Method for extracting proteins from precipitate and method for precipitating impurities

Abstract

The present invention relates to a method and system for extracting proteins, in particular recombinant and / or plasma-derived proteins, including immunoglobulins (Ig), such as immunoglobulin G (IgG), from precipitate.

Description

Technical Field

[0001] This invention relates to methods and systems for extracting proteins from precipitates, particularly recombinant and / or plasma-derived proteins, including immunoglobulins (Ig), such as immunoglobulin G (IgG).

[0002] Cross-reference to related applications

[0003] This application claims priority to Australian provisional patent application AU2019904386, the contents of which are incorporated herein by reference in their entirety. Background Technology

[0004] The demand for purified proteins, such as specific antibodies, has increased significantly. These purified proteins can be used for therapeutic and / or diagnostic purposes.

[0005] Human plasma has been used industrially for decades to produce widely established and accepted plasma-protein products, such as human albumin (HSA), immunoglobulins (IgG), coagulation factor concentrates (factor VIII, factor IX, prothrombin complex concentrate, etc.), and inhibitors (antithrombin, C1 inhibitors, etc.). In the development of drugs derived from this type of plasma, plasma fractionation methods have been established, resulting in intermediates enriched with certain protein components, which are then used as starting materials for plasma-protein products. Typical methods are reviewed, for example, in "Molecular Biology of Human Proteins" (Schultze HE, Heremans J.F.; Volume I: Nature and Metabolism of Extracellular Proteins 1966, Elsevier Publishing Company; pp. 236-317). These types of fractionation techniques allow the production of several therapeutic plasma-protein products from the same donor pool. This is economically advantageous compared to producing only one plasma-protein product from a single donor pool and has therefore been used as the industrial standard in plasma fractionation.

[0006] One example of this type of fractionation method is cold ethanol fractionation of plasma, pioneered by E. J. Cohn and his team during World War II, primarily for the purification of albumin (Cohn EJ, et al. 1946, J. Am. Chem. Soc. 62: 459–475). The Cohn fractionation method involves progressively increasing the ethanol concentration from 0% to 40% while simultaneously decreasing the pH from neutral (pH 7) to approximately 4.8, resulting in albumin precipitation. Although Cohn fractionation has been developed over the past 70 years, most commercial plasma fractionation methods are based on the original method or its variants (e.g., Kistler / Nitschmann), utilizing differences in pH, ionic strength, solvent polarity, and ethanol concentration to separate plasma into a series of major precipitated protein components (such as components I to V in Cohn).

[0007] Variations of the Cohn fractionation have been developed to improve the recovery of multivalent IgG. For example, Oncley et al. used Cohn fractions II+III as starting materials, with different combinations of cold ethanol, pH, temperature, and protein concentration as described by Cohn, to produce serum fractions of active immunoglobulins (Oncley et al., (1949) J. Am. Chem. Soc. 71, 541-550). Currently, the Oncley method is the classic method for producing multivalent IgG. However, it is known that approximately 5% of γ-globulin (the portion enriched with antibodies) is co-precipitated with fraction I through the fraction II+III step, and approximately 15% of the total γ-globulin present in plasma is lost (see Table III, Cohn EJ, et al. 1946, J. Am. Chem. Soc. 62: 459-475). The Kistler / Nitschmann method aims to improve IgG recovery by reducing the ethanol content in some precipitation steps (precipitate B vs fraction III). However, the increased yield comes at the cost of purity (Kistler & Nitschmann, (1962) Vox Sang. 7, 414-424).

[0008] Initially, immunoglobulin G (IgG) products derived from these fractionation methods were successfully used for the prevention and treatment of various infectious diseases. However, because ethanol fractionation is a relatively crude method, the IgG products contained impurities and aggregates, to the extent that they could only be administered intramuscularly. Since then, further improvements in purification methods have led to IgG products suitable for both intravenous (IVIg) and subcutaneous (SCIg) administration.

[0009] It is estimated that approximately 30 million liters of plasma were processed worldwide in 2010, providing a range of therapeutic products, including approximately 500 tons of albumin and 100 tons of IVIg. The IVIg market accounts for approximately 40-50% of the overall plasma fractionation market (P. Robert, Worldwide supply and demand of plasma and plasma-derived medicines (2011) J. Blood and Cancer, 3, 111-120). Therefore, with continued strong demand for IVIg (and increasing demand for SCIg), there remains a need to improve the recovery rate of immunoglobulins from plasma and related components. Preferably, this must be achieved in a manner that ensures it does not adversely affect the recovery rate of therapeutic proteins from other plasma sources.

[0010] From a commercial perspective, the initial fractionation process is crucial to the overall production time and cost associated with the production of therapeutic proteins, particularly plasma-derived proteins, as subsequent purification steps depend on the yield and purity of the protein of interest within these initial fractions. While several variants of cold ethanol fractionation methods have been developed for plasma-derived proteins to improve protein yields at lower operating costs, higher protein yields are often accompanied by lower purity.

[0011] An improved method and system is needed for the industrial-scale production of proteins, such as immunoglobulins derived from precipitates containing immunoglobulins (e.g., derived from plasma or serum), which must meet stringent safety standards. Currently used downstream technologies are relatively expensive, and their yields are not optimal. Therefore, there is an urgent need to develop more efficient and economical methods for the extraction and purification of proteins (such as immunoglobulins) from protein-containing suspensions.

[0012] Any prior art referenced in this specification is not an admission or representation that such prior art forms part of any general common knowledge under any law, nor is it an admission or representation that it can be reasonably expected that a person skilled in the art would understand such prior art, consider it relevant, and / or combine it with other prior art. Summary of the Invention

[0013] According to a first aspect of the present invention, a method for extracting a protein of interest from a precipitate is provided, the method comprising:

[0014] a. The precipitate is mixed with a liquid in a first container to form a first suspension having a first dilution factor;

[0015] b. The first suspension is fed into a first filtration unit, the first filtration unit comprising a dynamic filter element adapted to generate a first retention that consumes the protein of interest and a first permeate that enriches the protein of interest.

[0016] c. Dilute the first suspension in the first tank to a second dilution factor by adding liquid, optionally by introducing the first retainer into the first tank;

[0017] d. The first permeate enriched with the protein of interest was recovered in the second vessel;

[0018] e. Precipitate one or more impurities in a first permeate enriched with the protein of interest to produce a second suspension; and

[0019] f. Remove precipitated impurities from the second suspension to produce a solution containing the protein of interest.

[0020] In any embodiment of the foregoing aspects of the invention, a step is provided to concentrate and enrich the first permeate containing the protein of interest prior to the step of precipitating one or more impurities in the first permeate.

[0021] In any embodiment of the first aspect of the invention, removing precipitated impurities from the second suspension to produce a solution containing the protein of interest may include:

[0022] a) The second suspension is fed into a second filtration unit, the second filtration unit comprising a dynamic filter element adapted to produce a second retention containing one or more precipitated impurities and a second permeate enriched with the protein of interest.

[0023] b) Optionally, the second retained stream is fed into a tank containing the second suspension;

[0024] c) The second permeate enriched with the protein of interest is recovered in another container.

[0025] Preferably, removing precipitated impurities from the second suspension to produce a solution containing the protein of interest further includes:

[0026] a) The second permeate in another tank is continuously concentrated in a third filtration unit containing cross-flow filter elements, thereby producing a third retention enriched with the protein of interest and a third permeate that has consumed the protein of interest.

[0027] b) Optionally, the suspension in the tank containing the second suspension is diluted to a third dilution factor by introducing a third permeate into the tank containing the second suspension; and

[0028] c) Return the third retainer containing the protein of interest to the third container and / or collect the third retainer containing the protein of interest.

[0029] According to a second aspect of the present invention, a method for extracting a protein of interest from a precipitate is provided, the method comprising:

[0030] a) The precipitate is mixed with a liquid in a first container to form a first suspension having a first dilution factor;

[0031] b) The first suspension is fed into a first filtration unit, the first filtration unit comprising a dynamic filter element adapted to generate a first retention that consumes the protein of interest and a first permeate that enriches the protein of interest.

[0032] c) Dilute the first suspension in the first tank to a second dilution factor by adding liquid, optionally by introducing the first retainer into the first tank;

[0033] d) The first permeate enriched with the protein of interest was recovered in the second vessel;

[0034] d1) The first permeate in the second tank is subjected to a continuous concentration process in the second filtration unit, the second filtration unit comprising a cross-flow filter element adapted to produce a second retention enriched with the protein of interest and a second permeate that has consumed the protein of interest.

[0035] d2) Optionally, the suspension is diluted to a second dilution factor by diluting the first suspension in the first tank with a second permeate stream; and

[0036] d3) Return the second retainer containing the protein of interest to the second container and / or collect the second retainer containing the protein of interest.

[0037] e) Precipitate one or more impurities from a second retention enriched with the protein of interest to produce a second suspension; and

[0038] f) Remove precipitated impurities from the second suspension to produce a solution containing the protein of interest.

[0039] In any embodiment of the second aspect of the invention, removing precipitated impurities from the second suspension to produce a solution containing the protein of interest may include:

[0040] d) The second suspension is fed to a third filtration unit, the third filtration unit comprising a dynamic filter element adapted to produce a third retention containing one or more precipitated impurities and a third permeate enriched with the protein of interest.

[0041] e) Optionally, the third retained stream is introduced into a tank containing the second suspension to the third dilution factor;

[0042] f) Optionally, the third osmotic fluid enriched with the protein of interest can be recovered in another container.

[0043] Preferably, removing precipitated impurities from the second suspension to produce a solution containing the protein of interest further includes:

[0044] d) The third permeate in another tank is continuously concentrated in a fourth filtration unit, which includes a cross-flow filter element to produce a fourth retention enriched with the protein of interest and a fourth permeate that has consumed the protein of interest.

[0045] e) Optionally, the suspension in the tank containing the second suspension is diluted to a third dilution factor by introducing a fourth permeate stream into the tank containing the second suspension; and

[0046] f) Return the fourth retention containing the protein of interest to another container and / or collect the fourth retention containing the protein of interest.

[0047] In any embodiment of the first or second aspect of the invention, precipitation of one or more impurities in step e) (e.g., precipitation of impurities from a first permeate according to the first aspect of the invention or from a second retention according to the second aspect of the invention) can reduce the solubility of one or more impurities by altering the solubility potential of the solvent, and more specifically, by adding reagents and / or adjusting conditions (e.g., pH or conductivity).

[0048] In any embodiment of the first or second aspect of the invention, the reagent used to precipitate one or more impurities from the first permeate is an ionizable organic additive, such as a fatty acid. Preferably, the fatty acid comprises CH3(CH2). n The general structural formula of COOH. Preferably, the fatty acid is a C4-C10 carboxylic acid. The fatty acid can be saturated or unsaturated. More preferably, the fatty acid comprises enanthic (heptanoic) acid, caprylic (octanoic) acid, octenoic acid, pelargonic (nonanoic) acid, nonenoic acid, or capric (decanoic) acid. Most preferably, the fatty acid is caprylic acid. Salts or esters of any of the fatty acids described herein, such as octanoic acid esters, can also be considered as reagents.

[0049] In any embodiment, the amount of fatty acids, preferably caprylic acid, is about 0.1 g / g total protein, about 0.5 g / g total protein, about 0.75 g / g total protein, about 1 g / g total protein, about 1.5 g / g total protein, about 2.0 g / g total protein, about 2.5 g / g total protein, about 3.0 g / g total protein, about 3.5 g / g total protein, or about 4.0 g / g total protein. Preferably, the amount of fatty acids, preferably caprylic acid, is about 0.275 g / g total protein, about 0.280 g / g total protein, about 0.285 g / g total protein, about 0.290 g / g total protein, about 0.300 g / g total protein, or about 0.325 g / g total protein.

[0050] In any embodiment, the amount of fatty acids, preferably caprylic acid, is 0.1 g / g total protein, 0.5 g / g total protein, 0.75 g / g total protein, 1 g / g total protein, 1.5 g / g total protein, 2.0 g / g total protein, 2.5 g / g total protein, 3.0 g / g total protein, 3.5 g / g total protein, or 4.0 g / g total protein. Preferably, the amount of fatty acids, preferably caprylic acid, is 0.275 g / g total protein, 0.280 g / g total protein, 0.285 g / g total protein, 0.290 g / g total protein, 0.300 g / g total protein, or 0.325 g / g total protein.

[0051] In any embodiment of the first or second aspect of the invention, the reagent used to precipitate one or more impurities from the first permeate is a nonionic organic polymer. The nonionic organic polymer may be polyethylene glycol (PEG), polypropylene glycol, polyvinylpyrrolidone, dextran, cellulose, or other polymers.

[0052] In any embodiment of the first or second aspect of the present invention, the precipitation of one or more impurities in step e. can be carried out by sequential precipitation using different amounts of precipitating reagent.

[0053] Typically, the precipitate containing the protein of interest is an insoluble solid. It is usually in granular or paste form. Sometimes, the precipitate appears as a suspension. The solid fraction can then be collected, for example, by filtration and / or centrifugation. Alternatively, such a suspension can be added directly to the first tank to form a first dilution factor. Another option is to add the suspension to the first tank, and then add the liquid to the first tank to form a first dilution factor. Thus, in a particular embodiment, the precipitate containing the protein of interest is in suspension form when added to the first tank.

[0054] In any aspect of the invention, the precipitate containing the protein of interest is an intermediate product of an alcohol fractionation process, preferably plasma, more preferably human plasma. In a preferred embodiment, the precipitate is obtained from human plasma starting material. Even more preferably, the precipitate is obtained from 2500L-6000L of human plasma starting material.

[0055] In any aspect of the invention, the precipitate containing the protein of interest is a plasma component (intermediate product). In a particular embodiment, the component is a Cohn component. In a preferred embodiment, the plasma component is selected from the following: Cohn component I (Fr I), Cohn component II+III (Fr II+III), Cohn component I+II+III (Fr I+II+III), Cohn component II (Fr II), Cohn component III (Fr III), Cohn component IV (Fr IV), Cohn component V (Fr V), Kistler / Nitschmann precipitate A (KNA), Kistler / Nitschmann precipitate B (KNA), and Kistler / Nitschmann precipitate C (KNC). In a particularly preferred embodiment, the plasma component is selected from the following: Cohn component I (Fr I), Cohn component II+III (Fr II+III), Cohn component I+II+III (Fr I+II+III), or Kistler / Nitschmann precipitate A (KNA). The plasma component may be a combination of different components. For example, plasma components can be KNA and a combination of one or more of Fr I, FrII+III, and Fr I+II+III.

[0056] The method according to one aspect of the invention is suitable for extracting the protein of interest from other protein-containing solids. Examples include lyophilized and crystalline solid forms containing the protein of interest.

[0057] In any aspect of the invention, the protein-containing precipitate is obtained from culture supernatant or fermentation starting material. In some embodiments, the starting material is milk containing the protein of interest or a whey-containing composition. In other embodiments, the starting material is not milk.

[0058] In any aspect of this invention, the protein of interest is an immunoglobulin, preferably human immunoglobulin G (IgG), such as immunoglobulin G derived from human plasma or recombinant immunoglobulin G.

[0059] In any aspect of this invention, the protein of interest is albumin, preferably human albumin (HSA).

[0060] In any aspect of the invention, a starting composition having a first dilution factor is provided by mixing a protein-containing precipitate with a liquid such as a buffer or water to produce a first suspension. When the protein-containing precipitate is nearly solid (e.g., a very thick paste, granules, etc.), a liquid is added to the protein-containing precipitate to form a suspension as a starting composition.

[0061] The first suspension with a first dilution factor in step a) is a mixture in which solute-like particles, sometimes referred to herein as solids, are present in the solution. The particle size may vary and may include larger particles or smaller particles that do not settle (i.e., in colloidal form), wherein the larger particles would eventually settle if the solution were not mixed.

[0062] The first dilution factor can sometimes be referred to as the solids weight percentage (%w / v). This is defined as the weight of the dry solids in a given volume of suspension divided by the total weight of that volume of suspension, multiplied by 100. In a particular embodiment, the solids / suspension weight percentage in step a) is at least 5% (i.e., a first dilution factor of about 1:20), or at least 7.5% or at least 10%, or at least 12.5%, or at least 15%, or at least 17.5%, or at least 20% or at least 22.5%, or at least 25% or at least 27.5%, or at least 30%, or at least 35%, or at least 40%, or at least 50%. In some embodiments, the solids / suspension weight percentage in step a) is 10%-30%. In some embodiments, the solids / suspension weight percentage in step a) is 15%-25%. In a preferred embodiment, the solids / suspension weight percentage in step a) is 17.5%-22.5%. In one particular embodiment, the weight percentage of solids / suspension in step a) is 20%.

[0063] In any aspect of the invention, the first dilution factor is at least 3 (1:3; parts of precipitate:total), preferably 1-10, preferably 3-9, preferably 3-5, and preferably about 3, 5, 6, 7, 9, or 10. For example, when the protein-containing precipitate is agglomerates or paste, especially a very thick paste (with very high viscosity), a liquid is needed to suspend the paste or agglomerates.

[0064] For example, when the first dilution factor is 3 (1:3; 1 part protein-containing precipitate: total volume), this is equivalent to a dilution ratio of 1:2 (1 unit volume of solute (the substance to be diluted) to 2 unit volumes of diluent to give a total volume of 3 units).

[0065] In any aspect of the invention, the first dilution factor (containing protein precipitate: total amount) in the first container is at least 40, or at least 30, or at least 20, or at least 17.5, or at least 15, or at least 12.5, or at least 10, or at least 9, or at least 8, or at least 7, or at least 6, or at least 5.5, or at least 5, or at least 4.5, or at least 4, or at least 3.5, or at least 3, or at least 2.5, or at least 2, or at least 1.5, or at least 1.25. Preferably, the first dilution factor (containing protein precipitate: total amount) in the first container is at least 4.

[0066] In some implementations, the first dilution factor (containing protein precipitate:total) in the first container is 1:1-1:20, or 1:2-1:20, or 1:3-1:20, or 1:4-1:20, or 1:5-1:20, or 1:6-1:20, or 1:7-1:20, or 1:8-1:20, or 1:10-1:20, or 1:1-1:15, or 1:2-1:15, or 1:3-1:15, or 1:4-1:15, or 1:5-1:15, or 1:6-1:15, or 1:7-1:15, or 1:8-1:15, or 1:10-1:15, or 1:1-1 :10, or 1:2-1:10, or 1:3-1:10, or 1:4-1:10, or 1:5-1:10, or 1:6-1:10, or 1:7-1:10, or 1:8-1:10, or 1:9-1:10, or 1:3-1:7, or 1:3-1:8, or 1:3-1:9, or 1:4-1:7, or 1:4-1:8, or 1:4-1:9, or 1:5-1:7, or 1:5-1:8, or 1:5-1:9, or 1:3.5-1:5, or 1:4-1:5, or 1:1-1:3, preferably 1:9, 1:7, 1:5, or more preferably 1:3 or 1:1.

[0067] Suitably, the protein concentration in the protein-containing precipitate after resuspension is about 5-100 g / L, preferably 10-50 g / L, or more preferably 25-45 g / L. This includes 5, 10, 15, 20, 25, 30, 35, 40, 45, 50, 55, 60, 65, 70, 75, 80, 85, 90, 95, and 100 g / L, and any range between these amounts. In other embodiments, the protein concentration can be about 5-20 g / L, for example, about 8-12 g / L.

[0068] According to a preferred embodiment, the first suspension in the first tank or the second suspension (in the second or third tank) has a pH of about 3.0-9.0, preferably about 4.0-7.0, about 4.0-6.0, about 4.0-5.0, about 4.3-4.9, about 4.4-4.8, and more preferably about 5.0. Generally, the pH is measured in the solution before the protein precipitate is added to the solution; or the pH is measured directly after the protein precipitate is mixed with the solution. Typically, the pH of the solution is measured after mixing the precursor components. Alternatively, the pH can be determined by calculation based on the expected amount and concentration of the components in the mixture.

[0069] In some implementations, the conductivity of the first and second suspensions (and / or protein concentrates) can be adjusted. Typically, the conductivity of the protein suspension or concentrate is about 6-9 mS / cm when the pH is in the range of about 4.7-5.2, or adjusted to a similar conductivity before filtration.

[0070] In any aspect of the invention, the first suspension may be continuously fed into the first filtration unit. In a preferred embodiment, the first suspension is continuously fed into the first filtration unit until the first suspension has been diluted to at least a second dilution factor.

[0071] In a first aspect of the invention, the second suspension may be continuously fed into the second filtration unit. In a preferred embodiment, the second suspension is continuously fed into the second filtration unit until the second suspension has been diluted to at least a third dilution factor.

[0072] In a second aspect of the invention, the second suspension may be continuously fed into the third filtration unit. In a preferred embodiment, the second suspension is continuously fed into the third filtration unit until the second suspension has been diluted to at least a third dilution factor.

[0073] In another embodiment of the invention, a method for obtaining a permeate / filtrate enriched with the protein of interest (e.g., from first and second suspensions) is performed via a continuous separation process. This method is adapted to separate impurities from the first suspension to produce a filtrate enriched with the protein of interest and a retention containing one or more impurities. Furthermore, the method is adapted to separate precipitates from the second suspension to produce another filtrate enriched with the protein of interest (e.g., the second filtrate in the first aspect of the invention, the third filtrate in the second aspect of the invention). Preferably, the continuous separation process is a continuous filtration process, wherein one or more filter membranes or different types of filter membranes may be used. Continuous filtration processes, such as dynamic cross-flow filtration, can minimize the risk of filter element clogging.

[0074] Since one aspect of the method of the present invention may involve adding additional liquid to the suspension in the first tank, the second dilution factor is greater than the first dilution factor.

[0075] According to one embodiment of the invention, the second and third dilution factors (volume ratio of protein-containing precipitate to total recycled liquid volume) are 6-70, 10-70, about 10, about 20, about 30, about 40, preferably about 20-50. In other embodiments, the second and third dilution factors are about 60 or about 70 (1:70; parts of protein-containing precipitate:total). In specific embodiments, the second and third dilution factors are at least 20, or at least 30, or at least 40, or at least 50, or at least 60, or at least 70. The inventors of the invention have found that such high dilution factors can enhance extraction efficiency and thus increase yield.

[0076] In another embodiment, the predetermined protein concentration in the first suspension is below 0.1 g / L, preferably about 0.001-0.1 g / L; typically about 0.05-0.1 g / L. Providing such a value allows the separation process to be terminated immediately once this threshold is reached, thus avoiding inefficient extraction and filtration processes. For example, when the total protein concentration in the first suspension is below 0.1 g / L, the estimated total IgG concentration is below about 40-50 mg / L, making continued continuous extraction and filtration of the product of interest less economical.

[0077] In any aspect of the invention, the dynamic filter element in the filtration unit adapted to produce a permeate enriched with the protein of interest (i.e., from the first or second suspension) is a dynamic cross-flow filter element. It should be understood that the first and second filtration units comprise the filter element of the first aspect of the invention, and the first and third filtration units comprise the filter element of the second aspect of the invention.

[0078] In a preferred embodiment, the dynamic cross-flow filter element is a rotating cross-flow filter element. More preferably, the rotating cross-flow filter element includes a filter disc. The filter disc is typically mounted on a shaft component. In one embodiment, the rotating cross-flow filter element includes at least one filter disc and at least one shaft component.

[0079] According to a preferred embodiment of any aspect of the invention, the filter disc membrane is a ceramic membrane. More preferably, the ceramic membrane has a pore size in the range of greater than or equal to 5 nm to less than or equal to 2 μm. In a particular embodiment, the ceramic membrane has a pore size of about 0.2 μm to 2 μm. In a particular embodiment, the ceramic filter membrane has an average pore size in the range of greater than or equal to 5 nm to less than or equal to 200 nm (0.2 μm). In a particular embodiment, the ceramic filter membrane has an average pore size in the range of greater than or equal to 50 nm to less than or equal to 100 nm. Such filter discs are provided by Kerafol and Flowserve.

[0080] In a preferred embodiment, the filtration unit includes a pressure vessel. The suspension from the first tank can be continuously fed into the pressure vessel through an inlet. A distribution manifold can be used to achieve uniform distribution of the suspension within the vessel. Therefore, in a particular embodiment, the pressure vessel includes a distribution manifold.

[0081] In some embodiments, a filtration unit adapted to produce a permeate enriched with the protein of interest from a first and / or second suspension comprises a rotating cross-flow filter element. Preferably, the filter element comprises one or more filter discs evenly spaced along at least one hollow central collection axis. The filter discs may be arranged horizontally or vertically. When horizontal, they are spaced apart along a vertically oriented hollow collection axis. The collection axis and discs are rotatable. The suspension in the pressure vessel can then penetrate the outer membrane of the rotating filter discs into the hollow central portion of the discs, which is then conveyed to the central collection axis. Typically, the filtrate (i.e., the permeate enriched with the protein of interest) can then be removed from the axial portion of the filtration unit through a flange, while the remaining retention in the pressure vessel (i.e., the retention that has consumed the protein of interest) can be discharged from the vessel through an outlet. Typically, the retention is recycled to the first tank or the tank containing the second suspension to dilute the suspension. In this way, the retention from the first filtration unit can be used to dilute the suspension in the first tank to a second dilution factor. Furthermore, according to a first aspect of the invention, the residue from the second filtration unit can be used to dilute the suspension in the tank containing the second suspension. According to a second aspect of the invention, the residue from the third filtration unit can be used to dilute the suspension in the tank containing the second suspension.

[0082] Dynamic cross-flow filtration, such as rotary filtration, offers maximum filter efficiency. The cross-flow effect (tangential flow cleaning of the filter surface) is generated by the rotating filter discs, rather than by pumping a large volume across a fixed membrane as used in conventional (static) cross-flow filtration systems. Compared to conventional cross-flow techniques, the extreme cross-flow velocities generated on the rotating filter disc surface ensure highly efficient cleaning of the filter surface while consuming very little energy.

[0083] Temperature affects the viscosity of protein solutions, the solubility of proteins in suspensions, and also the flux when using membrane filtration.

[0084] The starting suspension used in the method of the present invention preferably has a temperature in the range of 0°C to the relevant protein denaturation temperature. The temperature is suitably in the range of about 10°C to about 50°C. In a particular embodiment, the temperature is in the range of about 18°C ​​to about 35°C, preferably in the range of about 18°C ​​to about 22°C.

[0085] The temperature of the suspension tank containing the reagent for precipitating one or more impurities in step e is preferably about 4°C to about 40°C. Preferably, the reagent is a fatty acid, more preferably octanoic acid.

[0086] Therefore, in any embodiment, fatty acids, preferably caprylic acid, are combined with a first permeate according to the first aspect of the invention or with a second retention product according to the second aspect of the invention at a temperature of about 4°C to about 40°C to produce a second suspension. In some embodiments, fatty acids, preferably caprylic acid, are combined with a first permeate according to the first aspect of the invention or with a second retention product according to the second aspect of the invention (e.g., in a second suspension tank) at a temperature of about 25°C to about 38°C, about 27°C to about 37°C, optionally about 27°C, about 32°C, or about 37°C. Optionally, the first permeate according to the first aspect of the invention or the second retention product according to the second aspect of the invention and fatty acids, when combined, are at a temperature of about 4°C to about 40°C, preferably about 25°C to about 38°C, about 27°C to about 37°C, and more preferably about 32°C.

[0087] The temperature in the filtration unit is controlled, preferably between 2°C and 25°C, more preferably between about 2°C and 15°C. Such a temperature ensures optimal extraction and separation processes while maintaining the bioreactivity of the protein of interest throughout the process.

[0088] It should be understood that the temperature in the suspension tank may be the same as or different from the temperature in the filtration unit. For example, while the permeate may circulate in the filtration unit at one temperature, the initial suspension temperature may be higher than the temperature at which filtration is performed. Similarly, the temperature of the permeate may be increased before and / or during incubation with fatty acids in a second suspension tank.

[0089] In one embodiment, the temperatures in the first and second suspension tanks and the filtration unit are the same. In another embodiment, the temperature in the first suspension tank and the filtration unit is lower than the temperature in the second suspension tank.

[0090] Filtration is performed at a transmembrane filtration pressure equal to or lower than the level the membrane can withstand, depending on the membrane material used herein, for example, with a pressure of about 0.2 to about 3 bar. The transmembrane pressure is typically 0.1 to 2.5 bar, preferably 0.2 to 2.4 bar, more preferably 0.4 to 2.0 bar, 0.5 to 1.8 bar, 0.6 to 1.6 bar, 0.6 to 1.5 bar, 0.7 to 1.5 bar, and most preferably 0.8 to 1.5 bar. According to another embodiment, a pressure of up to 2 bar is provided to the filtration unit, preferably 0.1 to 2.0 bar, or about 1.5 bar, 1.0 bar, or 0.5 bar.

[0091] According to another embodiment, the continuous extraction process in the filtration unit is further aided by adjusting the flow rate and / or residence time of the suspension or solution entering the filtration unit, and / or the flow rate of the retainer / raffinate containing impurities / precipitates, and / or the flow rate of the first permeate / filtrate enriched with the protein of interest. The filtration unit is adapted to separate impurities / precipitates from the first and second suspensions. For example, in one embodiment, the linear velocity of the suspension or solution entering the pressure vessel (filtration processing unit) can be about 0.27-1.66 m / s. In another example, the linear velocity of the retainer containing impurities / precipitates can be 0.25-1.33 m / s. In another example, the linear velocity of the permeate / filtrate enriched with the protein of interest can be 0.03-0.33 m / s. The linear velocity multiplied by the cross-sectional area yields the volumetric flow rate. Furthermore, turbulence can be generated in the first processing unit due to the speed of the rotating filter disc, where the velocity (sometimes referred to as the tangential velocity) can be about 1-7 m / s. According to one embodiment of the invention, the speed of the rotary disc filter is 1-10 m / s. In a preferred embodiment of the invention, the speed of the rotary disc filter is 5-7 m / s. More preferably, the speed of the rotary disc filter is 7 m / s at 60 Hertz (800 rpm). The rotational speed of the rotating cross-flow filter element is about 600 rpm (50 Hz) to about 1600 rpm (100 Hz), preferably about 800 rpm (60 Hz) to about 1200 rpm (80 Hz), preferably about 800 rpm (60 Hz), about 1000 rpm (70 Hz), or about 1200 rpm (80 Hz). As used herein, rotational speed in Hz refers to the speed of the electric motor. It can be correlated with speed in rpm using a suitable calibration curve.

[0092] This method allows for continuous extraction and separation processes to maximize the recovery of the protein of interest from the starting precipitate / feed (i.e., the first suspension) or from the second suspension. Due to the extraction process, virtually all of the protein of interest is extracted from the protein-containing precipitate and recovered in subsequent stages. This method also allows for the recirculation of liquids or diluents such as buffers or water in a closed system, thus maintaining the volume of liquid throughout the process while reducing footprint (i.e., large tank volume).

[0093] In a further embodiment, the invention includes a backwashing step in conjunction with dynamic cross-flow filtration. As used herein, the term "backwash" will be understood to mean that the process of liquid flowing into the filtration system is reversed in order to flush out contaminants that may have accumulated in the system. Preferably, backwashing is performed using the same buffer solution contained in the first or second suspension. It should be understood that the frequency, duration, and flow rate of backwashing can be adjusted to maximize filtration efficiency and the filtration period prior to the need for backwashing, as further described herein.

[0094] It is estimated that the methods and systems disclosed herein recover at least 75%, at least 80%, at least 85%, at least 90%, at least 95%, or generally at least 98% of the protein of interest from protein-containing precipitates. Therefore, in a specific embodiment of the first aspect of the invention, the method provides a recovery rate of at least 95%, or at least 96%, or at least 97%, or at least 98%, or at least 99% of the protein of interest from the precipitate. In a preferred embodiment, the recovery rate is at least 97% of the protein of interest recovered from the precipitate.

[0095] The first permeate / filtrate enriched with the protein of interest can be concentrated prior to the step of precipitating impurities in the first permeate. The concentration process can be ultrafiltration performed in a second filtration unit, such as as described in steps d1-d3 of the second aspect of the invention. Alternatively, the concentration process can use any standard concentration technique, including dialysis, filtration (including ultrafiltration), chromatography, or precipitation.

[0096] According to one embodiment of the first aspect of the invention, the method further includes a continuous concentration process in a third filtration unit of a second permeate (in a second vessel) enriched with the protein of interest, thereby producing a third retention enriched with the protein of interest and a third permeate having consumed the protein of interest.

[0097] According to one embodiment of a second aspect of the invention, the method further includes a continuous concentration process in a fourth filtration unit of a third permeate enriched with the protein of interest, thereby producing a fourth retention enriched with the protein of interest and a fourth permeate that has consumed the protein of interest.

[0098] According to embodiments of the first and second aspects of the present invention, a filtration unit for performing a continuous concentration process includes a dynamic cross-flow filter element, which includes a dynamic ultrafiltration device. Alternatively, the process includes a static ultrafiltration device.

[0099] In a preferred embodiment of the invention, the dynamic cross-flow filter element or ultrafiltration device for carrying out the concentration process comprises a membrane having a molecular weight cutoff value smaller than the molecular weight of the protein of interest. In these embodiments, the membrane cutoff value is selected to retain the protein of interest during the concentration process. As a general guideline, a nominal membrane cutoff value can be selected that is less than one-third of the molecular weight of the protein of interest to ensure that the protein is retained in the retentate.

[0100] In another embodiment, the dynamic cross-flow filter element or static ultrafiltration filter element used for the concentration process comprises a membrane having a molecular weight cutoff value greater than the molecular weight of the protein of interest. In such an embodiment, a nominal membrane cutoff value is selected to ensure that the protein of interest passes through the membrane and is collected in a second or fourth permeate instead of a second or fourth retention (in the case of the second aspect of the invention) or in a third permeate instead of a third retention (in the case of the first aspect of the invention).

[0101] In a dynamic implementation of the cross-flow filter element, preferably, the element is a rotating cross-flow filter element suitable for continuous concentration processes.

[0102] According to another preferred embodiment, the filter element for a continuous concentration process comprises a filter membrane having an average pore size of 5 nm-5000 nm, preferably 5 nm-2000 nm, 5 nm-1000 nm, 5 nm-500 nm, 5 nm-200 nm, 7 nm-1000 nm, more preferably 7 nm-500 nm, even more preferably 7 nm-100 nm, and most preferably 7 nm-80 nm. Of course, the average pore size can be in other combinations of the above ranges. Filter manufacturers typically assign terms such as nominal or average pore size ratings to commercial filters, which usually indicate compliance with certain retention criteria for particles or microorganisms rather than the actual pore geometry.

[0103] In one particular embodiment, the rotating cross-flow filter element for a continuous concentration process comprises a filter disc (such as a ceramic disc). In some embodiments, the filter disc comprises a membrane having the average pore size of a microfiltration filter. In other embodiments, the filter disc comprises a membrane having the average pore size of an ultrafiltration filter. In a further embodiment, the filter disc comprises a membrane having the average pore size of a permeate filter. In one embodiment, the average pore size of the filter disc membrane is in the range of greater than or equal to 5 nm to less than or equal to 2 μm. In a particular embodiment, the average pore size of the filter disc membrane is in the range of greater than or equal to 50 nm to less than or equal to 0.5 μm. In some embodiments, the filter disc membrane has an average pore size in the range of greater than or equal to 50 nm to less than or equal to 100 nm, or greater than or equal to 60 nm to less than or equal to 90 nm, or greater than or equal to 60 nm to less than or equal to 80 nm. In some embodiments, the filter disc membrane has an average pore size of 60 nm or 80 nm.

[0104] In a particularly preferred embodiment, the rotating cross-flow filter element for a continuous concentration process comprises a plurality of ceramic discs having pore sizes suitable for ultrafiltration and / or percolation. For example, the element preferably comprises at least one ceramic membrane having a pore size of 3 nm. Alternatively, the element preferably comprises at least one ceramic membrane having a pore size of 5 nm. Alternatively, the element preferably comprises at least one ceramic membrane having a pore size of 7 nm. Alternatively, the element preferably comprises at least one ceramic membrane having a pore size of 30 nm. The element may comprise a plurality of ceramic discs with different pore sizes, including those with pore sizes of 3 nm and 5 nm. The element may comprise a plurality of ceramic discs with different pore sizes, including those with pore sizes of 5 nm and 7 nm. The element may comprise a plurality of ceramic discs with different pore sizes, including those with pore sizes of 3 nm and 30 nm. The element may comprise a plurality of ceramic discs with different pore sizes, including those with pore sizes of 3 nm, 5 nm, 7 nm, and 30 nm.

[0105] According to another preferred embodiment, the filter element for the continuous concentration process comprises an ultrafiltration device, which includes a membrane in the form of a polymer membrane, such as polyethersulfone or regenerated cellulose. This membrane preferably has an average molecular weight cutoff of less than 50 kDa, more preferably less than 30 kDa, more preferably less than 10 kDa, or most preferably less than 5 kDa.

[0106] According to one embodiment of the first aspect of the invention, the method further includes diluting the suspension in the tank containing the second suspension by continuously flowing a second permeate and / or a second retention that has consumed the protein of interest into the tank containing the second suspension, thereby facilitating the dilution of the suspension to a third dilution factor.

[0107] Similarly, according to an embodiment of the second aspect of the invention, the method further includes diluting the suspension in the tank containing the second suspension by continuously flowing a third permeate and / or a third retention that has consumed the protein of interest into the tank containing the second suspension, thereby facilitating the dilution of the suspension to a third dilution factor.

[0108] According to one embodiment of one aspect of the invention, the method further includes diluting the suspension in the first tank to a second dilution factor by continuously flowing the retention from the first filtration unit and the second permeate from the second filtration unit into the first tank.

[0109] According to another preferred embodiment of the first aspect of the invention, a second container is provided to receive the first permeate, wherein the flow rate of the first permeate is controlled to maintain a substantially constant product volume in the second container. In a particular embodiment, fresh buffer solution is added to the first container in addition to the first permeate.

[0110] According to another preferred embodiment of the second aspect of the invention, a second container is provided to receive the first permeate and / or the second retention, wherein the flow rates of the first permeate and the second retention are controlled to maintain a substantially constant product volume in the second container. In a particular embodiment, fresh buffer solution is added to the first container in addition to the first permeate and / or the second retention.

[0111] According to another preferred embodiment of the second aspect of the invention, a third (fourth) tank is provided to receive the third permeate and / or the fourth retention, wherein the flow rates of the third permeate and the fourth retention are controlled to maintain a substantially constant product volume in the tank containing the second suspension.

[0112] According to one embodiment of one aspect of the invention, a first permeate / extract / filtrate is collected in a receiving tank (second tank), and once the suspension in the first tank has been completely filtered / extracted, the first permeate / extract from the receiving tank undergoes a continuous concentration process. This method is particularly suitable for smaller industrial-scale processes where dead volume in production equipment and piping can significantly affect the yield of the protein of interest. One example is a hyperimmune immunoglobulin product.

[0113] According to a preferred embodiment of a second aspect of the present invention, an industrial-scale method for extracting a protein of interest from a precipitate in high yield is provided, the method comprising:

[0114] a) The precipitate is mixed with liquid in a first container to form a suspension having a first dilution factor;

[0115] b) The first suspension is fed into a first filtration unit, the first filtration unit comprising a rotating cross-flow filter element comprising a filter disc having a ceramic membrane having an average pore size of 5 nm to 5000 nm, the filter element being adapted to produce a first retention that consumes the protein of interest and a first permeate enriched with the protein of interest.

[0116] c) Adding liquid in part by introducing the first retainer into the first tank to dilute the first suspension in the first tank to the second dilution factor;

[0117] d) The first permeate, enriched with the protein of interest, was recovered in the second vessel; and

[0118] e) The first permeate in the second tank is continuously concentrated in a second filtration unit containing a cross-flow filter element, thereby producing a second retention enriched with the protein of interest and a second permeate that has consumed the protein of interest.

[0119] f) Optionally, the suspension in the first tank is diluted to a second dilution factor by continuously flowing the second permeate into the first tank; and

[0120] g) Return the second retainer containing the enriched protein of interest to the second container and / or collect the second retainer containing the enriched protein of interest.

[0121] h) Precipitate one or more impurities in a second retention containing the protein of interest to produce a second suspension; and

[0122] i) Remove precipitated impurities from the second suspension to produce a solution containing the protein of interest.

[0123] The precipitated impurities can be removed from the second suspension according to the methods described herein, including by feeding the second suspension to a third filtration unit comprising a dynamic filter element adapted to generate a third retainer containing one or more precipitated impurities and a third permeate enriched with the protein of interest. Optionally, the third retainer is fed into a tank containing the second suspension; the third filtrate enriched with the protein of interest is recovered in another tank.

[0124] Preferably, removing precipitated impurities from the second suspension to produce a solution containing the protein of interest further includes: continuously concentrating the third permeate in another tank in a fourth filtration unit comprising a cross-flow filter element to produce a fourth retention enriched with the protein of interest and a fourth permeate consuming the protein of interest; optionally diluting the suspension in the tank containing the second suspension by flowing the fourth permeate into the tank containing the second suspension to dilute the suspension to a third dilution factor; and returning the fourth retention enriched with the protein of interest to the other tank and / or collecting the fourth retention enriched with the protein of interest.

[0125] According to a preferred embodiment of this aspect of the invention, the first retainer and the second permeate are continuously flowed into the first tank to dilute the suspension to a second dilution factor. Furthermore, the third retainer and the fourth permeate can be continuously flowed into the tank containing the second suspension to dilute the second suspension to a third dilution factor.

[0126] According to one embodiment of the present invention, a filtration unit adapted to produce a permeate enriched with the protein of interest comprises one or more hollow shafts adapted to collect the permeate, each shaft being connected to at least one filter disc containing a ceramic membrane.

[0127] According to one embodiment of the invention, the filtration unit for performing the concentration process comprises a dynamic cross-flow filter element. In other embodiments, the filtration unit comprises a static cross-flow filter element. In a preferred embodiment, the static cross-flow filter element is an ultrafiltration device comprising a membrane that retains the protein of interest in the retentate.

[0128] According to one embodiment of the invention, steps b) to c) or b) to f) are repeated until a predetermined value of the second dilution factor or protein concentration of the suspension or solution in the first tank is reached. This predetermined second dilution factor (sometimes referred to as the final dilution factor) ensures that an optimal yield can be obtained before continuing the extraction process becomes too uneconomical, and it can also be determined using a predetermined value of the protein concentration in the tank containing the first suspension. The protein concentration in the tank containing the first suspension can be monitored by various methods known in the art, including UV absorbance, such as at 280 nm.

[0129] Optionally, a filter aid may be used at an appropriate stage of the process. For example, the filter aid may be used in one or more steps involved in the preparation of the precipitate. Thus, in one embodiment, the precipitate contains a filter aid. In another embodiment, the precipitate does not contain a filter aid. In this embodiment, the filter aid may not be used at all in the process (including any of the foregoing steps), or, if present, the filter aid may be removed before feeding the suspension containing the precipitate to the first filtration unit, i.e., before step b). Preferably, the filter aid is removed before step b).

[0130] The product of the above method can then be further processed, including one or more of chromatographic steps, virus inactivation steps, concentration and formulation, so that the final product is suitable for administration to a subject, preferably a human subject.

[0131] According to one embodiment of the present invention, a filtration unit adapted to produce a permeate enriched with the protein of interest (i.e., the first and second filtration units in the first aspect of the invention, or the first and third filtration units in the second aspect of the invention) further includes a scraper device adapted to control the bed height of the filter aid and / or the precipitated material on the outer surface of the filter disc membrane. This device can also assist in controlling the filtration flux and / or preventing filter clogging. In some embodiments, the height of the scraper device relative to the distance to the filter disc membrane surface is adjustable. In a particular embodiment, the scraper device is positioned at a distance of at least 20 cm, or at least 15 cm, or at least 10 cm, or at least 9 cm, or at least 8 cm, or at least 7 cm, or at least 6 cm, or at least 5 cm, or at least 4 cm, or at least 3 cm, or at least 2.5 cm, or at least 2 cm, or at least 1.5 cm, or at least 1 cm, or at least 0.5 cm, or at least 0.25 cm from the filter disc membrane.

[0132] According to another embodiment, the filtration unit is equipped with a rotating filter disc (dynamic filter element) and optionally, baffles for turbulent mixing of the contents of the first filtration unit, preferably with a tangential velocity of about 1-7 m / sec. Turbulence can be generated by the baffles, thereby increasing the extraction of the protein of interest and achieving a high protein recovery rate.

[0133] In a preferred embodiment of the invention, the rotating cross-flow filter element includes one or more filter discs, the filter discs comprising ceramic membranes.

[0134] Ceramic filters can be composed of, for example, Al2O3 or ZrO3, TiO2 or MgAl2O4. Ceramic disc filters are typically designed so that the filtrate is transported from the outside through a ceramic membrane to a hollow internal channel where it can be collected.

[0135] Ceramic disc filters are available in various sizes, including 374mm (surface area 0.2m²). 2 ), 312mm (surface area 0.14m²) 2 ) and 152mm (surface area 360m²) 2 The outer diameter of the ceramic disc filter. Typically, the thickness of a ceramic disc filter ranges from approximately 4.5 to 6 mm.

[0136] In embodiments of the present invention, the first and third units have a filtration capacity of at least 25 kg, at least 50 kg, at least 75 kg, at least 100 kg, at least 200 kg, at least 300 kg, at least 350 kg, at least 400 kg, at least 450 kg, at least 500 kg, at least 550 kg, at least 600 kg, at least 650 kg, at least 700 kg, at least 750 kg, or at least 1000 kg of initial precipitate / m³. 2 Filter surface area.

[0137] In another aspect, the present invention provides a method for precipitating one or more impurities from a solution containing a protein of interest, the method comprising:

[0138] a) Provide a solution containing the protein of interest and one or more impurities, wherein the solution has a conductivity of 6-9 mS / cm, a pH of 4.7-5.2, and a total protein concentration of 15-17 g / L; and

[0139] b) Precipitate one or more impurities by adding fatty acids to a solution, wherein the amount of said fatty acids added is about 0.1 g to about 0.4 g fatty acids / g total protein.

[0140] In another aspect, the present invention provides a method for removing one or more impurities from a solution containing a protein of interest, the method comprising:

[0141] a) Provide a solution containing the protein of interest and one or more impurities, wherein the solution has a conductivity of 6-9 mS / cm, a pH of 4.7-5.2, and a total protein concentration of 15-17 g / L;

[0142] b) Precipitate one or more impurities by adding fatty acids to a solution, wherein the amount of said fatty acids added is about 0.1 g to about 0.4 g fatty acids / g total protein;

[0143] c) Remove one or more precipitated impurities.

[0144] In another aspect, the present invention provides a method for removing one or more impurities from a solution containing a protein of interest, the method comprising:

[0145] a) Provide a solution containing the protein of interest and one or more impurities, wherein the solution has a conductivity of 6-9 mS / cm, a pH of 4.7-5.2, and a total protein concentration of 15-17 g / L;

[0146] b) Optionally, the solution is filtered in a filtration unit containing a rotating cross-flow filter element to obtain a permeate enriched with the protein of interest.

[0147] c) Precipitate one or more impurities by adding fatty acids to a solution, preferably to a permeate enriched with the protein of interest, wherein the amount of said fatty acids added is about 0.1 g to about 0.4 g fatty acids / g total protein;

[0148] d) Remove one or more precipitated impurities.

[0149] Advantageously, the above method enables the recovery of the protein of interest to achieve a loss of less than about 1%, about 2%, about 3%, about 4%, about 5%, about 6%, about 7%, about 8%, about 9%, less than about 10%, or less than about 20% of the protein of interest during the removal of one or more impurities.

[0150] Preferably, the fatty acid contains CH3(CH2). n The general structural formula of COOH. More preferably, the fatty acid is a C7-C10 carboxylic acid. More preferably, the fatty acid includes heptanoic acid, octanoic acid, octenic acid, nonanoic acid, nonenoic acid, or decanoic acid. Most preferably, the fatty acid is octanoic acid. Salts or esters of any of the fatty acids described herein may also be considered as reagents.

[0151] In any embodiment of one aspect of the invention, the amount of fatty acids, preferably caprylic acid, is about 0.1 g / g total protein, about 0.5 g / g total protein, about 0.75 g / g total protein, about 1 g / g total protein, about 1.5 g / g total protein, about 2.0 g / g total protein, about 2.5 g / g total protein, about 3.0 g / g total protein, about 3.5 g / g total protein, or about 4.0 g / g total protein. Preferably, the amount of fatty acids, preferably caprylic acid, is about 0.275 g / g total protein, about 0.280 g / g total protein, about 0.285 g / g total protein, about 0.290 g / g total protein, about 0.300 g / g total protein, about 0.325 g / g total protein, or about 0.35 g / g total protein.

[0152] In any embodiment of one aspect of the invention, the amount of fatty acid, preferably caprylic acid, is 0.1 g / g total protein, 0.5 g / g total protein, 0.75 g / g total protein, 1.0 g / g total protein, 1.5 g / g total protein, 2.0 g / g total protein, 2.5 g / g total protein, 3.0 g / g total protein, 3.5 g / g total protein, or 4.0 g / g total protein. Preferably, the amount of fatty acid, preferably caprylic acid, is 0.275 g / g total protein, 0.280 g / g total protein, 0.285 g / g total protein, 0.290 g / g total protein, 0.300 g / g total protein, 0.325 g / g total protein, or about 0.35 g / g total protein.

[0153] In any embodiment of one aspect of the invention, the solution may be a component derived from a plasma fractionation process. Preferably, the solution is obtained or can be obtained from steps a) to d) (optionally including steps d1 to d3), as described in the method of the invention.

[0154] In any embodiment of one aspect of the invention, the solution does not contain a filter aid.

[0155] In any embodiment of one aspect of the invention, the protein of interest is an immunoglobulin, preferably human immunoglobulin G (IgG), such as immunoglobulin G derived from human plasma or recombinant immunoglobulin G.

[0156] In any embodiment of one aspect of the invention, the solution contains IgG.

[0157] In any embodiment of one aspect of the invention, the solution contains one or more of the following impurities: IgA, IgM, albumin, α-2 macroglobulin, α-1 antitrypsin, lipids, and lipoproteins.

[0158] As used herein, unless the context otherwise requires, the term “comprise” and its variations, such as “comprising,” “comprises,” and “comprised,” do not imply the exclusion of other additives, ingredients, integers, or steps.

[0159] Other aspects of the invention and other embodiments described in the foregoing paragraphs will become apparent from the following description, which is given by way of example and reference to the accompanying drawings. Attached Figure Description

[0160] The following figures are not necessarily drawn to scale, but rather focus on illustrating the principles of various embodiments. In the following description, various embodiments of the invention are described with reference to the following figures:

[0161] Figure 1 A schematic flowchart of the system of the present invention is provided below, and it is described in more detail below.

[0162] Figure 1A schematic flowchart illustrating a preferred embodiment of the system 100 and method according to the present invention is provided. A protein-containing precipitate, for example in the form of a suspension, paste, or precipitate, is suspended in a liquid such as a buffer solution. The composition and concentration of the buffer solution are as described above to produce a starting composition, such as a suspension having a first dilution factor of, for example, 3-10 (1:3-1:10). The suspension is placed in a first tank 1. The suspension can be fed to a first filtration unit 5 via a pump 2, using several types of pumps (e.g., piston pumps; rotary pumps; centrifugal pumps and membrane pumps) and a flow control valve 3 on pipe 12. The first filtration unit 5 is equipped with a rotating hollow shaft on which filter discs are mounted (filtrate flows from the outside to the inside of the hollow shaft). The first filtration unit 5 is further provided with a height-adjustable scraper to maintain a constant filter cake thickness, thus achieving a constant filtrate flow. The desired filtration pressure is controlled and regulated by an overflow valve (unfiltered suspension outlet). The filter discs used can be ceramic membranes, depth filter layers, and sintered porous metal filter discs. Once the container of the first filtration unit 5 is filled with the suspension, continuous pressure extraction and separation can begin. The first filtration unit 5, which may include a pressure unit / container, has suitable internal settings and conditions to simultaneously improve extraction efficiency and the filtration process. Extraction efficiency is improved through turbulent mixing in unit 5, without necessarily involving a mixer. However, it is foreseeable that an additional mixer could be provided to assist the extraction process by generating turbulence. Furthermore, higher final dilution factors, such as 40 or 70 disclosed in this invention, also improve extraction efficiency, resulting in high protein (e.g., IgG) yields. Of course, any other higher final dilution factor (above 70) is also conceivable.

[0163] The filtrate flows through a flow meter 6 installed on pipe (or channel) 14 and is collected in a second tank 7. The unfiltered suspension is returned through a regulating outlet 3 installed on pipe 13 in tank 1. When the 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 into the ultrafiltration (UF) system 8 through pipe 15. This sets the transmembrane pressure so that the permeate flow rate 17 is the same as or nearly the same as the first filtrate flow rate in pipe 14. The permeate from the UF system 8 flows back to the first tank 1 through pipe (or line or channel) 17, while the retained product of the UF system (= concentrated protein) flows back to the second tank 7 through pipe 16.

[0164] According to the invention, the first processing unit 5 is equipped with one or more rotating filter discs, each containing one or more first filter elements for turbulent mixing of the contents of the first processing unit 5 to produce a first retention and a first permeate. The first retention can be fed back to the first tank 1 via a control valve 3 through a channel 13, while the first permeate can be fed to the second tank 7 via another channel 14. The first filter element can be a ceramic-based filter membrane having a pore size of about 5 nm to 5000 nm, preferably 20 nm to 100 nm, or more preferably 30 nm to 80 nm. It is also foreseeable that inorganic membranes or any other suitable membranes can provide similar effects to ceramic-based membranes. The first filtration unit 5 can be equipped with a pressure control device 4, such as a pressure gauge, to regulate the pressure therein. Similarly, a flow meter 6 can be installed in the system of the invention for measuring the flow rate of the suspension or solution.

[0165] The feed stream from the second tank 7 can then be fed through channel 15 to the second filtration unit 8 for a second separation process. This second separation process can be a continuous concentration process (e.g., UF). The second filtration unit 8 is equipped with one or more second cross-flow filter elements, wherein the second cross-flow filter elements may comprise an ultrafiltration membrane with an average molecular weight cutoff of less than 50 kDa. However, the membrane may also be less than 10 kDa, or more preferably less than 5 kDa. Thus, the ultrafiltration membrane produces a second retention, which is returned to the second tank 7 through channel 16, while the second permeate is fed to the first tank 1 through channel 17. For this purpose, it should be noted that the pressure of the second filtration unit 10 can be adjusted during the concentration step (ultrafiltration) so that the flow rates in channels 14 and 17 are substantially equal.

[0166] Figure 2 This is a schematic overview of another system of the present invention, which is described in more detail below. Detailed Implementation

[0167] Certain embodiments of the invention will now be described in detail. While the invention will be described in conjunction with the embodiments, it is to be understood that it is not intended to limit the invention to these embodiments. Rather, the invention is intended to cover all alternatives, modifications, and equivalents that are included within the scope of the invention as defined by the claims.

[0168] Those skilled in the art will recognize that many methods and materials similar to or equivalent to those described herein can be used in the practice of this invention. This invention is by no means limited to the methods and materials described. It should be understood that the invention disclosed and defined in this specification extends to all alternative combinations of two or more individual features mentioned or apparent from the text or drawings. All these different combinations constitute various alternative aspects of the invention.

[0169] For the purposes of this specification, terms used in the singular will also include the plural, and vice versa.

[0170] This invention relates to systems and methods for maximizing protein recovery and yield by utilizing a novel first processing unit or further using a second processing unit and a precipitation step for removing one or more impurities. In this invention, a combination of extraction and separation methods is used in conjunction with a precipitation step to process precipitates containing solid proteins, i.e., intermediate materials (e.g., pastes derived from starting materials), wherein the starting precipitate may be suspended in a liquid or diluent, such as water or a buffer, to form a suspension.

[0171] One advantage of the method of the present invention is the maximization of the total recovery of the protein of interest, thereby minimizing the loss of that protein during the steps required to remove impurities and other proteins of interest. These advantages stem in part from the application of a continuous extraction filtration method combined with a subsequent precipitation step. Continuous extraction methods are advantageous for downstream processing, and the conditions used can minimize the loss of the protein of interest, for example, by reducing the total amount of reagents required to precipitate impurities or proteins of interest.

[0172] Typical protein-containing precipitates form during protein purification after exposure to a precipitating agent such as ethanol. The solid is usually referred to as a precipitate or paste. The precipitate can be mixed with a liquid to form a suspension in which the solid particles are distributed throughout the liquid. Under specific extraction conditions, the proteins contained in these particles can gradually dissolve into the liquid phase.

[0173] Solubility issues arise in industrial-scale production processes due to the need for large amounts of water or buffer solutions. For plasma fractionation processes used to produce proteins such as albumin and immunoglobulins, this step can involve thousands of liters. Even when tanks are available to accommodate such large volumes for high solubility, the desired higher yields cannot be achieved due to the balance between proteins dissolved in solution and those retained in the precipitate or paste (Le Chatelier's principle). These proteins confined to the precipitate may not be recoverable for further processing into the final product. This phenomenon involves solubility equilibrium in part. Known solubility equilibria exist when compounds in the solid phase are in chemical equilibrium with those dissolved in solution. This equilibrium is an example of a dynamic equilibrium, where some individual molecules migrate between the solid and liquid phases, resulting in equal rates of dissolution and precipitation.

[0174] This invention aims to solve the problem of protein recovery from precipitates by (a) continuously altering the dissolution equilibrium and (b) precipitating one or more impurities. Continuous alteration of the dissolution equilibrium is achieved by: 1) increasing extraction efficiency through built-in components (utilizing a dynamic filtration system that can incorporate hollow rotating disc filter elements) to allow for repeated close phase contact; 2) continuously removing dissolved protein from protein-containing precipitates (by applying the Le Chatelier principle—when the concentration (e.g., volume), temperature, or pressure of any system in equilibrium changes, the system readjusts to offset (partially) the effects of the applied change and establish a new equilibrium. This means that by continuously increasing the volume of the resuspension point and continuously removing dissolved protein through a dynamic filter, Le… The Chatelier principle can be used to ensure maximum transfer of protein from the precipitate to the liquid phase. In some cases, volume increases can be achieved by circulating permeate during successive protein concentration steps, thereby reducing buffer consumption. Furthermore, one or more impurities can be precipitated by altering the solvation potential of the solvent, and more specifically, by adding reagents and / or adjusting conditions (e.g., pH or conductivity) to reduce the solubility of one or more impurities. The result is a higher recovery rate of the protein of interest with the same, substantially the same, or reduced impurity properties and / or amounts. In other words, in the context of this invention, the step of precipitating one or more impurities results in a higher recovery rate of the protein of interest with the same, substantially the same, or reduced impurity properties and / or amounts than if the precipitation step were not performed.

[0175] Furthermore, the present invention relates to determining the amount of reagent that significantly reduces the solubility of one or more impurities, said reagent being used to precipitate one or more impurities from a solution containing the protein of interest. In particular, a smaller amount of reagent can be used when the solution containing the protein of interest has certain properties. The advantage is that process costs can be reduced by using a smaller amount of reagent, and the recovery rate of the protein of interest is higher due to the reduced amount of protein of interest co-precipitated with one or more impurities.

[0176] definition

[0177] The term "precipitate containing a protein" refers to any material containing the protein of interest. In the context of immunoglobulins as the protein of interest, this term can refer to plasma, serum, precipitates derived from plasma or serum, fermentation broth, inclusion bodies, cell culture supernatants, or precipitates derived from such materials. Typically, in the context of this invention, it refers to precipitates derived from plasma, such as Cohn or Oncley ethanol precipitates, or Kistler-Nitschmann precipitates.

[0178] The term "starting composition" refers to a suspension or solution that arises from a protein-containing precipitate, typically produced by dilution with water or a buffer according to a (first) dilution factor. In some cases, if dilution of the protein-containing precipitate is not required, the protein-containing precipitate may be a starting suspension.

[0179] "High yield" means that the yield of the protein of interest, such as immunoglobulin G (and other proteins and immunoglobulins), is at least 95% of the amount of the protein of interest in the precipitate containing the protein, preferably at least 96%, more preferably at least 98%, and most preferably more than 98%.

[0180] The concentration of immunoglobulins in a sample (e.g., in a precipitate or its pharmaceutically purified product) can be measured by any method known to those skilled in the art. It should be understood that the method used to measure immunoglobulins can depend on the nature of the sample. For example, it should be understood that when the sample is a precipitate containing immunoglobulins, it may be necessary to dissolve the precipitate (or its sample) in a suitable buffer prior to measurement. Examples of suitable assays for measuring the protein of interest include high-performance liquid chromatography (HPLC; e.g., size exclusion HPLC), enzyme-linked immunosorbent assay (ELISA), and quantitative immunoturbidimetry.

[0181] "About" or "approximately" in relation to a given value of percentage, pH, amount, time, or other reference indicates a value that includes up to 10% of the specified value.

[0182] Large-scale or industrial-scale processes or systems

[0183] The large-scale or industrial-scale nature of this invention refers to a production process based on a starting material such as human plasma of at least 200 L, preferably at least 500 L, and even more preferably at least 2000 L. For example, typical commercial plasma donor pool sizes used in industrial-scale protein production range from 2500 L to 6000 L of plasma per batch. In certain embodiments of the invention, the precipitate is obtained from 2500 L to 6000 L of plasma. Some commercial manufacturing processes are capable of using larger plasma donor pool sizes, including up to 7500 L, up to 10000 L, and / or up to 15000 L of plasma.

[0184] The method and system of the present invention can be used not only for large-scale industrial applications, but also as a stand-alone system and / or a method for smaller-scale production applications (where the starting material may be less than 200L).

[0185] Precipitates and proteins of interest

[0186] Many different methods can be used to selectively precipitate proteins from solution, such as by adding salt, alcohol, and / or polyethylene glycol in conjunction with pH adjustment and / or cooling steps. Therefore, the present invention is expected to be applicable to most protein precipitates, such as those containing immunoglobulin G, regardless of how they were originally prepared. It should be noted that the present invention can also be used to separate other types of proteins, including albumins, immunoglobulins (Ig), such as IgA, IgD, IgE, or IgM, each type individually or mixtures thereof. Recombinant proteins are also expected to be applicable in this regard.

[0187] Therefore, it should be noted that if the method is used to produce IgG, the protein-containing precipitate can be any substance containing IgG (e.g., in the form of a paste, precipitate, or inclusion body) or derived from starting materials, such as solutions of IgG that can be precipitated by, for example, one or more of the methods described above, whether from human or animal plasma or serum, fermentation broth, cell culture, protein suspension, milk, or other original sources. Materials or solutions containing immunoglobulins can contain monoclonal or polyclonal immunoglobulins. In some embodiments, the starting material containing immunoglobulins is a solution containing polyclonal antibodies. In other embodiments, the starting material contains monoclonal antibodies or fragments thereof. Therefore, to the knowledge of those skilled in the art, the term "immunoglobulin" as used herein can also be identified as an antibody, including natural or recombinant monoclonal or polyclonal antibodies.

[0188] For example, immunoglobulins (e.g., IgG) can be isolated from human or animal blood, or produced by other means such as recombinant DNA technology or hybridoma technology. In a preferred embodiment, immunoglobulins are obtained from plasma, typically from a pool of plasma derived from multiple donors. To obtain immunoglobulins from plasma, the plasma is typically subjected to alcohol fractionation, which can be combined with other purification techniques such as chromatography, adsorption, or precipitation. However, other methods may also be used. For example, the protein-containing precipitate can be a II+III precipitate according to the Cohn method, such as method 6, Cohn et al. J. Am; Chem. Soc., 68(3), 459-475 (1946), method 9, Oncley et al. J. Am; Chem. Soc., 71, 541-550 (1946), or an I+II+III precipitate, method 10, Cohn et al. J. Am; Chem. Soc., 72, 465-474 (1950); and the method of Deutschet. al. J. Biol. Chem. 164, 109-118 (1946), or precipitate A of Nitschmann and Kistler Vox Sang. 7, 414-424 (1962); Helv. Chim. Acta 37, 866-873 (1954). Optional precipitates containing the protein of interest include, but are not limited to, other Oncley fractions containing immunoglobulin G, Cohn fractions, and ammonium sulfate precipitates from plasma as described by Schulze et al. in U.S. Patent 3,301,842. Other optional precipitates containing the protein of interest include, but are not limited to, caprylic acid precipitates, such as those described in EP893450.

[0189] “Normal plasma,” “hyperimmune plasma” (such as hyperimmune anti-D, tetanus or hepatitis B plasma), or any plasma equivalent thereof can be used as the starting material in the cold ethanol fractionation process described herein.

[0190] The term "cryo-poor plasma" (also known as cryo-poor plasma, plasma depleted of cryoprecipitate, etc.) refers to plasma (derived from whole blood donation or plasma exchange) from which cryoprecipitate has been removed. Cryoprecipitation is the first step in most plasma protein fractionation methods used today for the large-scale production of plasma protein therapeutics. The method generally involves mixing frozen plasma thawed under controlled conditions (e.g., at or below 6°C) and then collecting the precipitate by filtration or centrifugation. The supernatant component is typically retained for use and is referred to by those skilled in the art as "cryo-poor plasma." The resulting cryo-poor plasma has reduced levels of factor VIII (FVIII), von Willebrand factor (VWF), factor XIII (FXIII), fibronectin, and fibrinogen. While FVIII levels are significantly reduced, fibrinogen levels can be as high as 70% of their original levels. Low-temperature supernatant provides commonly used raw materials for the manufacture of many therapeutic proteins, including α1-antitrypsin (AAT), apolipoprotein AI (APO), fibrinogen, antithrombin III (ATIII), prothrombin complex containing coagulation factors (II, VII, IX, and X), albumin (ALB), and immunoglobulins such as immunoglobulin G (IgG).

[0191] The supernatant of 8% ethanol-precipitate (Cohn et al. method; Schultze et al. (see above), p. 251), precipitate II+III (Oncley et al. method; Schultze et al. (see above), p. 253) or precipitate B (Kistler and Nitschmann method; Schultze et al. (see above Schultze), p. 253) is an example of IgG source compatible with industrial-scale plasma fractionation. Alternatively, the starting material used for the purification process to obtain IgG in high yield can be any other suitable material from different sources, such as fermentation and cell culture or other protein suspensions.

[0192] In the Cohn fractionation method, the first fractionation step results in fraction I, which mainly contains fibrinogen and fibronectin. The supernatant from this step is further processed to precipitate fractions II+III, then fractions III and II. Typically, fractions II+III contain approximately 60% IgG, along with impurities such as fibrinogen, IgM, and IgA. Most of these impurities are then removed from fraction III, which is considered a waste fraction and is usually discarded. The supernatant is then processed to precipitate the predominantly IgG-containing fraction, fraction II, which may contain more than 90% IgG. The percentage values ​​above refer to the percentage purity of IgG. Purity can be measured by any method known in the art, such as gel electrophoresis or immunoturbidimetry. In the Kistler & Nitschmann method, fraction I is equivalent to fraction I in the Cohn method. The next precipitate / fraction is called precipitate A (fraction A). Although different, this precipitate is roughly equivalent to Cohn fractions II+III. The precipitate is then redissolved and conditions are adjusted to precipitate precipitate B (fraction B), which is equivalent to Cohn fraction III. Similarly, this is considered a waste component and is typically discarded. The supernatant of precipitate B is then further processed to produce precipitate II, which corresponds to Cohn component II.

[0193] Specific protein-containing precipitates may include plasma proteins, peptide hormones, growth factors, cytokines, and polyclonal immunoglobulins. Plasma proteins are selected from human and animal blood coagulation factors, including fibrinogen, prothrombin, thrombin, prothrombin complex, FX, FXa, FIX, FIXa, FVII, FVIIa, FXI, FXIa, FXII, FXIIa, FXIII and FXIIIa, and von... Willebrand factors, transport proteins including albumin, transferrin, ceruloplasmin, haptoglobin, hemoglobin, and heme-binding proteins; protease inhibitors including β-antithrombin, α-antithrombin, α-2-macroglobulin, C1 inhibitor, tissue factor pathway inhibitor (TFPI), heparin cofactor II, protein C inhibitor (PAI-3), protein C and protein S, α-1 esterase inhibitor protein, α-1 antitrypsin; antiangionetic proteins including potential antithrombin; highly glycosylated proteins including α-1-acid glycoprotein, antichymotrypsin, α-trypsin interproteases, α-2-HS glycoprotein, and C-reactive protein; and other proteins including histidine-rich glycoproteins, mannan-binding lectins, C4-binding proteins, fibronectin, GC-globulins, plasminogen, blood factors such as erythropoietin, interferon, tumor factors, tPA, and γCSF.

[0194] In a particular embodiment, the protein-containing precipitate is used to prepare therapeutic proteins derived from plasma, including immunoglobulins such as immunoglobulin G, albumin, fibrin, thrombin, prothrombin complex, fibrinogen, plasminogen, α1-antitrypsin, C1-inhibitor, apolipoprotein A1, alpha-acid glycoprotein, haptoglobin, heme-binding protein, transferrin, and coagulation factors such as factor VII, factor VIII, and factor IX.

[0195] Dynamic filter element

[0196] In any aspect of the invention, the dynamic filter element filtration unit suitable for removing impurities from a first suspension and / or for removing precipitates from a second suspension is a dynamic cross-flow filter element. In a preferred embodiment, the dynamic cross-flow filter element is a rotary cross-flow filter element. More preferably, the rotary cross-flow filter element comprises a filter disc. The filter disc is typically mounted on a shaft member. In one embodiment, the rotary cross-flow filter element comprises at least one filter disc and at least one shaft member.

[0197] According to a preferred embodiment of any aspect of the invention, the filter disc membrane is a ceramic membrane. More preferably, the ceramic membrane has a pore size in the range of greater than or equal to 5 nm to less than or equal to 2 μm. In a particular embodiment, the ceramic membrane has a pore size of about 0.2 μm to 2 μm. In a particular embodiment, the ceramic filter membrane has an average pore size in the range of greater than or equal to 5 nm to less than or equal to 200 nm (0.2 μm). In a particular embodiment, the ceramic filter membrane has an average pore size in the range of greater than or equal to 50 nm to less than or equal to 100 nm. Such filter discs are provided by Kerafol and Flowserve.

[0198] It should be understood that a dynamic filter element filtration unit may include multiple filter disc membranes, the filtration unit being adapted to remove impurities from a first suspension and / or precipitates from a second suspension. Therefore, this method contemplates using one, two, three, four, five, six, or more filter disc membranes to remove impurities from the first suspension and / or precipitates from the second suspension. Multiple filter disc membranes may have the same or different pore sizes.

[0199] In a preferred embodiment, the filtration unit, suitable for removing impurities from a first suspension and / or for removing precipitates from a second suspension, comprises a pressure vessel. The suspension from the first tank can be continuously fed into the pressure vessel through an inlet. A distribution manifold can be used to achieve uniform distribution of the suspension within the vessel. Therefore, in a particular embodiment, the pressure vessel includes a distribution manifold. In some embodiments, the first filtration unit comprises a rotating cross-flow filter element. Preferably, the filter element contains one or more filter discs evenly spaced along at least one hollow central collection axis. The filter discs can be arranged horizontally or vertically. When horizontal, they are spaced apart along a vertically oriented hollow collection axis. The collection axis and the discs are rotatable. The suspension in the pressure vessel can then penetrate the outer membrane of the rotating filter discs, entering the hollow central portion of the discs, which is then conveyed to the central collection axis. Typically, the filtrate (i.e., the first permeate enriched with the protein of interest) can then be removed from the axial portion of the first filtration unit through a flange. Meanwhile, the remaining residue in the pressure vessel can be discharged from the vessel through an outlet. The residue is typically recycled back to the first tank to dilute the suspension. In this way, the residue from the first filtration unit can be used to dilute the suspension in the first tank to a second dilution factor.

[0200] Dynamic cross-flow filtration, such as rotary filtration, provides maximum filter efficiency. The cross-flow effect (tangential flow cleaning of the filter surface) is generated by the rotating filter discs, rather than by pumping a large volume across a fixed membrane as used in conventional (static) cross-flow filtration systems. Compared to conventional cross-flow techniques, the extreme cross-flow velocities generated on the rotating filter disc surface ensure highly efficient cleaning of the filter surface while consuming very little energy.

[0201] Dynamic filter elements can also be used in the third filtration unit of the first aspect of the invention, and / or the second or fourth filtration unit of the second aspect of the invention, in other words, to perform a continuous concentration process. Such dynamic filter elements typically comprise one or more ultrafiltration or permeation membranes.

[0202] Cross-flow filter elements for continuous concentration processes may include dynamic ultrafiltration devices. Alternatively, the method includes static ultrafiltration devices.

[0203] In a preferred embodiment of the invention, the dynamic cross-flow filter element or ultrafiltration device for carrying out the concentration process comprises a membrane having a molecular weight cutoff value smaller than the molecular weight of the protein of interest. In these embodiments, the membrane cutoff value is selected to retain the protein of interest during the concentration process. As a general guideline, a nominal membrane cutoff value can be selected that is less than one-third of the molecular weight of the protein of interest to ensure that the protein is retained in the retentate.

[0204] In another embodiment, the dynamic cross-flow filter element or static ultrafiltration filter element used for the concentration process comprises a membrane having a molecular weight cutoff value greater than the molecular weight of the protein of interest. In such an embodiment, a nominal membrane cutoff value is selected to ensure that the protein of interest passes through the membrane and is collected in a second or fourth permeate instead of a second or fourth retention (in the case of the second aspect of the invention) or in a third permeate instead of a third retention (in the case of the first aspect of the invention).

[0205] In a dynamic implementation of the cross-flow filter element, preferably, the element is a rotating cross-flow filter element suitable for continuous concentration processes.

[0206] According to another preferred embodiment, the filter element for a continuous concentration process comprises a filter membrane having an average pore size of 5 nm-5000 nm, preferably 5 nm-2000 nm, 5 nm-1000 nm, 5 nm-500 nm, 5 nm-200 nm, 7 nm-1000 nm, more preferably 7 nm-500 nm, even more preferably 7 nm-100 nm, and most preferably 7 nm-80 nm. Of course, the average pore size can be in other combinations of the above ranges. Filter manufacturers typically assign terms such as nominal or average pore size ratings to commercial filters, which usually indicate compliance with certain retention criteria for particles or microorganisms rather than the actual pore geometry.

[0207] In one particular embodiment, the rotating cross-flow filter element for a continuous concentration process comprises a filter disc (such as a ceramic disc). In some embodiments, the filter disc comprises a membrane having the average pore size of a microfiltration filter. In other embodiments, the filter disc comprises a membrane having the average pore size of an ultrafiltration filter. In a further embodiment, the filter disc comprises a membrane having the average pore size of a permeate filter. In one embodiment, the average pore size of the filter disc membrane is in the range of greater than or equal to 5 nm to less than or equal to 2 μm. In a particular embodiment, the average pore size of the filter disc membrane is in the range of greater than or equal to 50 nm to less than or equal to 500 nm (i.e., 0.5 μm). In some embodiments, the filter disc membrane has an average pore size in the range of greater than or equal to 50 nm to less than or equal to 100 nm, or greater than or equal to 60 nm to less than or equal to 90 nm, or greater than or equal to 60 nm to less than or equal to 80 nm. In some embodiments, the filter disc membrane has an average pore size of 60 nm or 80 nm.

[0208] In a particularly preferred embodiment, the rotating cross-flow filter element for a continuous concentration process comprises a plurality of ceramic discs having pore sizes suitable for ultrafiltration and / or percolation. For example, the element preferably comprises at least one ceramic membrane having a pore size of 3 nm. Alternatively, the element preferably comprises at least one ceramic membrane having a pore size of 5 nm. Alternatively, the element preferably comprises at least one ceramic membrane having a pore size of 7 nm. Alternatively, the element preferably comprises at least one ceramic membrane having a pore size of 30 nm. The element may comprise a plurality of ceramic discs with different pore sizes, including those with pore sizes of 3 nm and 5 nm. The element may comprise a plurality of ceramic discs with different pore sizes, including those with pore sizes of 5 nm and 7 nm. The element may comprise a plurality of ceramic discs with different pore sizes, including those with pore sizes of 3 nm and 30 nm. The element may comprise a plurality of ceramic discs with different pore sizes, including those with pore sizes of 3 nm, 5 nm, 7 nm, and 30 nm.

[0209] According to yet another preferred embodiment, the filter element for carrying out a continuous concentration process includes an ultrafiltration device comprising a membrane in the form of a polymer membrane, such as polyethersulfone or regenerated cellulose.

[0210] protein recycling

[0211] The concentration of protein in a sample (e.g., in the supernatant or its subsequently purified product) can be measured by any method known to those skilled in the art. Examples of suitable assays include high-performance liquid chromatography (HPLC; e.g., size exclusion HPLC), enzyme-linked immunosorbent assay (ELISA), and immunoturbidimetry. These techniques can be used to evaluate sample purity. Furthermore, gel electrophoresis, such as SDS-PAGE with staining and densitometric methods, can be used to evaluate sample purity and detect the presence of contaminating proteins. Reducing agents such as dithiothreitol can be used in SDS-PAGE to cleave any disulfide-linked polymers.

[0212] The starting material containing immunoglobulin G preferably has a total protein concentration of about 0.5-6.5% w / v, more preferably about 1.0-4.0% w / v, even more preferably about 1.5-3.0% w / v, and most preferably about 1.8-2.5% w / v, for example about 2.0% w / v.

[0213] In one embodiment, the liquid comprises a buffer solution containing one or more of sodium acetate, phosphate, and citric acid. In one embodiment, the phosphate is sodium phosphate, such as anhydrous sodium dihydrogen phosphate. Preferably, a buffer solution with low conductivity is used, such as a buffer solution with a conductivity of less than 5 mS / cm, more preferably less than 4 mS / cm, and more preferably 0.01 mS / cm to 4 mS / cm.

[0214] In any implementation, the temperature at which the conductivity of the solution is measured can be between about 4°C and 37°C, preferably, said temperature is between about 20°C and 25°C (room temperature).

[0215] The method according to the invention allows for the recovery of the protein of interest from a protein-containing precipitate (e.g., paste) in high yield. In the post-concentration stage, the recovery rate (ultrafiltration product) is typically at least 75% (w / w), at least 80% (w / w), at least 85% (w / w), at least 90% (w / w), at least 95% (w / w), preferably at least 96% (w / w), more preferably at least 97% (w / w), and most preferably up to 98% (w / w), defined as the total amount of immunoglobulin G in the final filtered solution compared to the total amount of immunoglobulin G in the starting material.

[0216] The following is an example illustrating how the recovery rate of immunoglobulin G content can be calculated according to the present invention. The first step involves determining the IgG content in the starting material (i.e., a precipitate containing protein; completely dissolved), and the second step involves determining the IgG recovery rate obtained when using the continuous extraction method or system of the present invention.

[0217] As a first step, the protein-containing precipitate (approximately 50 g per experiment) is dissolved in a buffer (e.g., 0.12 M–0.25 M phosphate buffer, pH 7.6–8.0) to give a final dilution factor of 20 (a final dilution ratio of 1:20 or 1:19 by weight). After resuspending using an impeller mixer for 2 h, the suspension is centrifuged at 4500 G. This results in a first supernatant and a first precipitate. The volume of the supernatant can be determined using standard methods, and the IgG content of the supernatant can be determined, for example, by turbidimetry. The resulting precipitate is resuspended and treated again with the same buffer as described above to give a final dilution factor of 20 (1 part of the obtained supernatant: 19 parts of new buffer). The volume of the resulting supernatant and the IgG content are determined again. This process is repeated, for example, 5 times, or as many times as needed, to ensure that the IgG content in the final supernatant is below 10 mg / L (the quantification limit is approximately 3.6 mg / L). This process ensures complete or optimal dissolution or extraction of IgG from protein-containing precipitates using buffer. Repeat this experiment several times (in this case, 12 separate experiments). Repeating this process with different starting precipitates produced by fractionation of source plasma yields similar, reliable results. Table 1 below shows the total protein and IgG content recovered from protein-containing precipitates.

[0218] Table 1: Determination of IgG content in protein-containing precipitates (total dissolution)

[0219]

[0220] In the second step, the same protein-containing precipitate is used for experiments, employing the continuous extraction and separation method or system according to the invention. A total volume of 1 kg of the protein-containing precipitate (precipitate A) is dissolved in a buffer solution (e.g., 10 mM phosphate, 10 mM acetate & 2 mM citric acid) for 30 minutes to provide an initial suspension with a first dilution ratio of 5 (1:6 by weight; or a first dilution factor equal to 6 (1:6)). The pH of the suspension is 4.6. The suspension is transferred from the first tank to the first filtration unit for the continuous extraction and separation process. For every 100-200 ml of filtrate collected, 100-200 ml of fresh buffer solution (or recirculation buffer after the UF step (i.e., the second permeate)) is added to the first tank, thereby maintaining a constant volume of the filtered suspension in the first tank. Filtration is terminated after 4 hours, thus expecting a total protein concentration of less than 0.1 g / L and / or an IgG concentration of less than 50 mg / L in the suspension.

[0221] Table 2 below shows the recovery rate of IgG.

[0222] Table 2: Determination of IgG recovery rate using the continuous extraction and filtration system of the present invention

[0223] Starting precipitate used Precipitate A Amount of initial precipitate (kg) 1.0 First dilution factor (paste: total amount (by weight)) 1:6 Total protein content (at first dilution factor) (g / kg paste) 116.4 IgG level (at first dilution factor) (g / kg paste) 61.7 Final dilution factor (after continuous extraction & filtration) 1:31 Total protein content (after reaching the final dilution factor) (g / kg paste) 168 IgG level (after reaching the final dilution factor) (g / kg paste) 78.9 IgG level (after ultrafiltration) (g / kg paste) 78.6

[0224] In this example, experiments were conducted offline using a continuous extraction and filtration unit to demonstrate increased yields of total protein and IgG. "Offline" indicates that the added buffer was not obtained from the second filtration unit. Ultrafiltration in the second filtration unit was performed separately. Table 2 also shows a lower amount of IgG at the first dilution factor (1:6 = 61.7 g / kg) compared to the amount of IgG at the final (second) dilution factor (1:31 = 78.9 g / kg). This is due to the fact that not all IgG is extracted or dissolved in the buffer at once, but rather through a period of time or repeated dissolution processes. Therefore, the continuous extraction and filtration method according to the invention improves the IgG yield.

[0225] As shown in Tables 1 and 2 above, an immunoglobulin G recovery rate of at least 95% was achieved according to the present invention. The recovery rate (based on the continuous extraction and filtration according to the present invention) was calculated by multiplying the ratio (total amount of IgG in the continuous filtrate: average amount of total dissolved IgG in Table 1) by 100.

[0226] Total IgG level (after reaching the final dilution factor) = 78.9 g / kg

[0227] Average total IgG extraction (Table 1) = (76.5 + 85.3) / 2 = 80.9 g / kg

[0228] IgG recovery rate = 78.9 / 80.9 x 100% = 97.53%

[0229] Therefore, it is shown that at least 95% or about 98% IgG recovery rate can be achieved according to the present invention.

[0230] High recovery rates in this early processing step (before further downstream processing steps) are a prerequisite for achieving even higher yields in the final batch stage. This invention utilizes an extraction process in which protein-containing precipitates (e.g., pastes) are suspended virtually at high dilution factors (e.g., 40-70; 1:40 and 1:70). For example, resuspending 1 kg of protein-containing precipitate (e.g., paste) in 3 kg of liquid (e.g., buffer) results in an initial suspension with a first dilution factor of 4 (1:4). Recirculation of a 66 kg buffer feed stream results in a final dilution factor of 70 (1:70). The extraction method used in this invention allows for the release of higher amounts of immunoglobulin G into the suspension / solution, thereby altering the equilibrium (discussed below) and allowing for more efficient separation of immunoglobulin G from the suspension.

[0231] In a preferred embodiment, a crude protein precipitate containing immunoglobulin G (i.e., a protein-containing precipitate) is suspended in a buffer solution to generate a starting suspension. In some embodiments, the buffer solution may contain acetate or phosphate, or additional citric acid.

[0232] In the most preferred embodiment, the product of extraction and filtration of a suspension or solution enriched with immunoglobulin G contains human immunoglobulins, wherein at least 95% or up to 98% of the immunoglobulin G content is recovered from the starting precipitate, or an immunoglobulin G protein concentration of less than 0.1 mg / ml, preferably less than 0.05 mg / ml, is detectable in the final suspension after reaching a second dilution factor. The approximate distribution of immunoglobulin G subclasses generally resembles the average subclass distribution in human plasma.

[0233] Furthermore, typically 1 kg of precipitate A (the protein-containing precipitate) contains approximately 170 g of total protein (range: 150-190 g protein / kg precipitate). The total protein consists of approximately 50%-60% IgG (therefore, range: 75-95 g / kg precipitate). According to a method of the present invention, when an IgG recovery rate of approximately 98% is achieved, this means that a total amount of 73.5-93.1 g / kg of IgG is obtained from the protein-containing precipitate (precipitate A).

[0234] The method described above for calculating the percentage of IgG recovery can be further applied to this invention, including collecting the first filtrate in a second vessel before precipitation, and then further continuously extracting and separating it.

[0235] In chemistry, such as in protein separation, Le Chatelier's principle, or "equilibrium law," can be used to predict the effects of changes in conditions on chemical equilibrium. When any system in equilibrium experiences a change in concentration, temperature, volume, or pressure, the system will readjust itself to counteract (partially) the effect of the change and establish a new equilibrium. In other words, whenever a system in equilibrium is disturbed, it will adjust itself in a way that negates the effects of the change. For example, at equilibrium, the concentrations of immunoglobulins in the suspensions on both sides are constant. If a small amount of immunoglobulin is removed from the reaction at equilibrium, the change in immunoglobulin concentration will shift the equilibrium to the side with the reduced concentration change. According to Le Chatelier's principle, the system will attempt to partially resist the change affecting the original equilibrium state. In turn, the reaction rate, the extent of the product, and the yield will change accordingly with respect to the effect on the system.

[0236] If the system is in equilibrium and the concentration of one of the substances involved in the reaction increases, the system will readjust to decrease the concentration of that substance. Therefore, the reaction will proceed by consuming some of the increased concentration. Similarly, if the concentration of some substance decreases, the reaction will proceed to compensate for the loss of concentration.

[0237] In other words, by continuously removing immunoglobulins (e.g., IgG) from the system while simultaneously reducing the immunoglobulin concentration in the suspension solvent through dilution, this results in an increase in the removal of immunoglobulins from one phase of the suspension, i.e., precipitation to the liquid phase. By repeating this process, virtually all immunoglobulins included in the precipitate of the suspension can be extracted from the protein-containing precipitate, particularly with a high final dilution factor of at least 30 as disclosed in this invention, preferably, for example, 40 (1:40) or higher, and can be further aided by, for example, the suggested buffer composition, or by additional aiding with a higher pH, to maximize the recovery of immunoglobulin G from the protein-containing precipitate. Compared to the prior art, the method and system of this invention allow for the recovery of virtually all immunoglobulin G from protein-containing precipitates (e.g., pastes or precipitates).

[0238] The ultrafiltration product can then undergo further processing, such as chromatographic steps, virus inactivation steps, concentration, and formulation, to allow the final product to be administered, for example, to the human body. The final product can be used to treat immune conditions, particularly autoimmune diseases and certain neurological disorders. These conditions include rheumatoid arthritis, systemic lupus erythematosus (SLE), antiphospholipid syndrome, immune thrombocytopenic purpura (ITP), Kawasaki disease, Guillain-Barré syndrome (GBS), multiple sclerosis (MS), chronic inflammatory demyelinating polyneuropathy (CIDP), multifocal motor neuropathy (MMN), myasthenia gravis (MG), vesicular disease, scleroderma, dermatomyositis, polymyositis, Alzheimer's disease, Parkinson's disease, Alzheimer's disease associated with Down syndrome, cerebral amyloid angiopathy, Lewy body dementia, frontotemporal degeneration, or vascular dementia. Furthermore, the final IVIg and SCIg products can be used in other medical procedures such as cell and organ transplantation.

[0239] Therefore, it should be reiterated that, due to its unique design, the first processing unit according to the invention (and / or the third processing unit according to the second aspect of the invention) provides a continuous extraction and separation process, particularly a filtration process. The first processing unit can be equipped with a dynamically rotating filter element, such as a ceramic-based membrane disc. Compared to conventional cross-flow techniques, rotary filtration allows for extreme cross-flow velocities (due to its highly efficient cleaning of the filter surface) and has very low energy consumption. The cross-flow effect (tangential flow cleaning of the filter surface) is generated by the rotation of the filter disc rather than by pumping through a large volume. Ceramic filter discs offer better resistance to chemical and thermal stress, high filtration flux, and a very long service life, and can be regenerated by backwashing or thermal steam sterilization.

[0240] Dynamic cross-flow filter elements and systems

[0241] Dynamic cross-flow filtration, such as rotary filtration, provides maximum filter efficiency. The cross-flow effect (tangential flow cleaning of the filter surface) is generated by the rotating filter discs, rather than by pumping a large volume across a fixed membrane as used in conventional (static) cross-flow filtration systems. Compared to conventional cross-flow techniques, the extreme cross-flow velocities generated on the rotating filter disc surface ensure highly efficient cleaning of the filter surface while consuming very little energy.

[0242] In the dynamic cross-flow filtration unit and system of this invention, the rotating ceramic filter disc is typically mounted in a pressurized housing. The disc's design incorporates internal drainage channels. The filtrate is transported from the outside of the disc to the inside. The rotation of the disc generates shear forces on the membrane surface. This technique avoids the accumulation of filter cake, resulting in high filtration flux. Some key parameters of rotary filtration are the rotational speed of the ceramic filter disc and the solids content (liquid concentration due to filtrate removal).

[0243] Temperature affects the viscosity of protein solutions and also the flux when filtered through membranes. The temperature of the starting suspension used in the method of the present invention should be in the range of 0°C to the temperature at which the relevant protein denatures. Temperatures are typically in the range of about 10°C to about 50°C. In a particular embodiment, the temperature is in the range of about 18°C ​​to about 35°C. According to a preferred embodiment, the second suspension tank (i.e., a tank containing fatty acids, preferably caprylic acid, said fatty acids being used to precipitate one or more impurities in the first permeate of the first aspect of the present invention or the second retention of the second aspect of the present invention enriched with the protein of interest) is at a temperature of about 4°C to about 40°C. More preferably, the fatty acids, preferably caprylic acid, are combined with the first permeate of the first aspect of the present invention or the second retention of the second aspect of the present invention in the second suspension tank at a temperature of about 25°C to about 38°C, optionally about 27°C, about 32°C, or about 37°C. Optionally, the permeate or retention enriched with the protein of interest and the fatty acids, when combined, are at a temperature of about 4°C to about 40°C, preferably about 25°C to about 38°C, more preferably about 32°C. In other words, although the permeate can circulate in the filtration unit at one temperature, the temperature of the permeate can be increased before and / or during incubation with fatty acids in the second suspension tank.

[0244] According to yet another preferred embodiment, the temperature in the first filtration unit is controlled, preferably between 2°C and 25°C, more preferably between about 2°C and 10°C. Such a temperature ensures optimal extraction and separation processes while maintaining the bioreactivity of the protein of interest throughout the process.

[0245] Filtration is performed at a transmembrane filtration pressure equal to or lower than the level the membrane can withstand, depending on the membrane material used herein, for example, with a pressure of about 0.2 to about 3 bar. The transmembrane pressure is typically 0.1 to 2.5 bar, preferably 0.2 to 2.4 bar, more preferably 0.4 to 2.0 bar, 0.5 to 1.8 bar, 0.6 to 1.6 bar, 0.6 to 1.5 bar, 0.7 to 1.5 bar, and most preferably 0.8 to 1.5 bar. According to another embodiment, a pressure of up to 2 bar is provided to the filtration unit, preferably 0.1 to 2.0 bar, or about 1.5 bar, 1.0 bar, or 0.5 bar.

[0246] According to another embodiment, the continuous extraction process in the filtration unit is further aided by adjusting the flow rate and / or residence time of the suspension or solution entering the filtration unit, and / or the flow rate of the retainer / raffinate containing impurities / precipitates, and / or the flow rate of the first permeate / filtrate enriched with the protein of interest. The filtration unit is adapted to separate impurities / precipitates from the first and second suspensions. For example, in one embodiment, the linear velocity of the suspension or solution entering the pressure vessel (filtration processing unit) can be about 0.27-1.66 m / s. In another example, the linear velocity of the retainer containing impurities / precipitates can be 0.25-1.33 m / s. In another example, the linear velocity of the permeate / filtrate enriched with the protein of interest can be 0.03-0.33 m / s. The linear velocity multiplied by the cross-sectional area yields the volumetric flow rate. Furthermore, turbulence can be generated in the first processing unit due to the speed of the rotating filter disc, where the velocity (sometimes referred to as the tangential velocity) can be about 1-7 m / s. According to one embodiment of the invention, the speed of the rotary disc filter is 1-10 m / s. In a preferred embodiment of the invention, the speed of the rotary disc filter is 5-7 m / s. More preferably, the speed of the rotary disc filter is 7 m / s at 60 Hertz (800 rpm). The rotational speed of the rotating cross-flow filter element is about 600 rpm (50 Hz) to about 1600 rpm (100 Hz), preferably about 800 rpm (60 Hz) to about 1200 rpm (80 Hz), preferably about 800 rpm (60 Hz), about 1000 rpm (70 Hz), or about 1200 rpm (80 Hz). As used herein, rotational speed in Hz refers to the speed of the electric motor. It can be correlated with speed in rpm using a suitable calibration curve.

[0247] This method allows for continuous extraction and separation processes to maximize the recovery of the protein of interest from the starting precipitate / feed (i.e., the first suspension) or the second suspension. Due to the extraction process, virtually all of the protein of interest is extracted from the protein-containing precipitate and recovered in subsequent stages. This method also allows for the recirculation of liquids or diluents such as buffers or water in a closed system, thus maintaining the volume of liquid throughout the process while reducing footprint (i.e., large tank volume).

[0248] In a further embodiment, the invention includes a backwashing step in conjunction with dynamic cross-flow filtration to flush out contaminants that may have accumulated in the system. Typically, the methods and systems of the invention include alternating filtration and backwashing, such that filtration is temporarily paused (i.e., the feed pump stops), while backwashing occurs at regular intervals, wherein liquid flows backward into the filtration system.

[0249] It should be understood that the frequency, duration, and flow rate of backwashing can be adjusted to maximize filtration efficiency and filtration time before backwashing is required.

[0250] In some preferred embodiments, the frequency of backwashing (and therefore the filtration time) is determined based on the total protein concentration or amount of impurities in the starting material. Therefore, it should be understood that backwashing is required more frequently when filtering the first suspension compared to filtering the second suspension, which has relatively fewer impurities. In other words, as the total protein concentration and turbidity of the filtrate decrease, the frequency of backwashing intervals also decreases (i.e., the time interval between backwashes increases, allowing for a longer filtration period before backwashing is required).

[0251] Protein concentration and turbidity of the filtrate can be monitored using various methods known in the art. In some embodiments, the method and system of the present invention include the use of an online detection unit capable of measuring protein concentration and / or turbidity as the filtrate enters and / or exits the filtration unit. In a further embodiment, a dual-wavelength spectrophotometer can be used to facilitate the simultaneous assessment of protein concentration (e.g., by detecting the absorbance of the solution at a wavelength suitable for detecting protein concentration, e.g., in the range of 260-280 nm, preferably about 280 nm) and solution turbidity (e.g., by detecting the absorbance of the solution at a wavelength suitable for detecting light scattering caused by the presence of particulate matter, e.g., in the wavelength range of 400 nm-900 nm, preferably about 600 nm-about 880 nm). Dual-wavelength spectrophotometric devices used in conjunction with chromatography and filtration units are well known in the art.

[0252] In some instances, the recoil frequency is 15-second intervals, or intervals of 30, 45, 60, 75, 90, 105, 120, 135, 150, 200, 230, 260, 300, 330, 360, 400, 1000, 2000, 3000, 4000, or longer.

[0253] It should be understood that the duration of the backflush interval will vary depending on the filter area and the number of discs requiring backflush. Larger filter areas typically require a larger volume of backflush buffer, and the backflush duration will also be determined by the flow rate during backflush. Technicians will be able to determine the appropriate duration, frequency, and flow rate for backflushing, depending on the size of the system and the number of discs used. In some instances, backflush durations are approximately 5 seconds, 10 seconds, 15 seconds, 30 seconds, 45 seconds, 60 seconds, or longer.

[0254] It should be understood that, for practical reasons (and to maximize filtration efficiency), the duration of the backwash interval is typically shorter than the duration of the filtration interval. In some implementations, the duration of the backwash interval is at least one-quarter, one-eighth, one-tenth, one-sixteenth, or less of the duration of the filtration interval.

[0255] It should also be understood that the flow rate used during backwashing may be the same as or different from the flow rate used for filtration. In some embodiments, the flow rate during dynamic filtration is in the range of about 15-100 L / h, preferably in the range of about 20-50 L / h (about 200 ml / min to about 1 ml / min, preferably about 300-about 900 ml / min, more preferably about 300-600 ml / min). Preferably, the backwash flow rate is lower than the flow rate used for filtration, such that in some embodiments, the backwash flow rate is in the range of about 100 to about 400 times slower than the flow rate used for filtration.

[0256] In some implementations, backflushing is performed using the same buffer solution contained in the first or second suspension. Alternatively, backflushing can be performed using permeate obtained during concentration (e.g., when using ultrafiltration coupled with dynamic cross-flow filtration to concentrate filtrate obtained from dynamic cross-flow filtration).

[0257] Example

[0258] Example 1

[0259] Processing Step 1

[0260] The starting materials used to prepare the first suspension are precipitate I+II+III or precipitate A. Optimal filtration is defined as the retention of the filter aid in the suspension, while IgG is completely extracted from the paste and recovered in the filtrate in the shortest possible time. Process 1 results in a filtrate free of filter aids. However, impurities such as IgA, IgM, and albumin will still be present. These impurities will be largely removed using octanoic acid (OA) in processing step 2.

[0261] The starting material is resuspended in the required buffer solution (1:5) using an impeller mixer and then transferred to tank 1 (suspension tank). The suspension is pumped to the first processing unit to begin the continuous extraction (CE) process. The unfiltered suspension is returned to the suspension tank as a retainer. This cycle is maintained as the dissolved protein is filtered through a ceramic disc. The filtrate, containing IgG and various impurities, is collected in the second tank (filtrate tank). When the specified volume is reached in the second tank, the concentration process is initiated using systems 2 and 3 (containing ultrafiltration and permeation membranes). The transmembrane pressure (TMP) is adjusted to ensure that the combined permeate flow rate from systems 2 and 3 equals the filtrate flow rate from system 1, thus ensuring a constant volume in the suspension tank during extraction. Once the protein concentration in the filtrate falls below a defined threshold, the filtration unit stops. At this point, a final dilution ratio ≥1:25 is achieved. Concentration continues until the protein concentration in the collection tank reaches 17-20 g / L. During this final concentration process, the permeate is discharged as waste.

[0262] Processing Step 2

[0263] The first concentrate (filtrate / concentrate) containing IgG is transferred to the tank where OA treatment will take place. This can be the same tank used for the starting material suspension (i.e., the suspension tank), but after the tank has been cleaned. After adding OA to the concentrated protein, the OA-suspension is recirculated in System 1 to achieve excellent dispersion of OA and excellent extraction of IgG. The OA-suspension is pumped to the first processing unit to begin the continuous extraction process. The unfiltered suspension is returned to the suspension tank as a retention. This circulation is maintained as IgG is filtered through a ceramic disc. The OA-filtrate containing IgG is collected in a second tank (filtrate tank). When the specified volume is reached in the second tank, the concentration process is started using Systems 2 and 3 (containing ultrafiltration and permeation membranes). The transmembrane pressure (TMP) is adjusted to ensure that the combined permeate flow rate from Systems 2 and 3 equals the filtrate flow rate from System 1, thereby ensuring a constant volume in the suspension tank during extraction. Once the protein concentration in the filtrate falls below a defined threshold, the filtration unit stops.

[0264] Continue concentrating the OA-filtrate until the protein concentration in the collection tank reaches 20 ± 2 g / L. During this final concentration process, percolate is discharged into the waste. Percolation can be performed simultaneously with the final concentration.

[0265] Processing step 1 (continuous extraction and filtration of proteins) has been successfully applied to the continuous extraction system. To reduce development costs and maximize IgG recovery, the minimum amount of OA required for precipitation of impurities must first be determined through laboratory experiments, thereby minimizing the possibility of IgG loss. For this purpose, a first concentrate (process 1) was prepared as described in Example 1. The first concentrate obtained from the CE system was used as the starting material for laboratory-scale OA experiments.

[0266] Example 2

[0267] Reduce caprylic acid (OA) to remove impurities and recover IgG.

[0268] This embodiment describes the process for determining the minimum amount of OA required to remove impurities from the produced filtered IgG composition according to processing step 1. Two important criteria for the minimum amount of OA used for precipitation are: 1) achieving a high IgG yield, and 2) maintaining high quality properties (QAT).

[0269] Experiment 1

[0270] For this experiment, Cohn I+II+III paste (2 kg containing 240 g Celfure C100) was suspended at a first dilution ratio of 1:6 in a buffer solution of 10 mM sodium acetate and 10 mM sodium dihydrogen phosphate dihydrate, pH 4.3-4.4, at 4°C, in the first tank. The suspension in the first tank was stirred with a paddle stirrer at 4°C for approximately 4 hours. Before starting the experiment, the first filtration unit (containing six ceramic filter discs, AD 152 mm, with 0.2 μm membranes; filtration area = 0.216 m²) was prepared. 2 The solution was stored overnight in cold water (1°C). At the start of the experiment, water was drained from the unit, and the suspension was fed into the unit. The suspension was then recirculated between the first tank and the first treatment unit for a few minutes before starting the filtration process. The ceramic filter in the first treatment unit was operated at 80 Hz, 1200 rpm, and a feed pressure of 2.2 bar; the overflow pressure was 0.4–0.8 bar. The permeate from the first treatment unit was collected in the second tank and then fed into the second unit, referred to as the (ultrafiltration / percolation) UF / DF unit. The UF / DF unit is a Novoflow dynamic filtration device containing 15 ceramic discs with 7.0 nm membranes; the filtration area is 0.5 m². 2 The permeate flow rate of the UF / DF system is 125-180 mL / min. The UF / DF system is started once the first filtrate is collected in the second tank. The retaining stream from the UF / DF system is returned to the second tank, while the permeate is returned to the first tank. The volume of permeate returned to the first tank contributes to the total volume of liquid mixed with the paste (i.e., the final dilution factor). In this experiment, the total recirculation volume was ≥25 L / kg paste (i.e., a final dilution factor of 1:25).

[0271] The conductivity of the first concentrate in this experiment was 1.2 mS / cm. The conductivity of the protein solution was adjusted to 8.0 mS / cm using 3.15 M sodium acetate (NaAc) buffer. After LF adjustment, no further dilution of the protein solution was required.

[0272] Table 3 shows the results for the first concentrate.

[0273] Starting materials Cohn components I+II+III Amount of precipitates I+II+III (kg) 2 Initial dissolution ratio (paste + buffer) 1+5 Total protein (g) at the initial dissolution ratio 258.8 Final dissolution ratio (paste + buffer) 1+24 Total protein at the final dissolution ratio 284.8 Total mass of the first concentrate (kg) 19 Total process time (h) 6

[0274] Experiment 2-7

[0275] To determine the minimum required amount of OA, one kilogram of the first concentrate (from Experiment 1) was used for subsequent treatment with different concentrations of OA (0.1; 0.2; 0.25; 0.30; 0.35 to 0.4 g OA / g protein).

[0276] The conductivity of the first concentrate was adjusted to 8.0 mS / cm using 3.15 M sodium acetate buffer. OA addition was performed while homogenizing with an Ultraturax mixer (7000 rpm) for approximately 5 minutes. The solution was then homogenized again with an Ultraturax mixer for 10 minutes and stirred with a paddle stirrer for 120 minutes. The protein solution was then incubated with 0.6 g of calcium phosphate (CAPO) for 30 minutes and then incubated with Celpure 100 at 18 g / kg solution for 15 minutes. Filtration was performed using a CH9 filter. Subsequent washing was performed using 19%–20% of the initial volume of dissolution buffer.

[0277] Further purification was carried out according to the method described in Experiment 8.

[0278] Experiment 8

[0279] An experiment was conducted to compare the IgG yield and purity obtained using CE filtration (Experiments 2-7) with those obtained using the current laboratory process without CE filtration. In short, the current laboratory process involves OA precipitation of precipitates I+II+III or A, followed by clarification and deep filtration on a CH9 filter plate in the presence of a filter aid, followed by ultrafiltration / percolation at low pH and incubation. More specifically:

[0280] The same precipitate sample (i.e., the starting material from Experiment 1) was used in this experiment, with a portion of the precipitate resuspended in buffer solution. The pH range after resuspension was 4.8 ± 0.2 pH units.

[0281] OA was added to the resuspended precipitate at a concentration of ~0.12 mol / L, and the OA suspension was further incubated with stirring.

[0282] Subsequently, a filter aid (0.25 kg Celpure C100 / m) at pH 4.8 was used. 2 Filtration area; In the presence of Advanced Minerals, OA suspension is clarified and deeply filtered through Purafix CH9P (Filtrox) filter plates.

[0283] The solution was diluted to a protein concentration of 20 g / L, and the pH was adjusted to approximately pH 4.0 in the presence of polysorbate 80 (P80). The solution was then subjected to further clarification and deep filtration.

[0284] Table 4: Yields of total protein and IgG after OA precipitation and after incubation at low pH (intermediates). Results are shown using different OA concentrations.

[0285] process g OA / g protein Protein yield (g / g paste) IgG (g / g paste) Experiment 8 0.70 0.047 0.043 Experiment 2 0.10 0.074 0.067 Experiment 3 0.20 0.070 0.064 Experiment 4 0.25 0.068 0.062 Experiment 5 0.30 0.067 0.061 Experiment 6 0.35 0.055 0.050 Experiment 7 0.40 0.053 0.048

[0286] As the amount of OA per gram of total protein decreased (in Experiments 2–7), the yield of IgG increased compared to the yield obtained according to the procedure in Experiment 8. Therefore, the procedure in Experiments 2–7, using less OA per gram of protein, provided a higher IgG yield compared to the procedure in Experiment 8.

[0287] Furthermore, the impurity distribution of the processes in Experiments 2-7 was better than that obtained using Experiment 8.

[0288] For details on impurities such as IgA, IgM and albumin, protease activity and anticomplement activity, as well as QAT parameters such as subclass distribution and molecular size distribution, see Tables 5, 6, 7 and 8.

[0289] Table 5: IgA, IgM, and albumin were present as impurities (intermediates) after incubation at low pH. Results are shown using different OA concentrations.

[0290] process g OA / g protein IgA (g / g paste) IgM (g / g paste) Alb. (g / g paste) Experiment 8 0.70 0.0040 0.0018 <0.0009 Experiment 2 0.10 0.0068 0.0032 0.0057 Experiment 3 0.20 0.0059 0.0025 0.0040 Experiment 4 0.25 0.0056 0.0015 0.0040 Experiment 5 0.30 0.0051 0.0011 <0.0009 Experiment 6 0.35 0.0051 0.0010 <0.0010 Experiment 7 0.40 0.0047 0.0010 <0.0009

[0291] Table 6: Proteolytic activity and ACA (intermediate body) after low pH incubation. Results are shown using different OA concentrations.

[0292] process g OA / g protein Protein hydrolysis activity (nkat / g protein) ACA (%) Experiment 8 0.70 3 20 Experiment 2 0.10 454 15 Experiment 3 0.20 181 13 Experiment 4 0.25 34 8 Experiment 5 0.30 3 7 Experiment 6 0.35 3 4 Experiment 7 0.40 3 7

[0293] Table 7: γ-globulin purity. Results are shown using different OA concentrations.

[0294] process g OA / g protein γ-globulin [%] α / β-globulin [%) albumin[%] Experiment 8 0.70 99.3 0.6 0.1 Experiment 2 0.10 90.3 2 7.7 Experiment 3 0.20 93.3 0.6 6.1 Experiment 4 0.25 91.7 0.7 7.6 Experiment 5 0.30 99.4 0.3 0.3 Experiment 6 0.35 99.5 0.3 0.2 Experiment 7 0.40 99.5 0.3 0.2

[0295] Table 8: Molecular size distribution after low pH incubation (intermediates). Results are shown using different OA concentrations.

[0296] process g OA / g protein Agg(%) Dim. (%) Frag. (%) Mono. (%) Experiment 8 0.70 4.5 3.7 <0.1 91.7 Experiment 2 0.10 10.7 7.9 7.2 74.2 Experiment 3 0.20 4.4 7.2 4.4 84 Experiment 4 0.25 5 9.5 1.3 84.2 Experiment 5 0.30 2.8 7.4 <0.1 89.7 Experiment 6 0.35 2.3 7.5 <0.1 90.1 Experiment 7 0.40 2.1 6.7 <0.1 91.1

[0297] Further experiments were conducted using different batches of concentrate obtained according to the method in Experiment 1, and further purification was performed using reduced amounts of OA (0.25, 0.30, 0.35 or 0.40 g OA) / g protein and including anion exchange chromatography.

[0298] For the following examples, the same procedure as in Experiments 1-8 was used to further purify the low-pH incubation solutions to include Zeta. + Filtration and anion exchange chromatography.

[0299] In short, the additional steps are:

[0300] The filter aid was suspended in the intermediate Ig solution and passed through a Purafix CH9P filter plate followed by Cuno Z. + 90LP filtration. CH9 filtration in the presence of a filter aid represents a form of depth filtration and has been shown to effectively remove pathogens such as PRV and picornaviruses. The Z-shaped pattern following the CH9 filter... + This is to reduce impurities.

[0301] The protein solution was finally filtered online at a moderately acidic pH using a Pall Ultipor N66 0.1 μm membrane filter, followed by a Pall Ultipor VF FTK DV20 virus filter (a filter capable of removing particles, including viruses as small as approximately 20 nm).

[0302] Experiment 9:

[0303] See above (Experiment 1)

[0304] Experiments 10; 11; 12; 13 (see Experiment 2-7)

[0305] OA concentrations of 0.25, 0.30, 0.35, and 0.40 g OA / g protein were used, followed by Zeta... + Further purification was achieved through filtration and anion exchange chromatography.

[0306] Experiment 14

[0307] See Experiment 8

[0308] The results of these experiments are presented in Table 9-11:

[0309] Table 9: Results of the first concentrate from Experiment 9

[0310] Starting materials Cohn components I+II+III Amount of precipitates I+II+III (kg) 2 Initial dissolution ratio (paste + buffer) 1+5 Total protein (g) at the initial dissolution ratio 253 Final dissolution ratio (paste + buffer) 1+24 Total protein at the final dissolution ratio 267.4 Total mass of the first concentrate (kg) 19 Total process time (h) 5

[0311] Table 10: Yields of total protein and IgG: Experiments 10-14. Results are shown using different OA concentrations.

[0312] process g OA / g protein Protein yield (g / g paste) IgG (g / g paste) Experiment 14 0.70 0.043 0.041 Experiment 10 0.25 0.067 0.064 Experiment 11 0.30 0.063 0.060 Experiment 12 0.35 0.061 0.058 Experiment 13 0.40 0.060 0.057

[0313] As OA levels decreased, IgG yield increased compared to the current process (Table 10). Important impurities such as IgA, IgM, and albumin, protease activity, and anticomplement activity, as well as QAT parameters such as subclass distribution and molecular size distribution, are provided in Tables 11, 12, 13, and 14.

[0314] Table 11: IgA, IgM, and albumin as post-chromatographic impurities. Results are shown using different OA concentrations.

[0315] process g OA / g protein IgA (g / g paste) IgM (g / g paste) Alb. (g / g paste) Experiment 14 0.70 <0.0007 <0.0006 <0.0012 Experiment 10 0.10 <0.0009 <0.0008 <0.0015 Experiment 11 0.20 <0.0007 <0.0006 <0.0012 Experiment 12 0.25 <0.0008 <0.0007 <0.0014 Experiment 13 0.30 <0.0007 <0.0006 <0.0012

[0316] Table 12: Post-chromatographic protein hydrolysis activity and ACA. Results are shown using different OA concentrations.

[0317] process g OA / g protein Protein hydrolysis activity (nkat / g protein) ACA (%) Experiment 14 0.70 5 8 Experiment 10 0.25 65 5 Experiment 11 0.30 3 6 Experiment 12 0.35 3 4 Experiment 13 0.40 2 5

[0318] Table 13: Molecular size distribution, after chromatography. Results are shown using different OA concentrations.

[0319] process g OA / g protein Agg(%) Dim. (%) Frag. (%) Mono. (%) Experiment 14 0.70 <0.1 1.5 <0.1 98.3 Experiment 10 0.25 <0.1 1.2 <0.1 98.6 Experiment 11 0.30 <0.1 1.7 <0.1 98.1 Experiment 12 0.35 <0.1 1.7 <0.1 98.1 Experiment 13 0.40 <0.1 1.5 <0.1 98.3

[0320] Table 14: Subclass distribution, after chromatography. Results are shown using different OA concentrations.

[0321] process g OA / g protein IgG1 (%) IgG2 (%) IgG3 (%) IgG4 (%) Experiment 14 0.70 74.8 22.8 0.6 1.7 Experiment 10 0.25 70.1 23.9 2.4 3.2 Experiment 11 0.30 73.2 23.2 1.2 2.5 Experiment 12 0.35 74.1 23.1 0.7 2.1 Experiment 13 0.40 73.7 23.7 0.6 2.1

[0322] Example 3: Effect of temperature on protease levels in purified products

[0323] The precipitate was purified according to procedures 1 and 2 described in Example 1. In short, the starting material (precipitate I+II+III or precipitate A) was resuspended in resuspension buffer (1:5) at 18-22°C using an impeller mixer, and then transferred to a suspension tank.

[0324] The suspension is pumped to the first processing unit to begin the continuous extraction process. Unfiltered suspension is returned to the suspension tank as a retainer. This cycle is maintained as dissolved proteins are filtered through a ceramic disc. The filtrate containing IgG is collected in the second tank (filtrate tank). When the specified volume is reached in the second tank (filtrate tank), the concentration process begins using ultrafiltration and permeation membranes (systems 2 and 3). The transmembrane pressure (TMP) is adjusted to ensure that the combined permeate flow rate from systems 2 and 3 is slightly higher than the filtrate flow rate from system 1, thus ensuring a constant volume in the suspension tank during extraction. The permeate is also used to fill the backflushing tank. Once the protein concentration in the filtrate falls below a defined threshold, the filtration unit stops.

[0325] Adjust the first concentrate (filtrate / concentrate) to a total protein concentration of approximately 15-18 g / L. If necessary, adjust the conductivity of the concentrate to (8 ± 1 mS / cm) and the pH to the range of approximately 4.75-4.85.

[0326] The regulated concentrate is then transferred to a tank for precipitation with octanoic acid (OA). This can be the same tank used for the starting material suspension (i.e., the suspension tank), but after the tank has been cleaned.

[0327] In a series of experiments, OA was added and incubated at different temperatures, such as 22°C, 27°C, 32°C, and higher. The duration of OA addition, mixing intensity, and all other parameters were kept constant to investigate the effect of temperature on the reduction of protease activity.

[0328] After the amount of OA is added, the OA suspension is recirculated in system (1) so that OA can be dispersed and IgG can be extracted.

[0329] The OA suspension is pumped to the first processing unit to begin the continuous extraction-filtration process. The unfiltered suspension is returned to the suspension tank as a retainer. This cycle is maintained as IgG and other globulins are filtered through the ceramic disc. The OA filtrate is collected in a second tank. When the specified volume is reached in the second tank (filtrate tank), the concentration process is initiated using systems 2 and 3. The concentration step can also be started immediately if systems 2 and 3 are filled with buffer.

[0330] The transmembrane pressure (TMP) is adjusted to ensure that the combined permeate flow rate from systems 2 and 3 is slightly higher than the filtrate flow rate from system 1, thereby ensuring a constant volume in the suspension tank during extraction.

[0331] The combined permeate from systems 2 and 3 flows at a slightly higher rate than the filtrate from system 1, which is advantageous for using the permeate as a backflushing buffer.

[0332] The filtration unit will stop once the protein concentration in the filtrate falls below a defined threshold.

[0333] Continue concentrating the OA-filtrate until the protein concentration in the collection tank reaches 20±2 g / L. During this final concentration process, percolate is transferred to waste. Percolation can be performed simultaneously during the final concentration process.

[0334] The concentration and extent of protease activity in the final concentrate were determined using standard methods. In short, serine protease activity was measured by the ability of the protein concentrate to cleave the chromogenic substrate Ile-Pro-Arg-pNA (S-2288). During this reaction, p-nitroaniline (pNA) was released, and its activity was measured at 405 nm using a spectrophotometer. Serine protease activity was measured at 37 °C and pH 8.4. Kallikrein-like activity was measured by cleaving the chromogenic substrate HD-Pro-Phe-Arg-pNA (S-2302). During this reaction, p-nitroaniline (pNA) was released, and its activity was measured kinetically at 405 nm using a spectrophotometer.

[0335] Quantitative assessment of IgA, IgM, and IgG subclasses was determined using standard techniques.

[0336] result

[0337] As summarized in Tables 15 and 17 below, the results show that the OA-suspension reduced the amount of protease present in the final protein product at temperatures above 22°C, preferably in the range of 25°C–37°C, more preferably in the range of 25°C–32°C or 27°C–32°C, and most preferably at about 32°C. Tables 16 and 18 confirm the high levels of recovery of Ig and its subclasses at all test temperatures.

[0338] Table 15: Effect of temperature on protease impurities measured after OA precipitation, filtration, and UF / DF. Starting material: Precipitate A.

[0339] Octanoic acid incubation temperature (°C) Prokaryotic release enzyme activator (PA) (nkat / L) Kallikrein (KK) (ng / mL) 27 476 255 37 105 75

[0340] Table 16: Starting material precipitate A

[0341]

[0342] Table 17: Effect of temperature on protease impurities measured after OA precipitation, filtration, and UF / DF. Starting material components I+II+III

[0343] Octanoic acid incubation temperature (°C) Prokaryotic release enzyme activator (PA) (nkat / L) Kallikrein (KK) (ng / mL) 20 873 276 32 65 <3.9 37 56 <3.9

[0344] Table 18: Starting Material Components I+II+III

[0345]

[0346] It should be understood that the invention disclosed and defined in this specification extends to all alternative combinations of two or more individual features mentioned or apparent from the text or drawings. All these different combinations constitute various alternative aspects of the invention.

Claims

1. A method for extracting immunoglobulins from a precipitate, the method comprising: a) The precipitate containing immunoglobulins is mixed with a liquid in a first vessel to form a first suspension having a first dilution factor, wherein the precipitate is an intermediate product of the plasma fractionation process; b) The first suspension is fed into a first filtration unit, the first filtration unit comprising a dynamic cross-flow filter element adapted to generate a first retention material that consumes immunoglobulins and a first permeate enriched with immunoglobulins. c) Dilute the first suspension in the first tank to a second dilution factor by adding liquid, optionally by introducing the first retainer into the first tank; d) The first permeate enriched with immunoglobulins is recovered in the second tank; e) Add caprylic acid or its salt or ester to a first permeate enriched with immunoglobulins, and adjust the pH or conductivity of the first permeate to precipitate one or more impurities to produce a second suspension. as well as f) Remove precipitated impurities from the second suspension to produce a solution containing immunoglobulins.

2. The method of claim 1, wherein the method includes a step of concentrating and enriching the first permeate containing immunoglobulins prior to the step of precipitating one or more impurities in the first permeate.

3. The method of claim 1 or 2, wherein removing the precipitated impurities from the second suspension comprises: f1) The second suspension is fed into a second filtration unit, the second filtration unit comprising a dynamic cross-flow filter element adapted to produce a second retention containing one or more precipitated impurities and a second permeate enriched with immunoglobulins; f2) Optionally, the second retained stream is fed into a tank containing the second suspension; f3) The second permeate enriched with immunoglobulins is recovered in another container.

4. The method of claim 3, wherein removing the precipitated impurities from the second suspension further comprises: f4) The second permeate in another tank is continuously concentrated in a third filtration unit containing a dynamic cross-flow filter element, thereby producing a third retention enriched with immunoglobulins and a third permeate depleted of immunoglobulins. f5) Optionally, the suspension in the tank containing the second suspension is diluted to a third dilution factor by introducing a third permeate into the tank containing the second suspension; and f6) Return the immunoglobulin-enriched third retention to the third container and / or collect the immunoglobulin-enriched third retention.

5. A method for extracting immunoglobulins from a precipitate, the method comprising: a) The precipitate is mixed with a liquid in a first vessel to form a first suspension having a first dilution factor, wherein the precipitate is an intermediate product of the plasma fractionation process; b) The first suspension is fed into a first filtration unit, the first filtration unit comprising a dynamic cross-flow filter element adapted to generate a first retention material that consumes immunoglobulins and a first permeate enriched with immunoglobulins. c) Dilute the first suspension in the first tank to a second dilution factor by adding liquid, optionally by introducing the first retainer into the first tank; d) The first permeate enriched with immunoglobulins is recovered in the second tank; d1) The first permeate in the second tank is subjected to a continuous concentration process in a second filtration unit containing a dynamic cross-flow filter element adapted to produce a second retention enriched with immunoglobulins and a second permeate depleted of immunoglobulins. d2) Optionally, the suspension in the first tank is diluted to a second dilution factor by introducing the second permeate into the first tank. as well as d3) Return the second retention containing immunoglobulins to the second container and / or collect the second retention containing immunoglobulins; e) Add caprylic acid or its salt or ester to the second retention enriched with immunoglobulins to precipitate one or more impurities to produce a second suspension; as well as f) Remove precipitated impurities from the second suspension to produce a solution containing immunoglobulins.

6. The method of claim 5, wherein removing the precipitated impurities from the second suspension comprises: f1) The second suspension is fed to a third filtration unit, the third filtration unit comprising a dynamic cross-flow filter element adapted to produce a third retention containing one or more precipitated impurities and a third permeate enriched with immunoglobulins. f2) Optionally, the third retained stream is introduced into a tank containing the second suspension to the third dilution factor; f3) The third permeate enriched with immunoglobulins is recovered in another container.

7. The method of claim 6, wherein removing the precipitated impurities from the second suspension further comprises: f4) The third permeate in another tank undergoes a continuous concentration process in a fourth filtration unit containing dynamic cross-flow filter elements, thereby producing a fourth retention enriched with immunoglobulins and a fourth permeate depleted of immunoglobulins. f5) Optionally, the suspension in the tank containing the second suspension is diluted to a third dilution factor by introducing a fourth permeate into the tank containing the second suspension; and f6) Return the fourth retention material enriched with immunoglobulins to another container and / or collect the fourth retention material enriched with immunoglobulins.

8. An industrial-scale method for extracting immunoglobulins from a precipitate in high yield, the method comprising: a) The precipitate is mixed with a liquid in a first vessel to form a first suspension having a first dilution factor, wherein the precipitate is an intermediate product of the plasma fractionation process; b) The first suspension is fed into a first filtration unit, the first filtration unit comprising a rotating cross-flow filter element comprising a filter disc having a ceramic membrane having an average pore size of 5 nm to 5000 nm, the rotating cross-flow filter element being adapted to produce a first retention material that has consumed immunoglobulins and a first permeate that has enriched immunoglobulins. c) Adding liquid in part by introducing the first retainer into the first tank to dilute the first suspension in the first tank to the second dilution factor; d) The first permeate enriched with immunoglobulins is recovered in the second tank; e) The first permeate in the second tank is continuously concentrated in a second filtration unit containing a dynamic cross-flow filter element, thereby producing a second retention enriched with immunoglobulins and a second permeate that consumes immunoglobulins; f) Optionally, the first suspension in the first tank is diluted to a second dilution factor by continuously flowing the second permeate into the first tank; as well as g) Return the second retention material enriched with immunoglobulins to the second container and / or collect the second retention material enriched with immunoglobulins; h) Add caprylic acid or its salt or ester to the second retention medium enriched with immunoglobulins to precipitate one or more impurities to produce a second suspension; as well as i) Remove precipitated impurities from the second suspension to produce a solution containing immunoglobulins.

9. The method of claim 8, wherein the second suspension is fed to a third filtration unit to remove the precipitated impurities from the second suspension, the third filtration unit comprising a dynamic cross-flow filter element adapted to produce a third retainer containing one or more precipitated impurities and a third permeate enriched with immunoglobulins; optionally, the third retainer is fed into a tank containing the second suspension; and the third permeate enriched with immunoglobulins is recovered in another tank.

10. The method of claim 9, wherein removing the precipitated impurities from the second suspension further comprises continuously concentrating the third permeate in the other tank in a fourth filtration unit, the fourth filtration unit comprising a dynamic cross-flow filter element to produce a fourth retention enriched with immunoglobulins and a fourth permeate depleted of immunoglobulins; optionally diluting the suspension in the tank containing the second suspension to a third dilution factor by flowing the fourth permeate into the tank containing the second suspension; and returning the fourth retention enriched with immunoglobulins to the other tank and / or collecting the fourth retention enriched with immunoglobulins.

11. The method of claim 8, wherein the first retention and the second permeate flow continuously into the first tank to dilute the suspension to a second dilution factor.

12. The method of claim 9, wherein the first retainer and the second permeate are continuously flowed into the first tank to dilute the suspension to a second dilution factor, and / or the third retainer is continuously flowed into a tank containing the second suspension to dilute the second suspension to a third dilution factor.

13. The method of claim 10, wherein the first retainer and the second permeate are continuously flowed into a first tank to dilute the suspension to a second dilution factor, and / or the third retainer and the fourth permeate are continuously flowed into a tank containing the second suspension to dilute the second suspension to a third dilution factor.

14. The method of any one of claims 8-13, wherein the filtration unit adapted to produce an immunoglobulin-enriched permeate comprises one or more hollow shafts adapted to collect the permeate, each shaft being connected to at least one filter disc, the filter disc comprising a ceramic membrane.

15. The method of claim 14, wherein the caprylic acid is added in an amount of 0.1 g / g total protein, 0.25 g / g total protein, 0.275 g / g total protein, 0.280 g / g total protein, 0.285 g / g total protein, 0.290 g / g total protein, 0.300 g / g total protein, 0.325 g / g total protein, 0.5 g / g total protein, 0.75 g / g total protein, 1.0 g / g total protein, 1.5 g / g total protein, 2.0 g / g total protein, 2.5 g / g total protein, 3.0 g / g total protein, 3.5 g / g total protein, or 4.0 g / g total protein.

16. The method of claim 15, wherein the caprylic acid is added to the permeate or the retention enriched with immunoglobulins and incubated together at a temperature of 22°C–40°C.

17. The method of claim 16, wherein the caprylic acid is added to the permeate or the retention enriched with immunoglobulins and incubated together at a temperature of 25°C–38°C.

18. The method of claim 17, wherein the caprylic acid is added to the permeate or the retention enriched with immunoglobulins and incubated together at a temperature of 27°C–37°C.

19. The method of claim 18, wherein the caprylic acid is added to the permeate or the retention enriched with immunoglobulins and incubated together at 32°C.

20. The method of any one of claims 1, 5-13, 15-19, wherein the plasma fractionation is an alcohol fractionation of plasma.

21. The method of claim 20, wherein the intermediate product is a Cohn component.

22. The method of claim 20, wherein the intermediate product is selected from: Cohn component I (Fr I), Cohn component II+III (Fr II+III), Cohn component I+II+III (Fr I+II+III), Cohn component II (Fr II), Cohn component III (Fr III), Cohn component IV (Fr IV), Cohn component V (Fr V), Kistler / Nitschmann precipitate A (KNA), Kistler / Nitschmann precipitate B (KN B), and Kistler / Nitschmann precipitate C (KNC), or a combination of one or more thereof.

23. The method of claim 22, wherein the immunoglobulin is human immunoglobulin G (IgG).

24. The method of claim 23, wherein the first suspension is continuously fed into the first filtration unit.

25. The method of claim 24, wherein the first suspension is continuously fed into the first filtration unit until the first suspension has been diluted to at least a second dilution factor.

26. The method of claim 24 or 25, wherein the second suspension is continuously fed into another filtration unit.

27. The method of claim 26, wherein the second suspension is continuously fed into another filtration unit until the second suspension has been diluted to at least a third dilution factor.

28. The method of any one of claims 1, 5-13, 15-19, 21-25, and 27, wherein the dynamic cross-flow filter element is a rotating cross-flow filter element.

29. The method of claim 28, wherein the dynamic crossflow filter element comprises a filter membrane having an average pore size in the range of 5 nm to 2 μm.

30. The method of claim 29, wherein one or more of the filter units comprise a pressure vessel.

31. The method of claim 30, wherein at least the first filtration unit comprises a pressure vessel.

32. The method of claim 30 or 31, wherein the temperature in one or more of the filter units is controlled at a temperature of 2°C to 25°C.

33. The method of claim 32, wherein the transmembrane filtration pressure in one or more of the filtration units is 0.1-2.5 bar.

34. The method of any one of claims 1, 5-13, 15-19, 21-25, 27, 29-31, and 33, wherein the method comprises adjusting the flow rate of the suspension entering the filtration unit and / or the residence time; and / or the flow rate of the first permeate enriched with immunoglobulins.

35. The method of claim 34, wherein the method includes periodic backflushing intervals, wherein the liquid flows in reverse through the filter unit at regular intervals and within a defined time period.

36. The method of claim 35, wherein the time interval between the backwash intervals increases as the turbidity and protein concentration of the liquid in the filtration unit decreases.

37. The method of claim 35 or 36, wherein the duration of the backflush interval is less than the duration of the filtration interval.

38. The method of claim 37, wherein the duration of the backflush interval is at least one-quarter of the duration of the filtration interval.

39. The method of claim 38, wherein the duration of the backflush interval is at least one-eighth of the duration of the filtration interval.

40. The method of claim 39, wherein the duration of the backflush interval is at least one-tenth of the duration of the filtration interval.

41. The method of claim 40, wherein the duration of the backflush interval is at least one-sixteenth of the duration of the filtration interval.

42. The method of any one of claims 1, 5-13, 15-19, 21-25, 27, 29-31, 33, 38-41, wherein the method recovers at least 75% of the immunoglobulins from the protein-containing precipitate.

43. The method of claim 42, wherein the method recovers at least 80% of the immunoglobulins from the protein-containing precipitate.

44. The method of claim 43, wherein the method recovers at least 85% of the immunoglobulins from the protein-containing precipitate.

45. The method of claim 44, wherein the method recovers at least 90% of the immunoglobulins from the protein-containing precipitate.

46. ​​The method of claim 45, wherein the method recovers at least 95% of the immunoglobulins from the protein-containing precipitate.

47. The method of claim 46, wherein the method recovers at least 98% of the immunoglobulins from the protein-containing precipitate.

48. The method of any one of claims 43-47, wherein the product of the method is further processed, comprising one or more chromatographic steps, a virus inactivation step, a concentration and formulation step, such that the final product is suitable for administration to a subject.

49. The method of claim 48, wherein the precipitate comprises a filter aid.

50. The method of any one of claims 5-13, 15-19, 21-25, 27, 29-31, 33, 38-41, wherein the step of precipitating one or more impurities to produce a second suspension further comprises adjusting the pH or conductivity of the second precipitate to reduce the solubility of one or more impurities.

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