Process for concentration of antibodies and therapeutic products thereof

Background of the Invention

Methods for isolating, purifying, and concentrating biological materials are known and include, for example, chromatography, ultrafiltration, and lyophilization, see generally R. Hatti-Kaul et al., Downstream Processing in Biotechnology in Basic Biotechnology, Chapter 9, pp. 187- 211, 2nd ed., Cambridge University Press (2001). Processes for making concentrated preparations of monoclonal antibodies for administration in humans are known; see, for example, US Patent No. 6,252,055, which uses ultrafiltration and recirculates the resulting filtrate.

Some challenges related to available antibody concentration methods include, for example, low fluxes, long run times, large membrane areas, losses and mechanical recovery yield, intensive operator intervention or delivery, low mass transfer rates, energy inefficiencies, and hydraulic pressure limits in concentration equipment. These and other challenges can contribute to high total manufacturing cost and even higher downstream costs for therapeutic drug users.

There is a need for improved processes for preparing highly concentrated protein formulations, such as liquid antibody preparations and therapeutics thereof. Summary

In general terms, the present invention relates to processes for concentrating proteins, such as processes for concentrating an antibody preparation, pharmaceutical formulations containing such a preparation, and their use in therapy in humans or animals.

In certain embodiments, the present invention provides processes for preparing highly concentrated proteins, such as antibody preparations; and therapeutic products prepared by the process, such as therapeutic antibody products. Accordingly, the present invention provides a process for concentrating proteins comprising: a first ultrafiltration of a first antibody preparation to provide a second antibody preparation; a diafiltration of the second antibody preparation to provide a diafiltered intermediate antibody preparation; and a second ultrafiltration of the diafiltered intermediate antibody preparation to provide a third antibody preparation, wherein one or more of the first ultrafiltration, the second ultrafiltration, and the diafiltration are accomplished at elevated temperatures, for example, from about 30°C to about 50°C.

The present invention also provides, in embodiments, a process for concentrating proteins comprising: a first ultrafiltration of a first protein mixture to provide a second protein mixture; a diafiltration of the second protein mixture to provide a diafiltered protein mixture; and a second ultrafiltration of the diafiltered protein mixture to provide a third protein mixture, wherein one or more of the first ultrafiltration, the diafiltration, and the second ultrafiltration are accomplished at, for example, about 45°C.

The present invention also provides, in embodiments, a highly concentrated antibody composition prepared by the processes described above.

Brief description of the drawings

FIG. 1 illustrates an apparatus for accomplishing the preparative process, in embodiments of the present invention. FIGS. 2-17 illustrate various observed and measured process values ​​with respect to various process phases or modes, in embodiments of the present invention.

FIGS. 18 and 19 illustrate the effect of elevated temperature on product quality, in embodiments of the present invention.

FIGS. 20 and 21 illustrate the effect of elevated temperature on bioburden control, in embodiments of the present invention.

FIG. 22 illustrates the effect of elevated temperature on process flow and process time, in embodiments of the present invention.

FIGS. 23-25 ​​illustrate various observed or measured process values ​​with respect to various phases or scaled process modes, in embodiments of the present invention.

Detailed description

Various embodiments of the present invention will be described in detail with reference to the drawings. Reference to various embodiments does not limit the scope of the invention, which is limited only by the scope of the appended claims. Also, any examples mentioned in this specification are not intended to be limiting and merely illustrate some of the many possible embodiments of the claimed invention.

The following expressions, unless otherwise indicated:

Ultrafiltration, ultrafiltration, ultrafiltration, UF, and similar terms refer, for example, to the use of semipermeable synthetic membranes, with appropriate chemical and physical properties, to discriminate between molecules in the mixture, primarily on the basis of molecular size and shape, and achieve the separation of different molecules or achieve the concentration of similar molecules.

Diafilter, diafiltration, diafiltration, DF, and similar terms refer, for example, to the use of an ultrafiltration membrane to remove, replace, or reduce the concentration of salts or solvents from solutions or mixtures containing proteins, peptides, nucleic acids, or other biomolecules.

Transmembrane Pressure or PTM refers to the average applied pressure on the filtrate side of the membrane calculated as PTM [bar] = [(P_TO +P_R. )/2 P_F. where P_TO is the supply pressure, P_R.is the retent pressure, and P_F is the filtrate pressure.

Tangential flow filtration, cross flow filtration, FFT, and similar terms refer to the mode of filtration in which the solute-containing solution passes tangentially through the UF membrane and salts or lower molecular weight solutes pass through the membrane by applying pressure.

“Antibody” is used in the broadest sense and specifically covers intact monoclonal antibodies, polyclonal antibodies, multispecific antibodies (eg, bispecific antibodies) formed from at least two intact antibodies, and antibody fragments, provided they exhibit the biological activity desired. An antibody is a protein generated by the immune system that is capable of recognizing and binding to a specific antigen. Described in terms of its structure, an antibody is a Y-shaped protein consisting of four amino acid chains, two heavy and two light. In a simplified model sufficient for exemplification, each antibody primarily has two regions: a variable region and a constant region. The variable region, located at the ends of the arms of the Y, binds and interacts with the antigen sought. This variable region includes a complementary determining region (CDR) that recognizes and binds to a specific binding site on a particular antigen. The constant region, located in the tail of the Y, is recognized and interacts with the immune system (Janeway, C., Travers, P., Walport, M., Shlomchik (2001) Immuno Biology, 5th Ed., Garland Publishing, New York). A sought antigen usually has numerous binding sites, also called epitopes, recognized by the RDCs on multiple antibodies. Each antibody that specifically binds to different epitopes has a different structure. Thus, an antigen may have more than one corresponding antibody.

The term "about" which modifies, for example, the amount of an ingredient in compositions, concentration of an active, buffer volumes, diavolumes, pore size, apparent molecular weight, molecular weight cut-off, process temperature, process, yields, flow rates, pressures, bioburdens, and similar values, and ranges thereof, when used in the methods of the invention, refers to a variation in the numerical quantity that occurs, for example, through typical administration and measurement procedures used to make concentrates or use solutions; through inadvertent errors in these procedures, through differences in the manufacture, source or purity of the ingredients used to make the compositions or carry out the methods, or similar considerations. The term "about" also encompasses amounts that differ due to the age of a composition with a particular initial concentration or mixture. The term "about" also encompasses amounts that differ as a result of mixing or processing a composition with a particular initial concentration or mixture. Whether or not modified by the term "about" the claims include equivalent amounts.

Consisting essentially of refers to a process for obtaining a concentrated protein composition or antibody composition that includes steps and ingredients that do not materially affect the basic and novel properties of the composition, such as multiplicity of steps or buffer media. Ingredients that materially affect the basic properties of the composition and method of the present invention impart undesirable characteristics including, for example, bioburden, such as undesirable toxicity or irritability associated with contaminants.

The indefinite article "a" or "an" and its corresponding definite article "the" or "the" as used herein shall be understood to mean at least one, or one or more, unless otherwise specified.

The present invention provides, in embodiments, the aforementioned processes and concentrated antibody products thereof.

In embodiments of the present invention, the preparative processes and products thereof may be used to make highly concentrated antibody preparations and similar preparations such as to purify and concentrate proteins or similar substances from natural or synthetic sources, and the products of which may be useful for treating pathological conditions, such as asthma, cancer, psoriasis, inhibiting angiogenesis and similar pathological conditions.

Further exemplified below is how to make and use the preparative processes and products of the invention in embodiments of the processes described above for preparing highly concentrated antibody compositions of the invention.

In embodiments of the present invention, a process for preparing highly concentrated antibody compositions is provided, for example, according to accomplishing the following steps in the stated order, comprising:

a first ultrafiltration of a first antibody preparation, having a concentration of, for example, about 0.1 to about 10 grams per liter (g/L), to provide a second antibody preparation as the retentate, having a concentration of antibodies greater than, for example, about 10 to about 50 grams per liter;

a diafiltration of the resulting second antibody preparation to provide a diafiltered intermediate antibody preparation as the retentate, which has approximately the same concentration as the resulting second antibody preparation retentate, i.e., a diafiltration to achieve buffer exchange to a constant volume; and a second ultrafiltration of the diafiltered intermediate antibody preparation to provide a third antibody preparation as the retentate, having a higher antibody concentration of, for example, about 150 to about 200 grams per liter.

The preparative processes of the invention may further comprise an optional product recovery step(s), for example, and as illustrated and disclosed herein.

In embodiments of the above-described process of the invention, one or more of the first ultrafiltration, the diafiltration, and the second ultrafiltration, may be conducted at, for example, from about 30°C to about 70°C. In embodiments, these steps can be performed at, for example, from about 30°C to about 50°C. In embodiments, these steps can be performed at, for example, from about 35°C to about 50°C. In embodiments, these steps can be performed at, for example, about 45°C, such as about 45°C plus or minus 5°C. Depending on the type of antibody preparation, for processes accomplished at temperatures greater than approximately 70 °C, the preparation may show signs of deterioration, such as denaturation, agglomeration, and the like. For processes achieved at temperatures below about 30⁵C to about 35 °C, flow rates are generally undesirably low and process times are undesirably long, making processes at lower temperatures less attractive for efficient commercial production.

In embodiments, the first antibody preparation can have an antibody concentration of, for example, about 0.1 to about 100 grams per liter (g/L). The antibody concentration is, for example, a common concentration typically available from other preliminary antibody or protein purification methods or steps, such as centrifugation, filtration, chromatography, and the like. The resulting second antibody preparation obtainable from the first ultrafiltration may have an antibody concentration of, for example, about 10 to about 50 grams per liter and, for example, about 20 to about 40 grams per liter, such as 30 grams per liter. . A range for the antibody concentration of the intermediate antibody preparation may depend on, for example, a balance of factors, such as sample volume and sample flow obtainable with a particular buffer containing the second antibody preparation. The intermediate antibody preparation can have an antibody concentration of, for example, about 25 to about 35 grams per liter and the third antibody preparation can have an antibody concentration of, for example, about 170 to about 200 grams per liter. The third antibody preparation, in embodiments, can have an antibody concentration of, for example, about 50 to about 250 grams per liter, such as about 100 to about 230 grams per liter, and about 170 to about 200 grams per liter. per liter, such as 185 grams per liter.

It will be apparent to a person skilled in the art, upon understanding the present invention, that the intermediate antibody preparation and the third antibody preparation comprise the same ultrafiltrate retentate except for, for example, differences in antibody concentration resulting from the first. and second ultrafiltration concentration, and differences in the suspension buffer medium resulting from diafiltration buffer exchange. Thus, if it did exist, there is minimal change in composition, such as degradation, of the target protein or antibody, in embodiments of the present invention.

Conventional ultrafiltration concentration methods can generally have longer times and lower throughput inefficiencies with considerably longer process times such as processes from several days to several weeks, or processes of considerably lower volumes, or both.

In embodiments, the protein concentration process of the invention can be accomplished, for example, in about 1 to 10 hours, preferably in about 2 to 5 hours, and most preferably in about 3 hours. Preferences favor higher flux yields and smaller membrane areas.

In embodiments, the first ultrafiltration can be achieved, for example, in about 35% of the total process time. Consequently, for example, in a concentration and purification process of the invention with approximately 3 hours of total process time, the first ultrafiltration can be achieved in approximately 45 minutes. In embodiments, the second ultrafiltration can be accomplished, for example, in about 15% of the total process time. Thus, for example, in a process of the invention of approximately 3 hours of total process time, the second ultrafiltration can be accomplished in approximately 15 minutes. Diafiltration can be accomplished, for example, in about 50% of the total process time. Consequently, for example, in a process of the invention with about 3 hours of total process time, diafiltration can be accomplished in about 90 to about 120 minutes.

In embodiments, the first ultrafiltration and the second ultrafiltration can be accomplished, for example, with an ultrafilter membrane having a nominal pore size or cut-off molecular weight of from about 5 to about 50 kilo Daltons. Another suitable nominal pore size is, for example, from about 10 to about 40 kilo Daltons. Another suitable nominal pore size, or molecular weight cutoff, is about 30 kilo Daltons.

In embodiments, the first antibody preparation can contain, for example, an antibody having an apparent molecular weight of, for example, about 100 to about 200 kilo Daltons. In other embodiments, the first antibody preparation may contain an antibody having an apparent molecular weight of, for example, about 150 kilo Daltons, such as when the antibody preparation comprises anti-IgE or IgE antibodies, see, for example, US Pat. No. 6,172,213 assigned to Genentech, Inc.

Other suitable antibodies for use in the present invention include antibodies for the treatment of cancer, see generally, for example: PCT/US02/19592; PCT/US01/20118; PCT/US01/25464; PCT/US01 /26626; PCT/US02/28859; PCT/US02/41798; PCT/US02/12206; PCT/US03/11148; PCT/US02/12619; and PCT/US02/33050. Other antibodies suitable for use in the present invention include an anti-CD20 antibody and similar antibodies including chimeric, hybrid, murine, non-human and human forms. See, for example, US Patent No. 6,582,959 (VEGF) and US Patent Application No. 2002/0122797 A1 (human VEGF).

In embodiments, antibodies included within the scope of the invention include recombinant and hybrid antibodies (eg, humanized and "human" antibodies) regardless of species of origin or immunoglobulin class or subclass designation, as well as antibody fragments. (for example, Fab, F(ab')₂, and F_v). See US Patent No. 4,816,567; Mage and Lamoyi, in Monoclonal Antibody Production Techniques and Applications, 79-97, Marcel Dekker, Inc., New York, (1987).

Monoclonal antibodies can also be used and can be isolated from phage antibody libraries using the techniques described in Clackson et al. (1991) Nature, 352:624-628 and Marks et al. (1991) J. Mol. Biol., 222:581-597, for example, monoclonal antibodies include "chimeric" antibodies in which a portion of the heavy and/or light chain is identical or homologous to the sequences corresponding to antibodies derived from or belonging to a particular species. to a particular class or subclass of antibodies, while the rest of the chain(s) is identical or homologous to the corresponding sequences in antibodies derived from other species or belonging to another class or subclass of antibodies, as well as fragments of said antibodies, provided they exhibit the desired biological activity (US Pat. No. 4,816,567; and Morrison et al. (1984) Proc. Nati. Acad. Sci. USA, 81:6851-6855). Chimeric antibodies can include "primatized" antibodies comprising variable domain antigen-binding sequences, derived from a non-human primate (eg, Old World monkey, ape, etc.) and human constant region sequences.

Monoclonal antibodies are highly specific, being directed to a single antigenic site. Also, in contrast to polyclonal antibody preparations that include different antibodies directed to different determinants (epitopes) each monoclonal antibody is directed to a single determinant on the antigen. In addition to their specificity, monoclonal antibodies are advantageous in that they can be synthesized without being contaminated by other antibodies. For this reason, the monoclonal modifier indicates the character of the antibody as being obtained from such a substantially homogeneous population of antibodies, that is, the individual antibodies comprising the population are identical except for possible naturally occurring mutations, which may be present in higher amounts. minor, and should not be construed as requiring antibody production by any particular method. For example, monoclonal antibodies for use in this invention can be made using the hybridoma method first described in Kohler & Milstein, Nature, 256:495 (1975), or by recombinant DNA methods. Other known methods of antibody production are described, for example, in Goding, Monoclonal Antibodies: Principles and Practice, 59103, Academic Press (1986); Kozbor, J. Immunol., 133:3001 (1984). Brodeur, et al., Monoclonal Antibody Production Techniques and Applications, 51-63, Marcel Dekker, Inc., New York (1987).

Several methods have been used to produce monoclonal antibodies (mAbs). Hybridoma technology, which refers to a cloned cell line that produces a single type of antibody, uses cells from several species, including mice (murine), hamsters, rats, and humans. Another method for preparing mAbs uses genetic engineering including recombinant DNA techniques. Monoclonal antibodies formed from these techniques include, among others, chimeric antibodies and humanized antibodies. A chimeric antibody combines the coding regions of DNA from more than one type of species. For example, a chimeric antibody can be derived from the variable region of a mouse and the constant region of a human. A humanized antibody is predominantly from a human, even if it contains non-human portions. Like a chimeric antibody, a humanized antibody may contain a fully human constant region. However, unlike a chimeric antibody, the variable region may be partially derived from a human. Non-human synthetic portions of a humanized antibody are often derived from RDCs in murine antibodies. However, these regions are crucial in allowing the antibody to recognize and bind to a specific antigen.

As described, murine antibodies play an important role in antibody technology. While useful for short-term diagnostics and therapies, murine antibodies cannot be administered to patients long-term without increasing the risk of a deleterious immunogenic response. This response, called a human anti-mouse antibody (AHAR), exists when a human's immune system recognizes the murine antibody as foreign and attacks it. An AHAR response can cause toxic shock or even death. Humanized and chimeric antibodies reduce the likelihood of an AHAR response by minimizing the non-human portions of the administered antibodies. Likewise, humanized and chimeric antibodies have the additional benefit of activating secondary human immune responses, such as antibody-dependent cellular cytotoxicity.

"Antibody fragments" comprise a portion of an intact antibody, preferably comprising the antigen-binding or variable region thereof. Examples of antibody fragments include Fab, Fab', F(ab')2, and Fv fragments; diabodies; linear antibodies, single chain antibody molecules; and multispecific antibodies formed from (a) antibody fragment(s).

An "intact" antibody is one that comprises an antigen-binding variable region as well as a light chain (CL) constant domain and heavy chain constant domains, CP1, CP2, and CP3. The constant domains can be native sequence constant domains (eg, human native sequence constant domains) or amino acid sequence variants thereof. The intact antibody may have one or more "effector functions" which refers to those biological activities attributable to the Fe region (a native sequence Fe region or an amino acid sequence variant Fe region) of an antibody. Examples of antibody effector functions include C1q binding; complement-dependent cytotoxicity; Fe receptor binding; antibody-dependent cell-mediated cytotoxicity (ADCC); phagocytosis; downregulation of cell surface receptors (eg, B cell receptor; BCR), etc.

Depending on the amino acid sequence of the constant domain of their heavy chains, intact antibodies can be assigned to different "classes." There are five major classes of intact antibodies: IgA, IgD, IgE, IgG, and IgM, and many of these classes can be further divided into "subclasses" (¡sotypes), for example, IgG1, IgG2, IgG3, IgG4, IgA, and IgA2. . The heavy chain constant domains that correspond to different classes of antibodies are called α, δ, ε, y, and μ, respectively. The subunit structures and three-dimensional configurations of different classes of immunoglobulins are well known.

In embodiments, the first ultrafiltration concentrates the first antibody preparation to provide the second antibody preparation having an antibody concentration of approximately 30 grams per liter, and the second ultrafiltration concentrates the intermediate antibody preparation (obtained from diafiltration ) to provide the third antibody preparation having an antibody concentration of, for example, about 170 to about 200 grams per liter. First ultrafiltration and second ultrafiltration can be accomplished with the same ultrafilter membrane and, if desired, within the same vessel or process circuit, for example, to minimize handling, losses, leaks, and similar impacts on performance, efficiency, and economy. The first ultrafiltration and the second ultrafiltration can be accomplished with any suitable ultrafilter apparatus or ultrafilter membrane. Many ultrafilter apparatus or ultrafilter membranes, suitable for cross flow filtration (FFT) operation to achieve ultrafiltrations and diafiltrations, are commercially available, such as Millipore, Pall Corp., Sartorius, and other vendors. In embodiments, a suitable ultrafilter membrane can be, for example, any regenerated cellulose component that has a relatively low protein adsorption profile compared to other available ultrafilter membranes, such as polyethersulfone.

The diafiltration step exchanges a first buffer composition present in the first and second antibody preparations for a desired second buffer in the third antibody preparation. In embodiments, the first buffer may comprise, for example, a mixture of aqueous sodium chloride and TRIS buffer, and a second buffer may comprise, for example, a mixture of aqueous histidine chloride and arginine chloride. Diafiltration can achieve buffer exchange at constant volume, constant concentration, or both. In embodiments, diafiltration achieves constant volume, constant concentration buffer exchange. Diafiltration can achieve buffer exchange, for example, from about 5 to about 15 times by volume (ie, diavolumes). Diafiltration can also achieve a buffer exchange of, for example, approximately 8 times in volume (8 diavolumes), ie 8 times the volume of the sample containing the antibody preparation to be exchanged. For example, a 10-liter antibody preparation can be diafiltered with a 5-fold volume (diavolumes) or 50-liter volume of exchange buffer. Exchange volume and preferences for exchange volumes represent a balance of factors, for example, process throughput efficiencies, product purity, governmental and client/patient acceptance standards, and similar standards, and may depend on, for example , the concentration and type of buffer (eg, the first buffer) in the first antibody preparation, and similar considerations.

The first ultrafiltration, the second ultrafiltration and the diafiltration are preferably performed with tangential flow filtration (FFT mode) through an ultrafilter membrane, and the ultrafilter membrane is preferably the same membrane for each step. The yield of the product in the final combined fractions or simply pool (ie, the third antibody preparation) can be, for example, greater than about 70% by weight, such as about 80%. to about 100% by weight based on the weight of the antibodies in the first antibody preparation. The yield of the third antibody preparation can be, in embodiments, greater than about 90% by weight, in embodiments, greater than about 95% by weight, and in embodiments, still greater than about 98% by weight, based on on the weight of antibodies in the first antibody preparation.

The first ultrafiltration can have a recirculation rate of, for example, about 0.83 to 16.67 mL_/s (50 to 1,000 mL/min), and preferably about 1.67 to 16.67 mL/s (100 at 1,000 ml_/min). The recirculation rate can rise according to the available membrane area, eg, 0.46 membrane areas; 1.86; 18.6; and 92.9m² (5, 20, 200, and 1,000 square feet), and similar areas allow for increasingly higher recirculation rates. Thus, a suitable upward recirculation rate, in embodiments, may be, for example, about 0.09 L/s/m² (0.5 L/min/ft²) at approximately 0.9 L/s/m² (5 L/min/ft²). Ultrafiltration and diafiltration can be accomplished, for example, at transmembrane pressures from about 5 to about 344.7 kPa (50 psi). Ultrafiltration and diafiltration can be accomplished, for example, at transmembrane pressures from about 68.9 kPa (10 psi) to about 344.7 kPa (50 psi). In embodiments of the present invention, a process is provided for preparing an antibody concentrate for a more dilute antibody formulation, the antibody concentrate having a minimal bioburden, e.g., less than or less than a detectable limit, such as, less than approximately 100 CFU/mL.

The antibody compositions of the invention can be, for example, a concentrated monoclonal antibody preparation for administration to humans, such as at a concentration greater than or equal to about 100 g/L (mg/mL), such as about 120 to about 170 g/L.

Antibody compositions of the invention can be, for example, immunoglobulins, such as from the IgA, IgD, IgE, IgG and IgM group; subclasses thereof; recombinants thereof;

fragments thereof; and mixtures thereof from any of these. A preferred antibody composition of the invention includes recombinant humanized anti-IgE antibodies. Antibody compositions of the invention may include a buffer. A preferred buffer may be, for example, a mixture of aqueous histidine chloride and arginine chloride.

The preparative processes of the invention are preferably accomplished in the same apparatus and would operate with no intervention or minimal intervention, for example, as illustrated in FIG. 1.

The first antibody preparation can be obtained or prepared using a variety of chemical, physical, mechanical or non-mechanical, or biochemical methods, such as grinding, ultrasonication, homogenization, enzymatic digestion, solvent extraction, centrifugation, chromatography, and the like. , and combinations thereof, see, for example, the above-mentioned citation in relation to R. Hatti-Kaul et al., Downstream Processing in Biotechnology, in Basic Biotechnology, Chapter 9. The third antibody preparation can be further processed, if desired, using, for example, nanofiltration (to remove, for example, divalent ions), osmosis reverse (to extract, for example, monovalent ions), and similar liquid purification methods. The third antibody preparation of the invention can be packaged, stored or used directly. The third antibody preparation can be further processed, if desired, using, for example, additional concentration steps, such as drying, lyophilization, lyophilization/reconstitution, and similar methods. The concentrated third antibody product can be reconstituted at a later stage, if desired, with a suitable liquid.

Referring to the figures, FIG. 1 illustrates an apparatus, in embodiments of the present invention, for achieving the preparative process including an ultrafiltration/diafiltration system (100) having an ultrafiltration-diafiltration (UF-DF) FFT unit (110), which it has a UF-DF membrane (115), which communicates with the recirculation tank (120) that serves as the main feeding and retention tank. In embodiments, the tank (120) may have a temperature control system comprising, for example, an insulation jacket (125), a temperature-controlled heating element or thermostat (126), such as a heating element at resistor rheostat or a heated circulating fluid system including a heater (not shown), a flow regulator (127), such as a recirculation pump, and a suitable heat transfer fluid, such as either water, a glycol or mixtures thereof. To internal circuit components or components that contribute to internal circuit flow or processing, such as pipes, valves, pumps, tanks, and similar components, It can be optionally insulated or additionally adapted for external heating to maintain close control of temperature specifications and prevent temperature drift in the recirculating fluid path within and between the filter chamber (110) and the storage tank. recirculation (120). In embodiments, for example, when the system (100) is performing the first ultrafiltration, such as in a batch feed mode, the system may include an optional feed tank (128) that fluidly communicates with the feed tank. recirculation feed (120) and can be used to, for example, top up, refill, or supplement the depleted liquid phase from the recirculation tank (120).

A pump (130) pumps feed liquid from tank (120) through the UF/DF unit (110) and then recirculates the resulting retentate (the unfiltered or membrane-excluded portion of the feed liquid) to the recirculation tank. (120). A second tank (140) optionally holds and pumps (not shown) a buffer in the main circuit (110-120 loop) during constant volume diafiltration. For example, the rate of addition and the volume of the buffer introduced into the main loop is preferably at the same rate and volume at which the filtrate leaves the main loop through the membrane (115). The buffer tank (140) can optionally be insulated with the insulation jacket (143) and can include the equivalent of the aforementioned heating element and a recirculation pump (not shown). An optional inert gas source (145), such as nitrogen, or other sources of compressed gas can be used, for example, for product recovery, to pressurize retentate return, to exclude oxygen, to exhaust liquid, to clean , to assess the integrity of the membrane, and similar operations. A third tank (160) is used to collect and recover the filtrate leaving the unit (110). Valves (150, 170) can be used as appropriate to regulate the direction and, optionally, the flow rate of the liquid in the system. All valves and pumps can be activated manually, by coordinated computer control, or both. An optional tank (190) and outgoing stream can provide a supplemental wash liquid drain, product recovery, or monitoring system, for example, when equipped with an optional monitoring device (180), such as a liquid proportioner. optical density, optional filter(s) (185) such as a protective filter, product filter, and similar optional subsystems. In embodiments, the main fluid circuit (110-120 loop) may optionally be equipped with an internal line monitoring system.

The concentrated antibody preparations prepared by processes of the present invention can be used for therapeutic administration in humans, including immunoglobulin products, either for intramuscular (IMIG) or intravenous (IVIG) administration. Concentrated antibody preparations of the invention may include a stabilizer, for example, a buffered amino acid salt solution, simple sugars, or similar stabilizers, suitable ion chelating agents, such as EDTA or citrate ion, and combinations thereof, see, for example, Wang, Y.-CJ et al., Parenteral formulations of proteins and peptides: stability and stabilizers, J. Parenteral Sci. Technol., 42, Suppl. S3-S26 (1988). Derwent's published abstract of JP01268646A (AN89359879) states that the application describes an injection preparation of an lgG monoclonal antibody₃ having a concentration of 0.1 micrograms/mL to 100 mg/mL. The subject matter disclosed in these publications is believed to be outside the scope of the present invention.

Preparations according to the invention may be substantially free of aggregates. Acceptable levels of added contaminants would be less than, for example, about 5% by weight, and ideally less than 2% by weight. Levels as low as 0.2% by weight can be achieved, although added contaminants of about 1% by weight are more common. The preparation in embodiments may also preferably be free of excipients traditionally used to stabilize polyclonal formulations, eg glycine and/or maltose.

The present invention can provide a monoclonal antibody preparation for administration to a human characterized in that the antibody in the preparation is a recombinant antibody and can be present at a concentration of 100 mg/mL or higher, preferably higher than 150 mg/mL. The preparation is preferably substantially free of any protein aggregates.

The pH of the pharmaceutical formulations of the invention depends on the particular route of administration. However, in order to maximize the solubility of the antibody in the concentrated solution, the pH of the solution should be different from the pH of the isoelectric point (pl) of the antibody.

In embodiments of the invention, the monoclonal preparation can be used for use in therapy in humans. Various human disorders, such as cancer or infectious diseases, for example, those mentioned above, and immune dysfunction, such as T-cell related disorders, including severe vasculitis, rheumatoid arthritis, systemic lupus, also autoimmune disorders, can be treated. such as multiple sclerosis, graft-versus-host disease, psoriasis, juvenile diabetes, Sjogrens disease, thyroid disease, myasthenia gravis, transplant rejection, inflammatory bowel disease, asthma, IgE related disorders, and other similar disorders or conditions, or combinations thereof.

Therefore, the invention provides in the embodiments for the use of a concentrated monoclonal antibody preparation as described herein in the manufacture of a medicament for the treatment of any of the aforementioned disorders, or similar disorders. Also disclosed is a method of treating a human, afflicted with any of these disorders, which comprises administering to the individual a therapeutically effective amount of a preparation according to the invention. Dosages of such antibody preparations vary with the conditions being treated and the recipient of the treatment but may be, for example, in the range of about 50 to about 2,000 mg for an adult patient, preferably about 100 to about 1,000 mg administered per day or per week for a period between 1 and 30 days, and repeat if necessary. Doses can be administered as single or multiple doses.

Process description.

The formulation step typically exchanges the fully purified drug substance, eg, resulting from ion exchange chromatography, into the final excipient concentration and composition. Purification was generally not achieved at this stage except by small molecule extraction. Emphasis was placed on high throughput, buffer exchange, and formulation step robustness. During formulation via FFT (cross flow filtration), the protein-containing feed solution was pumped through the membrane system and back to the recirculation vessel. The FFT membrane retained the protein (as part of the retentate) while the filtrate (or permeate) was forced through the membrane by pressure. The pressure is called transmembrane pressure (TMP) and is typically controlled using a retentate pressure control valve. The process was generally achieved by a sequence of a first ultrafiltration (concentration), diafiltration (constant volume buffer exchange), and a second ultrafiltration (additional concentration). The number of diavolumes (volumetric equivalents) required to extract buffer components from the process can be easily calculated or determined experimentally.

UF/DF process generally for antilgE.

The pH of an anion exchange chromatography pool was adjusted to approximately pH 6 using 0.5 M aqueous phosphoric acid. The pH adjusted anion exchange pool was formulated by the ultrafiltration/diafiltration (UF/DF) process of the present invention using a membrane having a nominal molecular cutoff of 10,000-30,000 Daltons. Prior to processing, the UF membrane was equilibrated with diafiltration buffer (0.02 M histidine, 0.2 M arginine-HCl, pH 6).

The anion exchange product (anion exchange pool) was then loaded into the system and concentrated to an intermediate concentration by the first ultrafiltration. The pool was then diafiltered (8X or diavolumes) into their formulation (0.02 M histidine, 0.2 M arginine-HCl, pH 6). The pool was then concentrated by a second ultrafiltration to a final volume concentration of >170g/L and recovered through a sterile 0.22 micron filter. The entire UF/DF process was carried out at a temperature set point of approximately 45⁵C. This temperature control was achieved using a temperature control of the combined incoming anion exchange, diafiltration buffer, and the use of a coated recirculation vessel for the UF/DF process as illustrated in Fig. present.

After UF/DF, the recovered pool was diluted (ie, conditioned) to a volume concentration of approximately 150 g/L in 0.02 M histidine, 0.2 M arginine-HCl, 0.04% of polysorbate 20, pH 6 (final formulation). During the conditioning steps the temperature of the batch was allowed to return to room temperature. After conditioning, the formulated batch was again recovered through a sterile 0.22 micron filter.

The UF/DF system can be regenerated with 0.1 N sodium hydroxide and sanitized with 1.4% Minncare®. When not in use, the system can be stored in 0.1 N aqueous sodium hydroxide. UF/DF membranes can be stored, for example, in a 0.1% Roccal®/20% glycerol-water solution between campaigns.

General Ultrafiltration/Diaph¡tration Process Procedures

Operating Parameters: Feed Flow Rate @ 0.09 L/s/m² (0.5 L/min/ft²). A constant impingement pressure control (eg, 68.9 kPa (10 psig)) was used for pre-use and cleaning equilibration, while constant impingement pressure C was used for processing._wall, or constant PTM.

Pre-use equilibration: The following preparations were accomplished on clean Pellicon-2 cassette membranes prior to use to ensure that the membranes were properly equilibrated.

<td>Volume L/m²</td><td>Solution (temperature</td><td>Modality</td>

<td>LL/ft!)</td><td>atmosphere)</td><td></td>

<td></td><td> -</td><td>PUFA</td>

<td> 10,76(1.0)</td><td>WFI</td><td>PUFA</td>

<td> 10,76(1.0)</td><td>DF buffer</td><td>PUFA</td>

<td> 5,38 (0.5)</td><td>DF buffer</td><td>TRFO, 10 minutes</td>

<td> -</td><td> -</td><td>PUFA</td>

Process Use: The following was performed on the resulting anion exchange pool (Q-Pool) obtained from a previous separation step, for example, a Q-Sepharose chromatography step:

A first ultrafiltration (UF1) at a concentration of approximately 5 g/L at a concentration by diafiltration (C_D.F);

Diafiltration (DF1) with four (4) volumes of diafiltration (VD) with the DF buffer;

Subsequent diafiltration (DF2) with four (4) volumes of diafiltration (VD) of DF buffer;

A second ultrafiltration (UF2) at a final concentration (C_Fnai); and optional product recovery. Such steps were typically performed at low dP recycle (mix), eg, 15 minutes. Post Use Deep Cleaning: The following tabulated sequences and conditions were used to deep clean Pellicon-2 cassette membranes immediately after use.________

<td>Volume L/m²(L/ft²)</td><td>Solution (room temperature)</td><td>Modality</td>

<td> 10,76(1.0)</td><td>0.1N NaOH</td><td>PUFA</td>

<td> 5,38 (0.5)</td><td>0.1N NaOH</td><td>TRFO, 30 minutes</td>

<td> -</td><td> -</td><td>PUFA</td>

<td> 10,76(1.0)</td><td>WFI</td><td>PUFA</td>

<td> 5,38 (0.5)</td><td>Minncare® 300ppm</td><td>TRFO, 30 minutes</td>

<td> -</td><td> -</td><td>PUFA</td>

<td> 10,76(1.0)</td><td>WFI</td><td>PUFA</td>

<td> -</td><td> -</td><td>Integrity test at 206.8 kPa (30 psig)</td>

<td> 5,38 (0.5)</td><td>0.1N NaOH</td><td>TRFO, 15 minute storage</td>

Definitions for modes of operation in FFT.

Single pass with open filtering (PUFA). The retentate and filtrate are directed to a drain. Filter valve open.

Total recycling with open filtering (TRFO). The retentate and filtrate are directed to a recycling container. Filter valve open.

Ultrafiltration of batch feed (UF-AT). The retentate goes to a recycle tank, the filtrate goes to the drain, and the incoming mix is ​​transferred to the recycle tank.

Ultrafiltration in batches (UF-T). The retentate goes to the recycling tank, and the filtrate goes to the drain.

Diafiltration (DF). The retentate goes to the recycle tank, the filtrate goes to the drain, and the diafiltration buffer is transferred to the recycle tank.

PD refers to differential pressure.

Product transfer. The ultrafilter membrane unit and recycle tank open into the blender tank. Nitrogen overlay pressure is controlled. The mix is ​​first transferred using the recycle pump and then using a manual peristaltic pump.

Power transfer. The incoming mix is ​​pumped to the recycle tank.

Total recycling with closed filtering (RTFC). The retentate is directed to the recycling container. Filter valve closed.

Q-pool refers to the protein pool or pool of protein fractions originating from, for example, a previous Q-Sepharose chromatography step that has been buffer conditioned, also referred to as conditioned pool.

API refers to water for injection.

EXAMPLES

The following examples serve to more fully describe the manner of using the present invention, as well as reveal the best modes contemplated for carrying out various aspects of the invention. It is understood that these examples do not serve in any way to limit the true scope of this invention, but are presented for illustrative purposes.

Example 1

High concentration formulation of rhuMAb E25

A pilot scale UF system was used to concentrate/formulate rhuMAb E25 (a recombinant human monoclonal antibody with IgE as sought). A Millipore Pellicon ultrafiltration/diafiltration system was assembled with a 61.4 µm regenerated cellulose component membrane²(5.7 sqft), 10,000 Daltons. The system consisted of a diaphragm stand, Waukeskaw model 6 rotary lobe feed pump, 1/2" (1.27 cm) stainless steel recirculation tubing.¹¹) of 316L, and a recirculation container. Pressure gauges/transmitters (Anderson) were located at the inlet (FEED), outlet (RETAINED), and perforated (LEAK) of the membrane holder. Flow dosers (Yokogawa ADMAG) were located at the inlet (FEED) and permeate (FILTER) of the membrane support. A non-return regulating valve (Mikroseal) was located at the outlet of the membrane support to control the retentate pressure and effect transmembrane pressure (TMP). A 40 liter 316L stainless steel lined tank was used as the recirculation vessel. The tank was assembled with a level indicator, top mount agitator (Lightnin), vortex disruptor, and bottom valve (NovAseptic). Temperature control was achieved through the use of temperature modulated glycol fed into the insulation jacket of the tank.

During this stage, the feed fluid rate was set at a constant rate of 0.05 L/s (2.85 L/min) (0.09 L/s/m² (0.5 L/min/ft²)). During all pre-use and post-use operations, the retent pressure control was set at a constant 68.9 kPa (10 psig). During ultrafiltration and diafiltration operations, the system used a Cp control scheme_wall to control the flux through the membrane, see, for example, R. van Reis, et al., "Constant C_wan Ultrafiltration Process Control”, J. of Membrane Science, 130 (1997), 123-140.

Prior to the process, the system stock solution (0.1 N NaOH) was washed in a single step as a drain, first with 21.6 L/m² (2 L/ft²) of purified water (AP) and then 10.8 L/m² (1 L/ft²) of diafiltration buffer (50 mM Histidine / pH 6.0). After the washes, the system equilibrated to recirculate 5.4 L/m² (0.5 L/ft²) of diafiltration buffer for 10 min. The pH of the recirculated solution was checked to confirm equilibration. The tank level was then reduced to a minimum measurable value to minimize dilution of the incoming protein mix. The protein pool resulting from a preceding Q-Sepharose chromatography step was measured to be 3.2 g E25/L and had a volume of 43.1 L. The protein was comprised in a solution of 25 mM TRIS buffer and approximately 200 mM NaCl and the pH was adjusted to 6.2. To begin this step, the protein pool was transferred to the recirculation vessel. In the vessel, the mix was agitated by the top-mounted impeller and the temperature was maintained at room temperature (20-25°C).

During the run, the pool was concentrated in the UF1 mode to 50 g E25/L (approximately 2.8 L). At the start of diafiltration, the recirculation vessel temperature set point was increased to 40 Ό. The rise in temperature and control was affected by flowing warm glycol through the outer insulation jacket of the tank. The pool was then diafiltered with 8 diavolumes of diafiltration buffer. Diafiltration was carried out at a constant volume, which was achieved by combining the flow rate of the buffered solution transferred to the recirculation tank with the flow rate of the filtrate removed from the system. At the end of the diafiltration, the pool was further concentrated in UF2 mode. This phase was performed using an elevated temperature set point of 40 °C. The search for this final concentration was 110 g/L. This was accomplished without the need to reduce the feed flow rate. Subsequently, a low pressure drop mix was performed where the feed pump was controlled to maintain the 34.5-68.9 kPa (5-10 psig) pressure drop across the feed channel. A sample was withdrawn from the recirculation tank and a final volume concentration of approximately 120 g/L was measured. Table 1 summarizes the flow and performance results for UF1, DF(DF1+DF2) and UF2.

Table 1.

<td>process phase</td><td>of</td><td>Normalized Yield g/m²/s (g/ft²/hr)</td><td>Normalized Flow LMH/kPa (LMH/psig)</td>

<td colspan="2">UF1</td><td> 0,04(13.8)</td><td> 6,89 (4.97)</td>

<td colspan="2">DF</td><td> 0,04(13.8)</td><td> 0,42 (2.92)</td>

<td colspan="2">UF2</td><td> 0,54(181.4)</td><td> 0,24(1.64)</td>

The FIG. 2 shows the observed or measured process values ​​over time for the parameters feed flow rate (210), tank temperature (220), dP feed (230), PTM (240), and filtrate flow rate. (250) during the various phases or modalities of the process, including UF1 (10), DF (20), UF2 (30).

The FIG. 3 shows the observed or measured process values ​​over time for E25 concentration (310), flow (320) and PTM (240).

The FIG. 4 shows the observed or measured process values ​​over time for the pressure drop compared to the observed protein concentration for UF1 (410) and UF2 (420) at 37 Ό.

The protein pool was recovered by a series of steps. First, the blend in the recirculation tank was pumped into the tank through a 0.22 micron Millipac 200 sterilizing grade filter, using the rotary lobe feed pump. The protein solution was then flushed from the membrane and tubing unit with a 34.5 kPa (5 psig) nitrogen gas purge applied to the top of the retentate line. The final phase involved purging the tank and feed line, also using 34.5 kPa (5 psig) nitrogen gas.

It is believed that product recovery was improved compared to Example 1 when run at room temperature because the elevated temperature used in one or more of the ultrafiltration, diafiltration, or recovery steps reduced viscous effects. For example, when the temperature control was turned off during product recovery, the system cooled slowly during this operation causing difficulties with recovery from the membrane unit. Alternatively, recovery could be done first from the membrane support and then from the recirculation vessel.

To determine the mass loss during recovery, 1.74 L of the DF buffer was added to the system and recirculated for approximately 5 minutes and recovered using the same sequence as described above. This volume was then analyzed for protein concentration with other pools. Table 2 summarizes the results.

Table 2.

Volume (L) | Concentration | Mass (g) | Performance

<td></td><td></td><td>(g/L)</td><td></td><td>or {Loss} (%)</td>

<td>CombinedQ</td><td> 43,1</td><td> 3,2</td><td> 137,9</td><td> 100</td>

<td>combined recovered</td><td> 0,99</td><td> 120</td><td> 118,8</td><td> 86,1</td>

<td>buffer wash</td><td> 1,74</td><td> 9,8</td><td> 17,1</td><td> 12,4</td>

<td>Filtered out</td><td> 65,3</td><td> 0,04</td><td> 2,6</td><td> 1,9</td>

After processing, the membrane was regenerated using a single wash of 0.1 N NaOH at 10.8 L/m² (1 L/ft²), followed by 5.4 L/m² (0.5 LJft²) of total recirculation for 30 minutes. This was followed by 10.8 L/m² (1 L/ft²) washing with purified water (AP). This was followed by a total recirculation of 300 ppm Minncare® solution for 30 minutes. The system was flushed again with 10.8 b'm² (1 L/ft²) of purified water and finally recirculated for 15 minutes with 0.1 N NaOH and stored. The recovered pool was diluted in 80 g E25/L and conditioned in the final formulation of 50 mM histidine/150 mM trehalose/0.02% polysorbate 20/pH 6.0. Product quality was assessed by size exclusion chromatography (SEC) for both the incoming Q-Mix and the final recovered batch. These data are summarized in Table 3.

Table 3.

<td>Combined</td><td>CTE results (% monomer)</td>

<td>CombinedQ</td><td> 99,8</td>

<td>final batch</td><td> 99,8</td>

Comparative example 2

High concentration formulation of rhuMAb E25 at room temperature

Example 1 was carried out with the following exceptions. Prior to the process, the storage solution of the system (0.1 N NaOH) was washed in a single step in the first drain mode with 21.6 L/m² (2 L/ft²) of purified water (AP) and then with 10.8 L/m² (1 L/ft²) from diafiltration buffer (20 mM histidine / pH 6.0). After washing, the system was equilibrated by recirculating 5.4 Um² (0.5 L/ft²) from the diafiltration buffer for 10 min. The pH of the recirculated solution was checked to confirm equilibration. The level in the tank was then reduced to a minimum measurable value to minimize dilution of the incoming protein mix.

The protein pool from the above Q-Sepharose chromatography step was measured to be 3.3 g E25/L and had a volume of 33.3 L. The protein was comprised in a solution of 25 mM TRIS buffer and approximately 200 mM NaCI and the pH was adjusted to 6.2. To start the step, the protein pool was transferred to the recirculation vessel. In the vessel, the mix was agitated by the top-mounted impeller and the temperature was maintained at room temperature (20-25°C). During the run, the pool was concentrated in the UF1 mode to 50 g E25/L (approximately 2.2 L). The pool was then diafiltered with 8 diavolumes of diafiltration buffer. The diafiltration was carried out at a constant volume, this volume was reached by combining the flow rate of the buffer transferred to the recirculation tank with the flow rate of the filtrate extracted from the system. Diafiltration was also carried out at room temperature. At the end of the diafiltration, the pool was then concentrated in the UF2 mode. The search for this final concentration was 110 g/L. However, due to the high pressure drop across the feed channel, this concentration could not be achieved. In an attempt to achieve this concentration, the feed flow rate was reduced to 0.02 Us (1.4 L/min) at a volume concentration of approximately 80 g E25/L because the pressure drop across the feed channel had reached 344.7 kPa (50 psig). The UF2 continued until a high pressure drop of 344.7 kPa (50 psig) was again reached and the process was stopped. This was followed by low pressure drop mixing where the feed pump was used to maintain a 34.5 kPa (5 psig) pressure drop across the feed channel. Again, the viscous nature of the protein solution made this difficult to achieve as the rotary lobe pump reached excessive pressures. A sample was withdrawn from the recirculation tank and a final volume concentration of approximately 104 g/L was measured. Table 4 summarizes the throughput and flow measured during phases UF1, DF (DF1+DF2) and UF2.

Table 4.

<td>process phase</td><td>Normalized Yield g/m²/s (g/ft²/hr)</td><td>Normalized Flow LMH/kPa (LMH/psig)</td>

<td>UF1</td><td> 0,04(14,5)</td><td> 0,77 (5,31)</td>

<td>DF</td><td> 0,03 (9,5)</td><td> 0,21 (1,47)</td>

<td>UF2</td><td> 0,43 (144,6)</td><td> 0,11 (0,78)</td>

The FIG. 5 shows the observed and measured process values ​​over time for the parameters feed flow rate (210), tank temperature (220), dP fed (230), PTM (240), and filtered flow rate ( 250) during the various phases or modalities of the process including UF1 (10), DF (20) and UF2 (30).

The FIG. 6 shows the observed or measured process values ​​over time for E25 concentration (310), flow (320) and PTM (240).

The FIG. 7 shows the observed or measured process values ​​over time for pressure drop compared to protein concentration observed for UF1 (410) and UF2 (420) at 24 °C.

The protein pool was recovered stepwise. First, the blend in the recirculation tank was pumped from the tank through a 0.22 micron Millipac 200 sterilizing grade filter using the rotary lobe feed pump. The protein solution was then flushed from the membrane and tubing unit with a 34.5 kPa (5 psig) nitrogen gas purge applied to the top of the retentate line. Product recovery in this manner was very poor due to the viscous nature of the solution. The final phase was a purge of the tank and feed line, also using nitrogen gas at 34.5 kPa (5 psig).

To determine the mass loss during recovery, 1.85 L of DF buffer was added to the system and recirculated for approximately 5 min and recovered using the sequence of Example 1. This volume was then analyzed for its protein concentration with the other volumes. combined. Table 5 summarizes the results.

Table 5.

<td></td><td>Volume (L)</td><td>Concentration (g/L)</td><td>Mass (g)</td><td>Yield or {Loss} (%)</td>

<td>CombinedQ</td><td> 33,3</td><td> 3,3</td><td> 109,9</td><td> 100</td>

<td>combined recovered</td><td> 0,77</td><td> 104,4</td><td> 80,4</td><td> 73,1</td>

<td>buffer wash</td><td> 1,85</td><td> 14,7</td><td> 27,2</td><td> 24,7</td>

<td>Filtered out</td><td> 52,2</td><td> 0,03</td><td> 1,6</td><td> 1,5</td>

After the run, the membrane was regenerated using a single 0.1 N NaOH wash of 10.8 L/m²(1 L/ft²) followed by 5.4 L/m² (0.5 L/ft²) total recirculation for 30 minutes. This was followed by a washout of 10.8 L/m² (1 L/ft²) with purified water. This was followed by a total recirculation of 300 ppm Minncare® solution for 30 minutes. The system was washed again with 10.8 Um² (1 L/ft²) of purified water and finally recirculated for 15 minutes with 0.1 N NaOH and stored. The recovered pool was diluted to 80 g E25/L and conditioned in the final formulation of 20 mM histidine/250 mM sucrose/0.02% polysorbate 20/pH 6.0. Product quality was assessed by size exclusion chromatography (SEC) for both the incoming Q-Mix and the final batch recovered. This information is summarized in Table 6.

Table 6.

<td>Combined</td><td>CTE results (% monomer)</td>

<td>Combined-Q</td><td> 99,8</td>

<td>final batch</td><td> 99,8</td>

Example 3

High concentration formulation of rhuMAb E26 with initial feeding batch mode

Example 1 was repeated with the following exceptions. The concentrate/formula was E26 rhuMAb (a recombinant human monoclonal antibody has IgE as target). The products of this example were used in the toxicological evaluation. Millipore Pelicon Ultrafiltration/Diafiltration System Assembled With 1.06 μm Regenerated Cellulose Component Membrane² (11.4-sqft), 30,000 Daltons. The feeding flow rate was set at a constant rate of 0.08 Us (5.0 L/min) (0.08 Us/rn² (0.44 L/min/ft²)). During the ultrafiltration and diafiltration operations the retentate pressure was maintained between about 41.4-55.2 kPa (6-8 psig). The protein pool resulting from the preceding Q-Sepharose chromatography step was measured to reach 6.7 g E26/L and had a volume of 59.3 L.

Since the incoming blend was larger than the recirculation vessel, the UF1 process started in a batch feed mode. In this mode, the Q-Mix was added to the recirculation vessel at approximately the same rate at which the filtrate passes through the FFT membrane to the drain. After the remaining Q-Combine had been transferred to the recirculation vessel, the UF1 process continued in a batch mode. During UF1 the pool was concentrated to 50 g E26/L (approximately 7.9 L). At the start of diafiltration, the recirculation vessel temperature set point was increased to 40 Ό. The increase in temperature and control was affected by flowing warm glycol through the outer insulation jacket of the tank. The pool was then diafiltered with 8 diavolumes of diafiltration buffer. Diafiltration was carried out at a constant volume by combining the flow rate of a buffer transferred to the recirculation tank with the flow rate of the filtrate removed from the system. At the end of the diafiltration, the pool was further concentrated in UF2 mode to a final concentration of 109 g E26/L (3.6L). This phase was also carried out using an elevated temperature set point of 40 Ό. A low pressure drop mix was then performed where the feed pump was controlled to maintain a 5-10 psig pressure drop across the feed channel. Table 7 summarizes the performance and flow results of UF1, DF (DF1+DF2) and UF2.

Table 7.

<td>process phase</td><td>Normalized Yield g/m²/s (g/ft²/hr)</td><td>Normalized Flow LMH/kPa (LMH/psig)</td>

<td>UF1</td><td> 0,08 (26,1)</td><td> 0,54 (3,71)</td>

<td>DF</td><td> 0,06(19,2)</td><td> 0,34 (2,34)</td>

<td>UF2</td><td> 0,52 (174,2)</td><td> 0,26(1,80)</td>

The FIG. 8 shows the observed or measured values ​​over time for feed flow rate (210), tank temperature (220), dP fed (230), PTM (240) and filtrate flow rate (250).

The FIG. 9 shows the observed or measured values ​​over time for E26 concentration (910), flux (920), and PTM (940).

The FIG. 10 shows the observed or measured values ​​over time for pressure drop compared to the observed protein concentration for UF1 (1010) and UF2 (1020).

Just prior to product recovery, a 10 mL sample was analyzed for bioburden detection and typing. A typical rejection limit is 1,000 colony-forming units (CFU) per mL. The results of this evaluation were 1.8 CFU/mL, an adequate value in this step and well below the rejection limit. To determine the mass loss during recovery, 908.1 mL of DF buffer was added to the system and recirculated for approximately 5 min and recovered using the same sequence described above. This volume was then analyzed for protein concentration with other pools. Table 8 summarizes the results.

Table 8.

<td></td><td>Volume (L)</td><td>Concentration (g/L)</td><td>Mass (g)</td><td>Yield {Loss} (%)</td>

<td>Combined-Q</td><td> 59,3</td><td> 6,7</td><td> 397,3</td><td> 100</td>

<td>combined recovered</td><td> 3,41</td><td> 109,1</td><td> 372,0</td><td> 93,6</td>

<td>Buffer wash</td><td> 0,908</td><td> 20,4</td><td> 18,5</td><td> 4,7</td>

<td>Filtered out</td><td> 120</td><td>n/a</td><td>n/a</td><td>n/a</td>

The recovered pool was diluted to 80 g E26/L and conditioned in the final formulation of 50 mM histidine/150 mM trehalose/0.02% polysorbate 20/pH 6.0. Product quality was assessed by size exclusion chromatography (SEC) for the incoming Q-Pool, post-UF1 retentate pool, post-DF retentate pool, and final recovered batch. These data were summarized in Table 9.

Table 9.

<td>Combined</td><td>CET results (% monomer)</td>

<td>Combined-Q</td><td> 99,8</td>

<td>End of UF1</td><td> 99,8</td>

<td>end of the DF</td><td> 99,8</td>

<td>final batch</td><td> 99,8</td>

Example 4

High Concentration Formulation of rhuMAb E26 for Comparison-Toxicological Evaluation of 10kDy30kD

Example 3 was repeated with the following exceptions. Two pilot scale UF systems were used to concentrate/formulate rhuMAb E26. Two Millipore Pelicon Ultrafiltration / Diafiltration systems were assembled with a 1.06 μm regenerated cellulose component membrane² (11.4-sqft), one with a 10,000 Dalton pore size and the other with a 30,000 Dalton pore size. Retent pressures were maintained at approximately 41.4-62.1 kPa (6-9 psig).

10kD process

The protein pool resulting from the preceding Q-Sepharose chromatography step was measured to be 5.85 g E26/L and had a volume of 62.4 L. During UF1, the pool was concentrated to 50 g E26 /L (approximately 7.3 L). At the end of the diafiltration, the pool was further concentrated in the UF2 mode to a final concentration of 107.5 g E26/L (3.4 L). Table 10 summarizes the performance and flow results for UF1, DF, and UF2.

Table 10.

<td>process phase</td><td>of</td><td>Normalized Yield g/m²/s (g/ft²/hr)</td><td>Normalized Flow LMH/kPa (LMH/psig)</td>

<td colspan="2">UF1</td><td> 0,07 (21,8)</td><td> 0,52 (3,6)</td>

<td colspan="2">DF</td><td> 0,05(15,9)</td><td> 0,38 (2,6)</td>

<td colspan="2">UF2</td><td> 0,41 (137,4)</td><td> 0,28(1,93)</td>

To determine the mass loss during recovery, 987 mL of DF buffer was added to the system and recirculated for approximately 5 min and recovered using the same sequence described above. This volume was then analyzed for protein concentration with other pools. Table 11 summarizes the results.

Table 11.

<td></td><td>Volume (L)</td><td>Concentration (g/L)</td><td>Mass (g)</td><td>Yield or {loss} (%)</td>

<td>Combined-Q</td><td> 62,4</td><td> 5,85</td><td> 365,4</td><td> 100</td>

<td>combined recovered</td><td> 3,38</td><td> 107,5</td><td> 361,7</td><td> 98,9</td>

<td>washing of Buffer</td><td> 0,987</td><td> 19,9</td><td> 19,6</td><td> 5,4</td>

<td>Filtered out</td><td> 125</td><td>n/a</td><td>n/a</td><td>n/a</td>

The FIG. 11 shows the measured or observed process values ​​over time for feed flow rate (210), tank temperature (220), dP fed (230), PTM (240) and filtered flow rate (250), in the various modalities or phases of the process including UF1 (10), DF (20), UF2 (30) and low dP (40), for the 10 kD process.

The FIG. 12 shows the values ​​of the process observed or measured over time for the concentration of E26 (1210), flow (1220) and PTM (1240), in the various modalities or phases of the process including UF1 (10), DF (20), UF2 (30) and low dP (40), for the 10 kD process.

FIG. 13 shows the observed or measured process values ​​over time for the pressure drop compared to the observed protein concentration for UF1 (1310) and UF2 (1320), for the 10 kD process.

30kD process

The protein pool resulting from the preceding Q-Sepharose chromatography step was measured to be 5.85 g E26/L and had a volume of 64.5 L. During UF1 the initial pool was concentrated to 50 g E26/ L (about 7.5L). At the end of the diafiltration, the pool was further concentrated in the UF2 mode to a final concentration of 117.5 g E26/L (3.2 L). Table 12 summarizes the performance and flow results for UF1, DF, and UF2.

Table 12.

<td>process phase</td><td>Normalized Yield g/m²/s (g/ft²/hr)</td><td>Normalized Flow LMH/kPa (LMH/psig)</td>

<td>UF1</td><td> 0,08 (25,5)</td><td> 0,58 (4,01)</td>

<td>DF</td><td> 0,05(17,6)</td><td> 0,35 (2,39)</td>

<td>UF2</td><td> 0,54(180,5)</td><td> 0,15(1,57)</td>

To determine the mass loss during recovery, 918 mL of DF buffer was added to the system and recirculated for approximately 5 min and recovered using the same sequence described above. The recovered pool was diluted to 80 g E26/L and conditioned in the final formulation of 50 mM histidine/150 mM trehalose/0.02% polysorbate 20/pH 6.0. Table 13 summarizes the results.

Table 13.

<td></td><td>Volume (L)</td><td>Concentration (g/L)</td><td>Mass (g)</td><td>Yield or {Loss} (%)</td>

<td>CombinedQ</td><td> 64,5</td><td> 5,85</td><td> 377,3</td><td> 100</td>

<td>combined recovered</td><td> 3,20</td><td> 117,5</td><td> 376,0</td><td> 99,6</td>

<td>Buffer wash</td><td> 0,918</td><td> 22,7</td><td> 20,8</td><td> 5,5</td>

<td>Filtered out</td><td> 125</td><td>n/a</td><td>n/a</td><td>n/a</td>

The FIG. 14 shows the observed or measured process values ​​over time for feed flow rate (210), tank temperature (220), dP fed (230), PTM (240) and filtered flow rate (250), in the various phases or modalities of the process including UF1 (10), DF (20), UF2 (30) and low dP (40), for the 30 kD process.

The FIG. 15 shows the values ​​of the process observed or measured over time for the concentration of E26 (1510), flow (1520) and PTM (1540), in the various phases or modalities of the process including UF1 (10), DF (20), UF2 (30) and low dP (40), by the 30 kD process.

The FIG. 16 shows the observed or measured process values ​​over time for the pressure drop compared to the observed protein concentration for UF1 (1610) and UF2 (1620), for the 30 kD process.

Example 5

rhuMAb E25 Liquid Scale

Example 1 was repeated with the following exceptions.

A scale-up UF production system was used to concentrate/formulate a liquid rhuMAb E25 (a recombinant human monoclonal antibody that has IgE as sought). The product can be used in human bioequivalence trials and therapeutic application. Millipore Pericon Ultrafiltration / Diafiltration systems were assembled with a 21.0 μm regenerated cellulose component membrane² (226-sqft), with a pore size of 30,000 Daltons. Each system consisted of a diaphragm stand, Viking S3S rotary lobe feed pump, 3.81 cm (1¹/₂”) and a 250 L recirculation container.

A 250 liter 316L stainless steel lined tank was used as the recirculation vessel. Temperature control of this tank was achieved with a temperature modulated glycol fed to the tank's insulation jacket. The temperature of the glycol fed to the tank jacket was raised or lowered using either a steam fed heat exchanger or a cold glycol supply, respectively.

For this stage, the feed flow rate was set at a constant rate of 1.9 L/s (114 Umin) (0.09 L/s/m² (0.5 L/min/ft²). Diafiltration buffer (20 mM histidine/200 mM arginine chloride/pH 6.0) was prepared in a separate tank. The temperature of this buffer was set at 45 °C prior to the process. This allowed accurate temperature control throughout the process.

Prior to processing, the system stock solution (0.1 N NaOH) was washed a single step under drain mode with 10.8 Um² (1 LVft²) of water for injection (API) and then with 10.8 Um² (1 L/ft²) of diafiltration buffer. After the washes, the system equilibrated by recirculating 5.4 L/m² (0.5 L/ft²) of diafiltration buffer for 10 min. The pH of the recirculated solution was checked to confirm equilibration.

The protein pool resulting from the preceding Q-Sepharose chromatography step was measured to reach 5.2562 g E25/L and had a volume of 1.141 L. The protein was comprised in a solution of 25 mM TRIS buffer and approximately 200 mM of NaCI and the pH was adjusted to 6.2. Just prior to this stage, the temperature set point for this blend was set at 45°C. To begin the step, the protein mix was transferred to the recirculating vessel, through a 0.22 micron sterilizing grade filter to a level of approximately 200 L in the tank. In the vessel, the mix was agitated by a top-mounted impeller and the temperature was maintained at about (40-50 Ό). Since the incoming blend was larger than the recirculation vessel, the UF1 process started in batch feed mode. In this mode, the Q-Mix was added to the recirculation vessel at approximately the same rate that the filtrate passes through the FFT membrane to the drain. After the remaining Q-Combine was transferred to the recirculation vessel, the UF2 process continued in batch mode. During the UF1 mode the pool was concentrated to approximately 30 g E25/L (approximately 200 L). The pool was then diafiltered with approximately 8 diavolumes of diafiltration buffer. During the diafiltration the temperature was maintained between 40° and 50°C. Diafiltration was carried out at a constant volume, which was achieved by combining the flow rate of the buffer transferred to the recirculation tank with the flow rate of the filtrate extracted from the system. Upon completion of diafiltration, the pool was further concentrated in UF2 mode to a final concentration set point of >170 g E25/L (35 L). This UF2 modality phase was also performed at an elevated temperature set point of 45 °C +/- 5 Ό. Subsequently, a low pressure drop mix was performed where the feed pump was controlled to maintain a 34.5-68.9 kPa (5-10 psig) pressure drop across the feed channel. A sample was taken and a scan was performed to confirm the pre-recovery concentration. The concentration of this sample was 219 g E25/L. Table 14 summarizes the throughput and flow measured during the UF1, DF (DF1+DF2) and UF2 phases.

Table 14.

<td>process phase</td><td>Normalized Yield g/m²/s (g/ft²/hr)</td><td>Normalized Flow LMH/kPa (LMH/psig)</td>

<td>UF1</td><td> 0,13(43.8)</td><td> 0,48 (3.34)</td>

<td>DF</td><td> 0,08 (25.9)</td><td> 0,36 (2.46)</td>

<td>UF2</td><td> 0,24 (78.9)</td><td> 0,10(0.66)</td>

Just prior to product recovery, a 30 mL sample was taken and delivered for bioburden detection and typing. The result was &lt;0.13 CFU/mL. The protein pool was recovered by a series of steps. First, the product was dislodged from the membrane in a single pass fashion using 5 L of the DF buffer added to the retentate line. The product was filtered in a recovery tank through a protective 0.69 m sterilizing grade filter.² (7.4ft²) 0.22 micron followed by a final sterilizing grade 0.19 μm filter² (2ft²) of 0.22 microns. The blend in the recirculation tank was then pumped from the tank using the rotary lobe feed pump. Subsequently, the residual protein solution was displaced from the tank and feed line with a 34.5 kPa (5 psig) nitrogen gas purge. The final phase was a purge of the membrane unit, which contained mostly DF buffer from the initial product shift. This phase also used 34.5 kPa (5 psig) nitrogen gas applied to the top of the retentate line. The recovered pool was first diluted to approximately 153 g E25/L using DF buffer. Finally, the pool was conditioned in the final formulation of 20 mM histidine / 200 mM arginine-HCl / 0.04% polysorbate 20 / pH 6.0. The volumes of the recovered pool, diluted pool and conditioned pool (Q-Pool) were analyzed one by one for protein concentration. Table 15 summarizes the results.

Table 15.

<td></td><td>Volume (L)</td><td>Concentration (g/L)</td><td>Mass (g)</td><td>Yield or {loss} (%)</td>

<td>CombinedQ</td><td> 1.141</td><td> 5,2562</td><td> 5.997.3</td><td> 100</td>

<td>combined recovered</td><td> 35,0</td><td> 170,0</td><td> 5.950,0</td><td> 99,2</td>

<td>diluted mix</td><td> 39,0</td><td> 147,0</td><td> 5.726</td><td> 95,5</td>

The FIG. 17 shows the parameters of the feed flow rate (210), tank temperature (220), fed dP (230), PTM (240) and filtrate flow rate (250), during the various phases or modalities. process including UF1 (10), DF1 (20), DF2 (25), UF2 (30) and low dP (50).

Example 6

rhuMAb E25 liquid preparation

Example 5 was repeated with the following exceptions. A production scale UF system was used to concentrate/formulate rhuMAb from liquid E25 (a recombinant human monoclonal antibody to E25 targeting IgE). Millipore Pelicon Ultrafiltration / Diafiltration systems were assembled with a 21.0 μm regenerated cellulose component membrane² (226 sqft), with a pore size of 30,000 Daltons. Each system consisted of a diaphragm stand, Viking S3S rotary lobe feed pump, 3.81 cm (1 Y/) 316L stainless steel recirculation tubing, and a 250 L recirculation vessel. 250 liter 316L stainless steel as a recirculation container. The feed flow rate was set at a constant rate of 1.9 L/s (114 L/min) (0.09 L/s/m² (0.5 L/min/ft²)). During all pre-use and post-use operations, the retent pressure control was set at a constant 68.9 kPa (10 psig). During the ultrafiltration and diafiltration operations, the system used the control scheme of C_wall to control the flux across the membrane. Diafiltration buffer (20 mM Histidine/200 mM Arginine Chloride/pH 6.0) was prepared in a separate tank. The buffer temperature was set at 45 °C prior to the process. This allowed accurate temperature control throughout the entire process. The protein pool resulting from the preceding Q-Sepharose chromatography step was measured to be 5.5438 g E25/L and had a volume of 1.082 L. The protein was comprised in a solution of 25 mM TRIS buffer and approximately 200 mM NaCl and the pH was adjusted to 6.2. Just prior to this stage, the temperature set point for this blend was set at 45°C. To begin the step, the protein mix was transferred to the recirculating vessel, through a 0.22 micron sterilizing grade filter to a level of approximately 200 L in the tank. In the vessel the mix was agitated by a top mounted impeller and the temperature was maintained at room temperature (40-50°C). Since the incoming blend was larger than the recirculation vessel, the UF1 process started in batch feed mode. In this mode, the Q-Combined was added to the recirculation vessel at approximately the same rate at which the filtrate passes through the FFT membrane to the drain. After the remaining Q-Combine was transferred to the recirculation vessel, the UF1 process continued in batch mode. During UF1 the pool was concentrated to ca. 30 g E25/L (ca. 200 L). The pool was then diafiltered with 8 diavolumes of diafiltration buffer. During the diafiltration, the temperature was kept at room temperature between 40 and 50 °C. Diafiltration was carried out at a constant volume, which was achieved by combining the flow rate of the buffer transferred to the recirculation tank with the flow rate of the filtrate extracted from the system. At the end of the diafiltration, the pool was further concentrated in the UF2 mode to a final concentration set point greater than 170 g E25/L (35 L). This phase was also carried out at an elevated temperature set point of 45 °C +/- 5 °C. A low pressure drop mix was then effected where the feed pump was controlled to maintain a 34.5-68.9 kPa (510 psig) pressure drop across the feed channel. A sample was taken and a sweep (scan) was performed to confirm the concentration prior to recovery. The concentration of this sample was 191 g E25/L and the volume of the pool was 31.9 L. A plot of process parameters over time was comparable to that observed and summarized for FIG. 17. Table 16 summarizes the throughput and flow measured during the UF1, DF (DF1+DF2) and UF2 phases.

Table 16.

<td>Phase of process</td><td>Normalized Yield g/m²/s (g/ft²/hr)</td><td>Normalized Flow LMH/kPa (LMH/psig)</td>

<td>UF1</td><td> 0,13(45,1)</td><td> 0,47 (3,21)</td>

<td>DF</td><td> 0,08 (25,9)</td><td> 0,36 (2,51)</td>

<td>UF2</td><td> 0,36(121,4)</td><td> 0,11 (0,79)</td>

Just prior to product recovery, a 30 mL sample was taken and analyzed for bioburden typing. The results of this test were below the detection limit (&lt; 0.13 CFU/mL).

The protein pool was recovered by a series of passages. First, the product was displaced from the membrane in a single step using 5 L of DF buffer added to the retentate line. The product was filtered in a recovery tank through a 0.69 μm protective sterilizing grade filter.² (7.4ft²), 0.22 micron followed by a final sterilizing grade 0.19 μm filter² (2ft²), of 0.22 microns. The blend in the recirculation tank was then pumped from the tank using the rotary lobe feed pump. Subsequently, the residual protein solution was displaced from the tank and feed line with a 34.5 kPa (5 psig) nitrogen gas purge. The final phase comprised a purge in the membrane unit, which contained mostly DF buffer from the initial product displacement. This phase also used 34.5 kPa (5 psig) nitrogen gas applied to the top of the retentate line. The recovered pool was first diluted to approximately 153 g E25/L using DF buffer. Finally, the pool was conditioned in the final formulation of 20 mM histidine / 200 mM arginine-HCl / 0.04% polysorbate 20 / pH 6.0. The volumes of the recovered pool, diluted pool, and conditioned pool were then analyzed for protein concentration. Table 17 summarizes the results. After the process, the membrane was regenerated as described above.

Table 17.

<td></td><td>Volume (L)</td><td>Concentration (g/L)</td><td>Mass (g)</td><td>Yield or {Loss} (%)</td>

<td>CombinedQ</td><td> 1.082</td><td> 5,5438</td><td> 5.998,4</td><td> 100</td>

<td>combined recovered</td><td> 34,95</td><td> 167,08</td><td> 5.839,8</td><td> 97,4</td>

<td>diluted mix</td><td> 38,2</td><td> 152,14</td><td> 5.810,3</td><td> 96,7</td>

Example 7

Effect of elevated temperature on product quality

Samples of E25 at 30 g/L and 150 g/L were stored in histidine and Q buffers at various temperatures for 24 hours. Samples were taken for turbidity measurements and CTE assays. The turbidity vs. temperature results for E25 at 30 g/L in Q buffer are shown in FIG. 18. FIG. 19 shows the amount of soluble E25 aggregate (150 g/L in 50mM histidine buffer, pH 6.0) observed over time and at temperatures of 23 °C, 40 °C, 50 °C, 60 Ό and 70 °C. c. The four time intervals (0 hour, 4 hour, 7.5 hour and 24 hour time) for each of these temperatures are shown as a group of four bars from left to right as 1810 and 1910, in FIGS. 18 and 19. The turbidity of the solution remained essentially unchanged after 24 hours at 60°C. No significant soluble aggregation of E25 was observed below 70°C, suggesting that the product samples were substantially stable up to at least 60°C and at least 24 hours.

Example 8

Effect of elevated temperature on bioburden

Samples of E25 at 30 g/L were inoculated into both arginine and histidine buffers with 10³ of colony-forming units per mL for two study organisms: a Gram-positive strain (Staphylococcus aureusy, and a Gram-negative strain (Pseudomonas chlororaphis). The samples were taken after 1.5 hours and 6 hours. The results shown in the bar graphs in FIGS 20 and 21 indicate that both organisms under study decreased with increasing temperature. The three temperature intervals (25 °C, 40 °C, and 50 Ό temperatures) for each observed time interval are shown as a group of three bars from left to right as 2010 and 2110, in FIGS. 20 and 21. The inoculations shown were carried out in arginine buffer with protein concentrations of 30 g/L.

Example 9

Effect of elevated temperature on the process flow

Samples of E25 at 10 g/L in 0.2M arginine, 25 mM histidine, pH 6.0 buffer were evaluated for their influence on flow vs. transmembrane pressure (TMP) comparison. The FIG. 22 shows that raising the system temperature also increased the process flow during UF/DF operations. Flow peaks were performed at various volume concentrations and three different temperatures of 23°C (2210), 40°C (2220) and 46°C (2230). Mass transfer coefficient and filtrate flow were increased from about 2 to about 3 times, resulting in considerably reduced process times.

Example 10

High concentration formulation of rhuMAB anti-CD20 (2H7)

A pilot scale UF system was used to concentrate and formulate anti-CD20 rhuMAb (2H7; a recombinant human monoclonal antibody). Example 1 was repeated with the following exceptions. Millipore Pelicon Ultrafiltration / Diafiltration systems were assembled with a 1.6 m regenerated cellulose composite membrane² (17.5 sqft), with a pore size of 30,000 Daltons. The system consisted of a diaphragm stand, a Viking S1 L rotary lobe feed pump, 1/2" (Va") 316L stainless steel recirculation tubing, and a 40 L recirculation vessel. retention regulation were from HD Baumann, Inc. The temperature of the glycol fed into the tank's insulation jacket was regulated up or down as necessary using an electric heat exchanger, a cold glycol supply, or both.

During this stage, the feed flow rate was set at a constant rate of 0.14 L/s (8.5 L/min) (approximately 0.09 L's/m2 (0.5 L/min/ft²)). The FIG. 23 Illustrates that the trends of values ​​over time in terms of the feed flow rate (210) scaled from 0 to 20, the pH (212) scaled from 2 to 12, the filtrate flow rate (250) scaled from 0 to 5, the level of the recycle tank (2320) scaled from 0 to 45, and the dP of the retentate (2350) scaled from 0 to 100 during the various phases or modalities of the process including UF1 (10), DF1 (20 ) and UF2 (30).

During ultrafiltration and diafiltration operations the system used constant retentate pressure followed by a constant delta feed/retentate pressure control scheme to control flux through the membrane. Diafiltration buffer (30 mM sodium acetate / pH 4.9) was prepared in a separate tank. The temperature of this buffer was set at 45⁵C prior to the process for accurate temperature control throughout the process. Prior to processing, the system stock solution (0.1 N NaOH) was washed in a single step as a drain, first with 10.8 L/m² (1 L/ft²) of water for injection (API) and then with 10.8 L/m² (1 L/ft²) of diafiltration buffer. After the washes, the system was equilibrated to recirculate 5.4 LJm² (0.5 LVft²) of diafiltration buffer for 10 min. The pH of the recirculated solution was checked to confirm equilibration.

The protein pool resulting from a preceding Q-Sepharose chromatography step was measured to reach 2.31 g 2H7/L and had a volume of 356 L. The protein was presented in a 6 mM HEPES/19 mM free acid solution. of HEPES sodium salt and 25 mM sodium acetate where the pH had been adjusted to 5.3 with 0.5 M acetic acid. Just before the stage, the temperature set point was set at 45^SC. To begin the step, the protein mix was transferred to the recirculating vessel through a 0.22 micron sterilizing grade filter to a level of approximately 40 L in the tank. In the vessel, the mix was agitated by the top-mounted impeller and the temperature was maintained at 40-50⁹c.

Since the incoming blend was larger than the recirculation vessel, the UF1 process started in batch feed mode (see FIG. 23). In this mode, the Q-Mix was added to the recirculation vessel at approximately the same rate at which the filtrate passes the FFT membrane to the drain. After the remaining Q-Combine was transferred to the recirculation vessel, the UF1 process continued in batch mode. During UF1 the pool was concentrated to ca. 50 g 2H7/L (ca. 16 L). The pool was then diafiltered with 10 diavolumes of diafiltration buffer. During diafiltration the temperature was maintained between 40 and 50^QC. Diafiltration was carried out at a constant volume, which was achieved by combining the flow rate of the buffer transferred to the recirculation tank with the filtrate removed from the system. At the end of the diafiltration, the pool was further concentrated in UF2 mode to a target set point of final concentration of 190 g 2H7/L (4.3 L). See in FIG. 23 the addition of constant dP control at 344.7 kPa (50 psig) at the end of this phase. This phase was then carried out at an elevated temperature set point of 45 -C +/- 5⁹C. Subsequently, a low pressure drop mix was performed where the feed pump was controlled to maintain a 137.9 kPa (20 psig) pressure drop across the feed channel. A sample was taken and density measurements were made to confirm the pre-recovery concentration. The concentration of this sample was 189 g of 2H7/L. Table 18 summarizes the performance and flow results.

Table 18.

<td>Phase of process</td><td>Normalized Yield g/m²/s (g/ft²/hr)</td><td>Normalized Flow LMH/kPa (LMH/psig)</td>

<td>UF1</td><td> 0,10(32)</td><td> 0,70 (4,8)</td>

<td>DF</td><td> 0,17(56)</td><td> 0,35 (2,4)</td>

<td>UF2</td><td> 0,80 (267)</td><td> 0,23(1,6)</td>

The protein pool was recovered by a series of steps. First, the product was flushed from a membrane in a single-pass mode using 0.2 L of DF buffer added to the retentate line. The product was filtered into the recovery tank through a 0.22 micron sterilizing grade final filter. The blend in the recirculation tank was then pumped from the tank using the rotary lobe feed pump. The residual protein solution was then displaced from the tank and feed line with a 34.5 kPa (5 psig) nitrogen gas purge. The final phase comprised a purge in the membrane unit containing DF buffer of the initial product shift. This phase also used 34.5 kPa (5 psig) of nitrogen gas applied to the top of the retentate line.

If necessary, the recovered pool was first diluted to approximately 175 g 2H7/L using dilution buffer (30 mM sodium acetate, pH 5.3). Finally, the pool was diluted to a desired concentration of 150 g of 2H7/L and conditioned in the final formulation of 30 mM sodium acetate, 7% trehalose, 0.03% polysorbate 20, pH 5, using a buffer. 7 X conditioning solution (30 mM sodium acetate, 49% trehalose, 0.21% polysorbate 20, pH 5.3).

The volumes of the recovered pool, diluted pool, and conditioned pool were analyzed for protein concentration. Table 19 presents the results.

Table 19.

<td></td><td>Volume (L)</td><td>Concentration (g/L)</td><td>Mass (g)</td><td>Yield or { Loss} (%)</td>

<td>Combined-Q</td><td> 355,81</td><td> 2,31</td><td> 821,92</td><td> 100,0</td>

<td>combined recovered</td><td> 4,64</td><td> 180,02</td><td> 835,3</td><td> 101,6</td>

<td>final combo</td><td> 4,871</td><td> 149,40</td><td> 727,7</td><td> 88,5</td>

Note: Yields include loss due to sampling. The volume and concentration of the pool recovered includes addition of the buffer displacement.

After the run, the membrane was regenerated using a single step wash of 0.1 N NaOH,

10.8 L/m² (1 L/ft²) followed by 5.4 L/m² (0.5 L/ft²) of total recirculation for 30 minutes. This was followed by 10.8 L/m² (1 L/ft²) of purified water washing. This was followed by a total recirculation of 5.4 L/m² (0.5 L/ft²) of 14% Minncare solution for 30 min. The system was flushed again with 10.8 L/m² (1 L/ft²) of purified water and finally recirculated for 15 minutes with 0.1 N NaOH and stored.

Example 11

High concentration formulation of rhuMAb anti-CD20

A pilot scale UF system was used to concentrate and formulate anti-CD20 rhuMAb (2H7) for use in a phase I human clinical study in a GMP manufacturing facility. Example 10 was repeated with the following exceptions.

The protein pool resulting from a preceding Q-Sepharose chromatography step was measured to be 3.729 g 2H7/L and had a volume of 262 L. The protein was comprised in a 6 mM HEPES/19 mM free acid solution. of HEPES sodium salt and 25 mM sodium acetate whose pH was adjusted to pH 5.3 with 0.5 M acetic acid. Just before the stage, the temperature set point for this combo was set at 45⁹C. To begin the step, the protein mix was transferred to the recirculating vessel through a 0.22 micron sterilizing grade filter to a level of approximately 40 L in the tank. In the vessel, the mix was agitated by the top-mounted impeller and the temperature was maintained at 40-50²c.

During UF1 the pool was concentrated to ca. 50 g 2H7/L (ca. 20 L). The FIG. 24 shows the value trends over time: the level of the recycling tank (210) increased from -0.713963 to 295,989, the dP of the retentate (2420) increased from -0.237899 to 98.6629, the flow rate The feed rate (250) scaled from -0.356981 to 147.994, and the filtrate flow rate (2450) scaled from 0.118994 to 49.3315, during the process. The pool was then diafiltered with 10 diavolumes of diafiltration buffer. During diafiltration, the temperature was kept between 40 and 50²C. Diafiltration was performed at a constant volume, which was achieved by combining the flow rate of the buffer transferred into the recirculation tank with the flow rate of the filtrate removed from the system. At the end of the diafiltration, the pool was further concentrated in UF2 mode to a target final concentration set point of 190 g 2H7/L (5.25 L). Note in FIG. 24 the incorporation of the constant dP to the control of

275.8 kPa (40 psig) at the end of this phase. This phase was also carried out at an elevated temperature set point of 45⁹C+/-5²C. A low pressure drop mix was then performed where the feed pump was controlled to maintain a 137.9 kPa (20 psig) pressure drop across the feed channel. A sample was taken and density measurements were made to confirm the concentration prior to recovery. The concentration of this sample was 194 g of 2H7/L. Table 20 summarizes the performance and flow results.

Table 20.

<td>process phase</td><td>Normalized Yield g/m²/s (g/ft²/hr)</td><td>Normalized Flow LMH/kPa (LMH/psig)</td>

<td>UF1</td><td> 0,15(51)</td><td> 0,55 (3,8)</td>

<td>DF</td><td> 0,14(46)</td><td> 0,32 (2,2)</td>

UF2

0,86(286) 10,23 (1,6)___________|

Just before product recovery, a 30 ml_ sample was taken and delivered for bioburden detection and typing. Results were negative (ie, &lt;0.13 CFU/mL).

The protein pool was recovered by a series of steps of Example 10. The volumes of the recovered pool, diluted pool, and conditioned pool were then analyzed for protein concentration. Table 21 presents the results. The membrane was regenerated as in Example 10.

Table 21.

<td></td><td>Volume (L)</td><td>Concentration (g/L)</td><td>Mass (g)</td><td>Yield or {Loss} (%)</td>

<td>CombinedQ</td><td> 262</td><td> 3,72</td><td> 977</td><td> 100</td>

<td>combined recovered</td><td> 5,0</td><td> 174,0</td><td> 863,0</td><td> 88,3</td>

<td>diluted mix</td><td> 5,421</td><td> 149,6</td><td> 811,0</td><td> 83,0</td>

Example 12

High concentration formulation of rhuMAb anti-CD20 GMP

Example 11 was repeated with the following exceptions. The protein pool resulting from a preceding Q-Sepharose chromatography step was measured to reach 5.106 g of 2H7/L and had a volume of 196 L. The protein was contained in a solution of 6 mM HEPES free acid/19 mM of HEPES. HEPES sodium salt and 25 mM sodium acetate whose pH had been adjusted to 5.3 with 0.5 M acetic acid. Just before the stage, the temperature set point for this mix was set to 45^andC. To begin the step, the protein mix was transferred to the recirculating vessel through a 0.22 micron sterilizing grade filter to a level of approximately 40 L in the tank. In the vessel, the mix was agitated by the top-mounted impeller and the temperature was maintained at 40-50^andc.

During UF1 the pool was concentrated to ca. 50 g 2H7/L (ca. 20 L). The FIG. 25 shows value trends over time: recycle tank level (210) scaled from 0 to 300, retentate dP (2520) scaled from 0 to 100, feed flow rate (250) scaled from 0 to 150, and the flow rate of the filtrate (2550) scaled from 0 - 50 during the process. The pool was diafiltered with 10 diavolumes (10X) of diafiltration buffer. During diafiltration, the temperature was kept between 40 and 50⁵C. Diafiltration was performed at a constant volume that was achieved by combining the flow rate of buffer transferred to the recirculation tank with the flow rate of filtrate removed from the system. At the end of diafiltration, the pool was further concentrated in the UF2 mode to an intended set point of final concentration of 190 g 2H7/L (5.26 L) again using a constant dP control just at the end of diafiltration. this phase (see FIG. 25). This phase was carried out at an elevated temperature set point of 45⁹C +/- 5⁹C. A low pressure drop mix was then performed where the feed pump was controlled to maintain a 137.9 kPa (20 psig) pressure drop across the feed channel. A sample was taken and density measurements were made to confirm the concentration prior to recovery. The concentration of this sample was 191 g of 2H7/L. Table 22 summarizes the performance and flow results.

Table 22.

<td>process phase</td><td>Normalized Yield g/m²/s (g/ft²/hr)</td><td>Normalized Flow LMH/kPa (LMH/psig)</td>

<td>UF1</td><td> 0,20 (67)</td><td> 0,52 (3,6)</td>

<td>DF</td><td> 0,14(47)</td><td> 0,39 (2,1)</td>

<td>UF2</td><td> 0,87 (292)</td><td> 0,26(1,8)</td>

Just prior to product recovery, a 30 mL sample was taken and delivered for bioburden detection and typing. Results were negative (ie, &lt;0.13 CFU/mL). The protein pool was recovered by a series of steps as in Example 11. The volumes of the recovered pool, diluted pool, and conditioned pool were then analyzed for protein concentration. Table 23 presents the results. The membrane was regenerated as in Example 11.

Table 23.

<td></td><td>Volume (L)</td><td>Concentration (g/L)</td><td>Mass (g)</td><td>Yield or {Loss} (%)</td>

<td>Combined-Q</td><td> 196</td><td> 5,106</td><td> 1000</td><td> 100</td>

<td>combined recovered</td><td> 4,9</td><td> 187,1</td><td> 918,0</td><td> 91,8</td>

<td>diluted mix</td><td> 6,075</td><td> 150,9</td><td> 916,9</td><td> 91,7</td>

All publications, patents, and patent documents are incorporated herein in their entirety and individually by reference. The invention has been described with reference to various specific and preferred techniques and embodiments. However, it is to be understood that many variations and modifications can be made while maintaining the spirit and scope of the invention.