dissolution equilibrium in favour of increased extraction of immunoglobulin G from the precipitate material which could then be recovered in the first permeate of the first filtration unit. As shown in Table 3, the dynamic filtration system of the present invention enabled an increase in IgG yield of approximately 0.68 g/L plasma equivalent (PEQ) compared to the prior art method. This equates to about 10% of the IgG in each liter of pooled plasma.

TABLE-US-00003 TABLE 3 Comparison of IgG yield between current and new process at different process steps. (PEQ stands for plasma equivalent i.e. the amount of IgG from each liter of plasma) Control Present invention (prior art process) (new process) First dilution factor 6 (1:6) 10 (1:10) Resuspension time (h) 2-8 0.5  IgG yield (g/L PEQ) in 5.86 5.80 Starting composition (suspension) Final dilution factor 6 (1:6) 31 (1:31) Filtrate 5.53 Not applicable IgG yield (g/L PEQ) post Not applicable 6.20 continuous extraction and separation IgG yield (g/L PEQ) post 5.48 6.16 Ultra-concentration

EXAMPLE 2

[0151] In this example, IgG yields resulting from the use of different buffer compositions and different final dilution factors were compared using similar methods and equipment as described in Example 1.

[0152] As shown in Table 4, a final dilution factor of 6 was applied to the control (Sample A). This final dilution factor represents a common final dilution factor widely practiced in the art. For example, WO2016012803 (p. 15, line 30) suggests diluting by factors from about 1:2 to about 1:10. In contrast, the present invention provides a practical means to allow higher dilution ratios to be used. In the present example a final dilution factor of 40 was used for Samples B and C.

[0153] Buffer in Sample A comprises 0.22 M sodium acetate. Buffer in the Sample B comprises 5 mM acetate and 5 mM phosphate whereas the buffer in Sample C comprises 10 mM acetate and 10 mM phosphate. Both Samples B and C contained additionally 2 mM citric acid to maintain a constant pH after resuspension of the protein comprising precipitate. The pH of the starting composition in form of a suspension of all samples was approximately 4.8.

[0154] The final dilution factor for sample A of 6 was obtained by firstly dissolving approximately 1 kg of Precipitate I+II+III according to Cohn Method 10 in the buffer described above (0.22 M acetate; one part of precipitate and 5 parts of buffer; 1:6 wt/wt; precipitate:total). The suspension was mixed for 4 hours at an ambient room temperature. Thereafter, the suspension was filtered through a depth filter (0.2 to 0.45 μm, polypropylene), and finally ultra-concentrated through 10 kD membrane (Pellicon® 3) to 20 g protein/L.

[0155] A total amount of 1 kg protein-comprising precipitate was dissolved for Sample B as well as for Sample C in the above described buffer for 30 min to give a first dilution factor of 6, wherein the pH of the suspension was adjusted to about 4.6. The suspension was prepared in the first tank and was fed to the first process unit for a continuous extraction process. During the continuous extraction process, for each 100 to 200 ml filtrate collected in a second tank, 100 to 200 ml of buffer were returned to the first tank from the second permeate of the second process unit (ultrafiltration). The filtration process was stopped when the total dilution factor was about 40. The filtrate was further concentrated to 20 g/L in the second filtration unit using an ultrafiltration membrane having an average molecular weight cutoff value of 10 kD.

[0156] Six different lots were used for the experiments (comparing the same lot with each test buffer, respectively). The protein and immunoglobulin G yields were then compared. The yield results showed increases of 0.56 g immunoglobulin G per L PEQ (average) for both Samples B and C compared to Sample A (see Table 4).

TABLE-US-00004 TABLE 4 Comparison between prior art method (control) and methods used in the present invention post ultrafiltration concentration step. Control Present invention (prior art) (proposed method) Sample A B C Buffer Acetate & Acetate & Acetate Phosphate Phosphate (0.22 M) (5 mM; 5 mM) (10 mM; 10 mM) First dilution 6 (1:6) 6 (1:6) 6 (1:6) factor Final dilution 6 (1:6) 40 (1:40) 40 (1:40) factor IgG yield 4.38 ± 0.24 4.89 ± 0.18 5.04 ± 0.14 (g/L PEQ)

EXAMPLE 3

[0157] The impact of a different pH on the yield of immunoglobulin G, M and A and other impurities using the continuous extraction system of the present invention are demonstrated in this example. Two experiments were performed wherein IgG recoveries were compared using a citric acid buffer and a phosphate buffer. The protein-comprising precipitate used in this example was 1 kg Precipitate I+II+III derived from plasma treated with ethanol according to the Cohn Method 10 or according to the Kistler and Nitschmann method (1962, Vox Sang. 7, 414).

[0158] The lower pH sample was obtained by resuspending the above-described precipitate in a citric acid buffer (natrium citrate-citric acid) in order to give a first dilution factor of 5 (1:5) at a pH of 3.5 to 3.9. The suspension was stirred at 20° C. for 30 minutes.

[0159] The suspension was then transferred to the first tank which was subsequently fed to the first filtration process unit for a continuous extraction and separation process as described in Example 1. The first process unit was started as soon as the system was filled with the suspension. A first permeate/extract was produced from the first process unit, wherein the first permeate/first extract was collected in a second tank before it underwent an ultra-concentration step (second filtration unit). A second permeate depleted in the protein of interest obtained from the ultra-concentration step was fed back to the first tank. The continuous extraction and filtration process was stopped when the protein concentration in the first tank was less than 0.05 g/L and/or the final dilution factor was 40 (1:40).

[0160] The higher pH sample was obtained by resuspending the above-described precipitate in a phosphate buffer (disodium hydrogen phosphate Na.sub.2HPO.sub.4 and sodium dihydrogen phosphate NaH.sub.2PO.sub.4) in order to give a first dilution factor of 5 (1:5), and a pH of 8.0. The suspension was stirred at 20° C. for 30 minutes. Apart from the pH value, all other conditions and steps used in the higher pH sample were identical to low pH suspenion (as described above).

[0161] Tables 5 and 6 show the results after the suspension had undergone the ultra-concentration step in the second filtration unit.

TABLE-US-00005 TABLE 5 IgG, IgA and IgM yield at ultra-concentrated step. Citrate Phosphate pH 3.7 8.0 IgG (g/L PEQ) 6.38 6.42 IgA (g/L PEQ) 0.79 0.81 IgM (g/L PEQ) 0.48 0.50

[0162] The results demonstrate that the pH extraction conditions did not affect IgG, IgA or IgM yield. There were however effects observed in respect to other parameters with for example the low pH buffer conditions resulting in preparations with reduced levels of PKA and proteolytic activity. Such parameters can have a negative impact on the stability/quality of an immunoglobulin preparation. The parameters α1-Antitrypsin, α2-Macroglobulin, Transferrin, Albumin, Apo-Al, Ceruloplasmin, Haptoglobin, Fibrinogen, Fibronectin, Hemopexin and IgG-subclass distribution were determined by immunonephelometry assays. Phospholipid, triglyceride and cholesterol levels were determined by enzymatic test assays. Protein composition was performed by agarose gel electrophoresis. Molecular size distribution (Aggregate, Dimer, Monomer and Fragment) was determined by size exclusion chromatography. Determination of PKA and proteolytic activity were performed by chromogenic substrate assays.

TABLE-US-00006 TABLE 6 Impurity profile after ultra-concentration step Citrate buffer Phosphate buffer Impurities [g/L] [g/L] Alpha1-Antitrypsin 0.0931 0.0633 Alpha2-Macroglobulin 0.972 0.703 Transferrin 0.215 0.203 Albumin 1.32 1.29 Apo-Al 0.149 0.106 Ceruloplasmin 0.205 0.104 Haptoglobin 0.0763 <0.07 Fibrinogen 1.16 1.79 Fibronectin 0.052 0.030 Hemopexin 0.0617 <0.05 Phospholipid 0.2 0.25 Triglyceride 0.14 0.19 Cholesterol 0.19 0.28 Protein composition Percentage (%) Albumin 8.1 8.07 Alpha-/Beta-Globulin 16.8 23.9 Gamma-Globulin 75.1 68.1 Molecular size distribution Percentage (%) Aggregate 28.5 29.8 Dimer 6.5 6.5 Monomer 64.9 60.0 Fragment <0.1 3.7 IgG-subclass Percentage (%) IgG1 [%] 62.1 61.4 IgG2 [%] 27.9 30.3 IgG3 [%] 3.7 3.1 IgG4 [%] 6.3 5.2 Other parameter PKA [IU/mL] 1900 3300 Protease at Product pH [nkat/L] 46 97 Protease at Protease pH [nkat/L] 2407 10037

EXAMPLE 4

[0163] In this example IgG recovery from a precipitate was compared by i) dissolving the precipitate in a fixed volume of 220 mM sodium acetate (pH 4.8±0.2) resulting in a final dilution factor of 6 and recovering dissolved protein using depth filtration; ii) dissolving the precipitate in 220 mM sodium acetate (pH 4.8±0.2) and recovering dissolved protein using the continuous extraction process of the invention to achieve a final dilution factor of 31; and iii) using the continuous extraction process of the invention whereby the suspension in the first tank was continuously replenished with fresh buffer to achieve a second dilution factor of 31.

[0164] Part i: 1 kg of the same lot of precipitate was suspended in 5 kg of 220 mM sodium acetate (pH 4.8±0.2), using the same impeller mixer (ID 10 cm) to give a final dilution factor of 6 (i.e. 1 kg of the precipitate dissolved in 5 kg of buffer). The suspension was mixed for 8 hours. Prior to depth filtration, filter aid (FA=10 g/kg of Celpure C100, Advanced World Mineral) was added and mixed for 30 min. The depth filtration was performed using combined filter sheets (Polypropylene from Dolder CH, Cellulose, CH9 from Filtrox) in a filter press (20×20 cm frames; from Filtrox) at a maximum pressure of 2.5 bar. The filter area used was 3200 cm.sup.2. After the filtration was finished the post wash was started using 2.5 L of the resuspension buffer. This resulted in a total filtrate of 6.9 L and protein concentration of 18 g/L. Protein concentration was determined by Kjeldahl, Biuret and A280 assays. Finally, the filtrate was further concentrated to 20 g/L, using an ultrafiltration membrane having an average molecular weight cut-off value of 10 kD as described above. The yield in a final ultra-filtrate volume of 5.9 L and protein concentration of 20.7 g/L (Table 7).

[0165] Part ii: An amount of 1 kg of frozen protein-comprising precipitate in form of a precipitate containing around 100 g filter aid, from Kistler-Nitschman process (KN), was resuspended in 220 mM sodium acetate (pH 4.8±0.2) buffer for 30 min using an impeller mixer (ID 10 mm) to give a first dilution factor of 10 (i.e. 2 kg of the precipitate dissolved in 18 kg of buffer). The suspension in the first tank (20 L working volume) was pumped using a diaphragm pump at a flow rate of 1000 mL/min into the first dynamic filtration process unit. The process unit contained a double layer ceramic-based membrane disc (upper membrane layer 80 nm and lower layer 100 nm). The ceramic filter disc 152 (KERAFOL Keramische Folien GmbH, 92676 Eschenbach) had a diameter of 152 mm; thickness of 4.5 mm; and membrane surface area of 360 cm.sup.2. The tangential speed of the disc was approximately 7 m/s at 60 Hz (equivalent to 800 rpm). An average filtration rate was set to approximately 200 ml/min.

[0166] Once the first process unit was filled, the suspension was circulated for 10-15 minutes under constant pressure of 1 bar (range: 0 to 2 bar) using the overflow valve (which modulates the return flow of first retentate from the first process unit to the first tank). At this point the continuous extraction process was initiated with the transmembrane pressure (TMP) maintained between 0.5-1.5 bar. The first permeate was collected in a second tank (20 L working volume) at a flow rate 100-200 mL/min. The unfiltered retentate suspension flowed back to the first tank, at a flow rate of 800-900 mL/min, through the regulated outlet valve. When a defined volume (2000-4800 mL) was collected in the second tank (=filtrate tank), the ultrafiltration (second filtration unit) using a 0.2 m.sup.2, cut-off 10 kD, Ultracel®/or Biomax® filter (Milipore) was started. The transmembrane pressure (TMP: 0.8-1.5 bar) was set such that the permeate flow rate of the UF system was similar to the permeate filtrate flow rate (100-200 mL/min) to ensure a continuous extraction process. The permeate of the UF-second filtration unit was returned to the first tank at a flow rate of 100-200 mL/min. After 4 hours the filtration was stopped, wherein the predetermined value of protein concentration in the first tank (suspension tank=feed tank) was less than 0.1 g/L. This equates to a final dilution ratio of 1:31. The ultrafiltration process was continued until the protein concentration reached 20 g/L. During this final concentration, the second permeate was sent to waste. The final ultra-filtrate volume was 7.3 L at a protein concentration of 21.4 g/L (Table 7). The protein concentration was determined by the Kjeldahl assay.

[0167] Part iii: In a third part of this experiment, the first filtration unit was used as a stand-alone system (i.e. disconnected from second UF system). An amount of 1 kg of the same precipitate was resuspended in 5 kg to give a dilution factor of 6 (i.e. 1 kg of the precipitate dissolved in 5 kg of buffer), wherein the pH of the suspension was about 4.6 to 5.0. All other parameters for this experiment were the same as described in part ii) with the exception that fresh buffer was added to the first tank instead of permeate from the UF system. Once the first filtration unit was filled, the suspension was recirculated for 10-15 minutes under constant pressure of 1 bar (range: 0 to 2 bar) using the overflow return valve before starting the continuous extraction process. The filtrate was collected in a second tank (50 L working volume) at a flow rate 100-200 mL/min. The unfiltered suspension flowed back to the first tank, at a flow rate of 800-900 mL/min, through the regulated return valve. When a defined volume (2000-4800 mL) was collected in the second tank (=filtrate tank) fresh buffer was added to the suspension tank at a flow rate similar to that of the first permeate filtrate flow rate (i.e. 100-200 mL/min). After about 4 hours the filtration was stopped, wherein the predetermined value of protein concentration in the first tank (suspension tank=feed tank) was less than 0.1 g/L. The volume of collected filtrate was around 31 L at protein concentration of 4.8 g/L. This volume is equal to final dilution ratio of 1:31. The filtrate was further concentrated to a protein concentration of 20.6 g/L to give a final volume of 7.2 L (Table 7).

[0168] The results of the IgG yield are shown in Table 7. The IgG yield according to the present invention method (which involves a continuous extraction and filtration process) gave a higher yield than the prior art method. As explained above, by using the extraction and filtration method as disclosed in the present invention, a change in the concentration by increasing the volume (or final dilution factor) is achieved, whereby the dissolution equilibrium is shifted in favour of increased extraction of immunoglobulin G to achieve a higher overall yield.

TABLE-US-00007 TABLE 7 Comparison of IgG yield between current and new process at different process steps. (Initial dissolution time: is the mixing time prior to the start of the continuous extraction and filtration unit; PEQ stands for plasma equivalent i.e. the amount of IgG from each liter of plasma). Continuous Continuous recovery recovery Control process process Part i) Part ii) Part iii) Initial dilution factor 1:6 1:10 1:6  Initial dissolution time (h) 8 0.5 0.5 Volume of suspension (L) 6 10 6 Protein concentration in 24.3 10.9 22.1 suspension (g/L) after 8 h after 0.5 h after 0.5 h Final dilution factor 1:6 1:31 1:31 Process time for extraction 9 4.5 4.5 and filtration (h) Filtrate volume (L) including 6.9 Not applicable 31 post wash for the prior art Protein conc. (g/L) 18.0 Not applicable 4.8 Total protein (g) 123.6 Not applicable 148.8 Volume post UF (L) 5.9 7.1 7.2 Protein conc. post UF (g/L) 20.7 21.4 20.6 Protein yield post UF (g/L 12.1 14.9 14.5 PEQ) IgG yield post UF (g/L PEQ) 6.1 7.3 7.2

EXAMPLES 5, 6 AND 7

[0169] In these examples the impact of rotation speed of the rotating filter discs in the first process unit and the overall recirculation volume (final dilution factor) on the protein yield in the second tank was investigated. Table 8 provides an overview of the conditions used.

[0170] The results show that the higher the speed of rotation of the discs and the higher the recirculation volume the higher the extracted target proteins recovered in the second tank without increasing co-extraction of the unwanted impurities such as IgA, IgM, lipid and high molecular weight proteins .

EXAMPLE 5A

[0171] For the experiments Cohn I+II+III paste (1 kg containing 120 gram of Celpure C100) was suspended at a first dilution ratio of 1:6 in 10 mM sodium acetate and 10 mM Sodium dihydrogen phosphate dihydrate, pH 4.3-4.4 buffer at 4° C. in the first tank. The suspension in the first tank was stirred at 4° C. with a paddle stirrer for approximately 15-20 hours. Prior to starting the experiments the first filtration unit (a Novoflow dynamic filtration device containing three ceramic filters with 0.2 μm membranes; filter area=0.1 m.sup.2) was stored overnight in cold water (1° C.). At the start of the experiments the water was drained from the unit and the suspension was fed into the unit. The suspension was then recirculated for several minutes between the first tank and the first process unit prior to beginning the filtration process. The remaining suspension was gradually added to the first process unit during the filtration process. The ceramic filters in the first process unit were operated at a rotation speed of 1200 rpm and a TMP of 1.2 bar. The permeate from the first process unit was collected in a second tank and then fed into a second unit referred to as the UF/DF unit. The UF/DF unit was a Novoflow dynamic filtration device containing 6 ceramic disks, with 7.0 nm membranes; filter area 0.2 m.sup.2. The permeate flow rate of the UF/DF system was 50-70 mL/min. The UF/DF system was started once the first filtrate was collected in the second tank. The retentate of the UF/DF system flowed back into the second tank while the permeate flowed back into the first tank. The volume of permeate fed back into the first tank contributed to the overall volume liquid mixed with the paste (i.e. the final dilution factor). In this experiment, the overall recirculation volume was 107 L per kg of paste (i.e. 1:107 final dilution factor).

EXAMPLE 5B

[0172] In this Example the same procedure was used as described in Example 5A with the exception that the overall recirculation volume was 16 L/kg paste.

EXAMPLE 6A

[0173] In this Example the same procedure was used as described in Example 5A with the exception that the rotation speed of the ceramic filters in the first process unit were operated at 1000 rpm and the overall recirculation volume was 102 L/kg paste.

EXAMPLE 6B

[0174] In this Example the same procedure was used as described in Example 6A with the exception that the overall recirculation volume was 16 L/kg paste.

EXAMPLE 7A

[0175] In this Example the same procedure was used as described in Example 6A with the exception that the the rotation speed of the ceramic filters in the first process unit were operated at 800 rpm and the overall recirculation volume was 93 L/kg paste.

EXAMPLE 7B

[0176] In this Example the same procedure was used as described in Example 7A with the exception that the overall recirculation volume was 16 L/kg paste.

[0177] Table 8 provides the experimental parameters and the target protein yield in the second tank following the filtration process.

TABLE-US-00008 Example 5A 5B 6A 6B 7A 7B Rotation speed (rpm) 1200 1200 1000 1000 800 800 Amount of Paste (kg) 1 1 1 1 1 1 First dilution factor 1:6 1:6 1:6 1:6 1:6 1:6 (Paste:buffer) Final dilution factor = 107 16 102 16 93 16 Overall Recirculation volume (L/kg paste Membrane area (m.sup.2) 0.1 0.1 0.1 0.1 0.1 0.1 Filtration membrane Ceramic Ceramic Ceramic Ceramic Ceramic Ceramic Pore diameter (μm) 0.2 0.2 0.2 0.2 0.2 0.2 Transmembrane Pressure 1.2 1.2 1.2 1.2 1.2 1.2 TMP (bar) UF/DF Membrane area (m.sup.2) 0.2 0.2 0.2 0.2 0.2 0.2 Pore diameter (nm) 7 7 7 7 7 7 Membrane type MgAl.sub.2O.sub.4 MgAl.sub.2O.sub.4 MgAl.sub.2O.sub.4 MgAl.sub.2O.sub.4 MgAl.sub.2O.sub.4 MgAl.sub.2O.sub.4 Transmembrane Pressure 1.6 1.6 1.6 1.6 1.6 1.6 TMP (bar) Protein Yield (%)* 96.2 92.2 95.5 91.3 92.0 87.8 IgG Yield (%)* 98.9 95.6 97.4 95.2 95.1 94.4 MgAl.sub.2O.sub.4 = Magnesium Aluminium Oxide *The protein yield is calculated in relation to the starting suspension before the start of the continuous filtration process. The protein-and IgG yield in suspension was set as 100%, (100% total protein in suspension = 13.29 g protein/L plasma equivalent) and (100% total IgG in suspension = 7.34 g IgG/L plasma equivalent)

[0178] Impurities were determined by nephelometry, ELISA (IgA, IgM), or enzyme test methods (lipids). The results showed that no significant increases in impurities were observed, but as is clearly shown in Table 8, the protein yield as well as the IgG yield significantly increased with higher rotation speeds, and the yields of both protein and IgG were significantly higher when using a high final dilution factor (Examples 5A, 6A, 7A), as compared to a low final dilution factor (Examples 5B, 6B, 7B)

EXAMPLE 8A AND 8B

[0179] In this example the dynamic cross flow filtration process was compared in the presence (Example 8A) and absence (Example 8B) of filter aid. The filter aid used was (Celpure C300; Advanced Minerals).

[0180] For the experiments Cohn I+II+III paste (1.2 kg for Example 8A and 1.5 kg for Example 8B was used). Each kilogram of paste contains 120 grams of filter aid. The paste was resuspended at an initial ratio of 1:6 in 10 mM sodium acetate & 10 mM sodium dihydrogen phosphate dihydrate, pH 4.3-4.4 buffer at 4° C. The suspensions were stirred at 4° C. with a paddle stirrer for approximately 15-20 hours. Prior to starting the experiments the first filtration unit (a Novoflow dynamic filtration device containing three ceramic filter 0.2 μm membranes; filter area=0.1 m.sup.2) was stored overnight in cold water (1° C.). At the start of the experiments the water was drained and the suspension was fed into the dynamic filter device. The suspension was then recirculated for several minutes between the first tank and the device prior to beginning the filtration process. The remaining suspension was gradually added during the filtration processes (the ceramic filters were operated at 1200 rpm with a TMP of 1.2 to 1.6 bar. In each experiment, about 600 ml of buffer was exchanged 16-18 times. This corresponds to a buffer amount of about 9.6-10.8 L and serves to simulate the buffer recovery of the ultrafiltration/diafiltration unit during online operation. The filtrates were collected in an ice-cooled container under stirring (paddle stirrer) and after the filtrations were completed the filtrates were stirred for a further hour. Subsequently, the filtrates were concentrated to 20 g/L (±5 g/L) using an Äkta Crossflow device.

[0181] In Example 8B the filter aid was removed using a Mecaplex pressure filtration sleeve and polypropylene filter layers located between the first tank and the first filtration unit. All other conditions were the same as described in Example 8A.

[0182] The experiments demonstrated that the removal of the filter aid resulted in a higher filtration rate through the first filtration unit (FIG. 2). In addition the total protein and IgG recoveries in the filtrates were similar irrespective of the presence or absence of the filter aid. Moreover the removal of the filter aid prior to the first filtration unit also assisted with the removal lipids and other hydrophobic molecules present in the suspensions from the filtrate (data not shown). Thus the filtrate obtained from the first filtration unit when operated in the absence of the filter aid remained more stable (i.e. the filtrate turbidity was relatively lower and remained stable upon storage as compared to the filtrate obtained with the first filtration unit operated in the presence of filter aid). These results further suggest that removal of the filter aid prior to the first filtration unit will improve the throughput capability of the system.