Computer-Aided Process Analysis and Economic Evaluation for Biosynthetic Human Insulin Production-A Case Study

Demetri Petrides,'* Elpida Sapidou? and John Calandranis' 'Department of Chemical Engineering, Chemistry, and En vironm en tal Science, New Jersey Institute of Technology, Newark, New Jersey 07702; and 21ntelligen, lnc., 2326 Morse Ave., Scotch Plains, New Jersey 07076; e-mail In tellinfo @aol. co m

Received November 22, 1994lAccepted June 15, 1995

Human insul in was the first mammalian protein pro- duced in bacteria using recombinant DNA technology. Two technologies were developed; the first based on the separate expression of precursors of chains A and B of insulin, and the second based on the expression of a precursor of proinsulin as a Trp-E fusion protein. Both technologies utilized Escherichia coli as an expression system. Later, a third technology was developed based on a strain of yeast that can secrete a precursor of insu- lin. The second E. coli process, a variation of which has been commercialized by Eli Lilly and Co., is analyzed in this article from a process design and economic evalua- tion viewpoint. The objective of this work is to elucidate the technical complexity and high cost associated with the manufacturing of biopharmaceuticals. Human insulin is a good example of a protein-based biopharmaceutical produced in large quantities (a few tons per year) that requires large scale equipment and presents a multitude of scale-up challenges. Based on the analysis, a number of conclusions are drawn regarding the cost breakdown and cost dependency on process parameters. Recom- mendations are made for cost reduction and environ- mental impact minimization. This analysis was per- formed using a software tool for computer-aided biopro- cess design. 0 1995 John Wiley & Sons, Inc. Key words: human insulin, biosynthetic process design economic evaluation insu l in

INTRODUCTION

Insulin facilitates the metabolism of carbohydrates and is essential for the supply of energy to the cells of the body.' Impaired insulin production leads to the disease diabetes mellitus, which is the third largest cause of death in indus- trialized countries after cardiovascular diseases and can- cer. ' ,' '

Human insulin is a polypeptide consisting of 51 amino acids arranged in two chains: A with 21 amino acids, and B consisting of 30 amino acids. The A and B chains are con- nected by two disulfide bonds. Human insulin has a molec- ular weight of 5734 and an isoelectric point of 5.4.

* To whom all correspondence should be addressed; Internet: Intellinfo@aol.com.

Human insulin can be produced by four different meth- ods':

Extraction from human pancreas. 0 Chemical synthesis via individual amino acids. 0 Conversion of pork insulin or "semisynthesis."

Fermentation of genetically engineered microorganisms.

Extraction from the human pancreas cannot be practiced due to the limited availability of raw material. Total syn- thesis, while technically feasible, is not economically viable due to very low yield. Production based on pork insulin, also called "semisynthesis ," transforms the porcine insulin (which differs only in one amino acid) molecule into an exact replica of the human insulin molecule by substituting the amino acid threonine for alanine in the G-30 position. This technology has been developed and implemented by Novo Nordisk A/S (Denmark). However, this option is also quite expensive because it requires the collection and pro- cessing of large amounts of porcine pancreases. In addition, its supply is limited by the availability of porcine pancreas.

At least three alternative technologies have been devel- oped for producing human insulin based on fermentation and utilizing recombinant DNA technology. '

Two-chain method. This was the first successful tech- nique of biosynthetic human insulin (BHI) production based on recombinant DNA technology, developed by Genentech, Inc. (South San Francisco, CA) and scaled-up by Eli Lilly and Co. (Indianapolis, IN)." Each insulin chain is pro- duced as a p-galactosidase fusion protein in Escherichiu coli forming inclusion bodies. The two peptide chains are recovered from the inclusion bodies, purified, and com- bined to yield human insulin. Later, the P-galactosidase operon was replaced with the tryptophan (Trp) operon, re- sulting in a substantial yield increase.

Proinsulin method (intrucellular). This method elimi- nates the need for two different fermentation and purifica- tion trains that the previous option requires. In this case, intact proinsulin is produced instead of the A or B chains separately. The proinsulin route has been commercialized by Eli Lilly and C0.3,498x'0 Figure 1 shows the key manu- facturing process steps. The E. coli cells overproduce Trp-

Biotechnology and Bioengineering, Vol. 48, Pp. 529-541 (1995) 0 1995 John Wiley & Sons, Inc. CCC 0006-3592/95/050529-13

Biomass Cell Harvesting Cell Disruption

IB recovery IB solubilization

1 t t t t 4 4

Inclusion Bodies

CNBr cleavage Trp-LEI-Met-Proinsulin

Oxidative sulfitolysis Proinsulin (unfolded)

Folding, S-S bond formation Proinsulin-SS03

Enzymatic conversion

Proinsulin (refolded)

Insulin (crude)

Purification

Purified Human Insulin

Figure 1. Human insulin from proinsulin fusion protein

LE-Met-proinsulin in the form of inclusion bodies which are recovered and solubilized. Proinsulin is released by cleaving the methionine linker using CNBr. The proinsulin chain is subjected to a folding process to allow intermolec- ular disulfide bonds to form, and the C peptide is then cleaved with enzymes to yield human insulin. A number of chromatography and membrane filtration steps are utilized to purify the product.

Proinsulin method (secreted). Novo Nordisk AJS has de- veloped a technology based on yeast cells that secrete in- sulin as a single-chain insulin precursor. Secretion simpli- fies product isolation and purification. The precursor con- tains the correct disulfide bridges and is therefore identical to those of insulin. It is converted to human insulin by transpeptidation in organic solvent in the presence of a thre- onine ester and trypsin followed by a de-esterification. An- other advantage of this technology is the ability to reuse the cells by employing a continuous bioreactor-cell separator

In the article, we propose a flowsheet for the production of BHI based on the intracellular proinsulin method. The generation of the flowsheet was based on information avail- able in the patent and technical literature combined with our engineering judgment and experience with other recombi- nant products. We use this flowsheet to draw general con- clusions on the manufacturing cost of biophmaceuticals. Furthermore, we address the issue of waste minimization, by considering in-process and on-site recycling opportuni- ties.

loop.

MATERIALS AND METHODS

Process Description

The plant analyzed in this example has a capacity of around 1500 kg of purified BHI per year. This represents 10% to 15% of the current world demand and corresponds to a

relatively large plant for producing polypeptide-based biop- harmaceuticals. 6.

The entire flowsheet for the production of BHI is shown in Figure 2. Several holding tanks and other auxiliary pieces of equipment were omitted for simplicity. However, their approximate costs were taken into account in the economic evaluation.

Media Preparation and Fermentation Fermentation media are prepared in a stainless-steel tank (V-101) and sterilized in a continuous heat sterilizer (E- 101). The filling of the production fermentor with sterilized media takes 5 h. The axial compressor, G-101, and the absolute filter, AF-101 provide sterile air to the fermentors at an average rate of 0.4 vvm. Transformed E . coli cells are used to produce the Trp-LE-Met-proinsulin precursor of insulin. The fermentation time in the production fermentor is about 18 h, while the turnaround time is 12 h . The fer- mentation temperature is 37C. A two-step seed fermentor train (not shown in Fig. 2) is used to inoculate the produc- tion fermentor. The final concentration of E . coli in the pro- duction fermentor is about 37.5 g/L dry cell weight ( ~ c w ) . ~

The Trp operon is turned on when the E . coli fermenta- tion runs out of tryptophan.6 The chimeric protein, Trp- LE-Met-proinsulin, accumulates intracellularly as insolu- ble aggregates (inclusion bodies) and this decreases the rate at which the protein is degraded by proteolytic enzymes. In the base case, it was assumed that the inclusion bodies (IBs) constitute 20% of total dry cell mass. After the end of fermentation, the recombinant microorganisms are killed using a thermal inactivation

Primary Recovery Stages The broth is transferred into a surge tank (V-102) which isolates the upstream from the downstream section of the plant. A disk-stack centrifuge (CF-101) is used for cell har- vesting to reduce the volume of the broth by a factor of 1.85 and remove the extracellular impurities and other undesir- able components. The cell sludge temperature is maintained at around 10C. The cell sludge is diluted by a factor of 2.5 with a buffer solution (buffer composition in percent w/w: DI H,O 96.4, EDTA 0.7, and Tris-base 2.9) using a mixer tank (M-102). The buffer facilitates the separation of the cell debris particles from inclusion bodies. A high pressure homogenizer (HG- 10 I ) is used to break the cells and release the inclusion bodies. The broth undergoes three passes un- der a pressure drop of 800 bar. The exit temperature is maintained at around 10C.

A second disk-stack centrifuge (CF-102) is used for in- clusion body recovery. The inclusion bodies are recovered in the heavy phase while most of the cell debris particles remain in the light phase. The inclusion body sludge, which contains approximately 20% w/w solids, is washed with DI water containing 0.66% w/w Triton-X- 100 detergent (the volume of solution is three times the volume of inclusion body sludge) and recentrifuged (CF- 103). The detergent

530 BIOTECHNOLOGY AND BIOENGINEERING, VOL. 48, NO. 5, DECEMBER 5, 1995

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solution facilitates purification (dissociation of debris and soluble proteins from inclusion bodies). The exit tempera- ture is maintained at 10C. The overall yield of the primary recovery stages is 94%.

Inclusion Body Solubilization

The inclusion body suspension is transferred to a well- mixed reactor (R-102) and is mixed with urea and 2-mer- captoethanol to final concentrations of 300 glL ( 5 M ) and 40 g/L, respectively. Urea is a chaotropic agent that dissolves the denatured protein in the inclusion bodies and 2-mercap- toethanol is a reductant that reduces disulfide bonds. A reaction time of 8 h is required to reach a solubilization yield of 95%. Inclusion bodies contain 80% w/w Trp-LE'- Met-proinsulin with the rest being other contaminant pro- teins. After the end of the reaction, urea and 2-mercaptoet- hanol are replaced with DI water using a diafiltration unit (DF-101). This operation is performed in 4 h. All remaining fine particles (biomass, debris, inclusion bodies) are re- moved using a polishing dead-end filter (DE-101). This polishing filter protects the chromatographic units that are used further downstream.

CNBr Cleavage

The chimeric protein is cleaved with CNBr (cyanogen bro- mide) into the signal sequence Trp-LE'-Met, which con- tains 121 amino acids, and the denatured proinsulin (82 amino acids) in a well-mixed reactor (R-103). The reaction is carried out in a 70% v/v formic acid solution containing 30-fold molar excess CNBr (stoichiometrically, 1 mol of CNBr is required per mole of Trp-LE'-Met-proinsulin). The reaction takes 6 h at 20C and reaches a yield of 95%.

The mass of the released proinsulin is approximately 3 1 % of the mass of Trp-LE'-Met-proinsulin. A small amount of cyanide gas is formed as a byproduct of the cleavage reac-

tion. Detailed information on CNBr cleavage is available ir the patent literature (U.S. Patent No. 4,451,396, 1984) The formic acid, the unreacted CNBr, and the small amoun of generated cyanide gas are removed by applying vacuum and raising the temperature to around 35C (the boiling point of CNBr). This operation is carried out in a rotaq vacuum evaporator (M-104) and takes 1 h. Because cyanide gas is toxic, the vapor stream must be handled appropri- ately. Oxidation with hypochloride and wet air oxidatior are cited in the literature as effective treatment methods foi cyanogen gas.4 The cost of cyanogen gas treatment wa: considered as part of the cost for treatment/disposal ol CNBr and formic acid. Values of $2/kg and $l/kg were assumed for the unit costs of treatment/disposal of CNBi and formic acid, respectively.

Sul fit0 I ysis

Sulfitolysis of the denatured proinsulin takes place in a well-mixed reactor (R- 104) under alkaline conditions (pH 9 to 11). This operation unfolds the proinsulin molecules (by breaking all disulfide bonds) and adds SO, moieties to all sulfur residues on the cysteines. The product of interest is human proinsulin (S-S03-), (protein-S-sulfonate). The sulfitolysis mixture contains 50% w/w guanidine - HCI (6 M ) , 0.35% w/w ammonium bicarbonate (NH,HCO,), 3% w/w Na,SSO,, and 1.56% w/w Na20,S, (U.S. Patent No. 4,923,967, 1990). A reaction time of 6 h is required to reach a yield of 95%. The presence of the denaturing re- agent (guanidine . HCl) prevents refolding and cross- folding of the same protein molecule onto itself or two separate protein molecules onto each other. Urea may also be used as a denaturing reagent. After the completion of the sulfitolysis reaction, the sulfitolysis solution is exchanged with distilled water in a diafilter (DF-102) to a final guanidine . HCl concentration of 20% w/w.

Water Glucose Urea Guanidine.HC1 HCOOH NaCl NaOH NH 3 Acetonitrile Acetic Acid MrEtOH EDTA Triton-X-100 CNBr Na2S03 Na2 06 54

589,650 3,810 18,960 3,335 7,390 3,865 1,515

3 2 0 6,580 2,810

845 818 37 20 194 97

Process StepS

Media Preparation Fermentation Cell Harvesting Cell Disruption IB Recovery IB Solubilization CNBr Cleavage Sulfitolysis Refolding Enzymatic Conversion Final Purification

Z n = Cout = 640,250 kg/batch Figure 3. Overall material balances per batch (kgil80 h).

Insulin

c02

Total Waste Organics H20

15

2,300

637,930 46,480 591,450

532 BIOTECHNOLOGY AND BIOENGINEERING, VOL. 48, NO. 5, DECEMBER 5, 1995

The human proinsulin (S-S03-), is chromatographically purified using an ion-exchange column (C-101). A cation exchange resin is used (SP sepharose big beads from Phar- macia) operating at pH 4.0. The eluant solution contains: H,O 69.5% wlw, urea 29% wlw, and NaCl 1.5% wlw. Urea, a denaturing agent, is used to prevent incorrect re- folding and cross-folding of proinsulin (S-S03-),. The fol- lowing operating assumptions were made: (1) total resin binding capacity of 20 mglmL; (2) the eluant volume is equal to 2 column volumes (CVs); (3) the total volume of the solutions for column wash, equilibration, regeneration, and storage is 8 CVs; and (4) the protein of interest is recovered in 1 CV of eluant buffer with a recovery yield of 90%.

Refolding

Folding of proinsulin( S-S03-), and disulfide bond forma- tion takes place in a well-mixed reactor (R-105). This pro- cess step involves treatment with mercaptoethanol (a reduc- tant) which facilitates the disulfide interchange reaction to occur. Sufficient dilution (to a concentration of proinsu- lin(S-S03-), of less than I glL) is required to prevent cross- folding of proinsulin molecules (U .S . Patent No. 4,923,967, 1990). The product of interest is human proin- sulin. The reaction is carried out at 8C for 12 h and reaches a yield of 85%. After the completion of the refolding, the refolding reagents are replaced with water for injection (WFI) and the protein solution is concentrated using a di-

Table 1. Waste treatmentidisposal costs (1994 prices).

Waste component

Biomass H2O Glucose NH, Incl. body Trp. proinsulin Cont. protein Proinsulin SSO, Proinsulin Cell debris Triton-X 100 CNBr

NaZO6S, Guanidine HCI HCOOH NaCl Acetonitnle NaOH Urea MrEtOH Salts NH,HCO, EDTA Total

Na,SO,

Unit cost ( $ i k g )

Annual amount (kg)

~

1.000e + 00 2.000e - 04 7.000e - 01 2.000e - 01 1.000e + 00 1.000e + 00 1.000e + 00 1.000e + 00 1.000e + 00 1.000e + 00 1.000e + 00 2.000e + 00 1.000e + 00 1.000e + 00 1.000e + 00 1.000e + 00 5.000e - 02 5.000e - 01 5.000e - 02 1.000e + 00 1.000e + 00 1.000e + 00 1.000e + 00 1.000e + 00

2982 59,145,632

15,244 6148 2512 1154

63,387 1456 890

7 1,682 3695 1988

19,192 9496

333,455 739,167 386,506 657,844 151,571

1,895,738 84,535

5175 2298

81,826

cost (Uyr.)

3000 12,000 11,000

1000 3000 1000

63,000 1000 1000

72,000 4000 4000

19,000 9000

333,000 739,000

19,000 3 29,000

8000 1,896,000

85,000 5000 2000

82,000 3,702,000

afiltration unit (DF- 103) which has product recovery yield of 95%.

The human proinsulin is chromatographically purified in a hydrophobic interaction chromatography column (C- 102). The primary impurities removed at this stage include the unfolded proinsulin (S-SO3-),, incorrectly folded proinsu- lin, and general E . coli proteins. The following operating assumptions were made: (1) total resin binding capacity is 20 mglmL; (2) the eluant volume is equal to 6 CVs; (3) the total volume of the solutions for column wash, equilibra- tion, regeneration, and storage is 10 CVs; and (4) the pro- tein of interest is recovered in 1 CV of eluant buffer with a recovery yield of 90%.

Enzymatic Conversion

The removal of the C peptide from human proinsulin is carried out enzymatically (using trypsin and carboxypepti- dase B) in a well-mixed reactor (R-106). The cleavage of

v-101

E-101

G-101

M-101

AF-101

R-1 01

AF-102

v-102

CF-101

M-102

HG-101

CF-102

M-103 CF-103

R-102 DF-101

DE-101

R-103

M-104

R-104

DF-102

c-101

R-105

DF-103

c-102

R-106 DF-104

C-103 DF-105

C-104

DF-106

C-105

UF-101

M-105

CF-104

M-1 OE 1

1 1

0.0 24.0 48.0 72.0 96.0 120.0 144.0 168.0

Time [hours]

Figure 4. Process scheduling Gantt chart.

PETRIDES, SAPIDOU, AND CALANDRANIS: BIOSYNTHETIC HUMAN INSULIN PRODUCTION 533

Table 11. Major equipment specification and FOB cost (1994 prices).

Quantity/ stand-by

Cost Unit cost ($1 ($1 Description

1/0 v-101

1/0 E-101

110 (3-101

110 AF-101

1/0 R-101

110 R-101

110 R-101

1/0 AF-102

110 v-102

110 CF-I01

210 HG-101

2/0 CF-102

110 CF-103

110 R-102

310 DF-101

110 DE-101

110 R-103

110 R-104

210 DF-102

810 C-101

110 R-105

1310 DF-103

810 C-102

110 R-106

110 DF-104

810 C-103

1/0 DF-105

410 C-104

110 DF-106

10/0 c-105

110 UF-101

1/0 M-105

Blending tank vol. = 30.92 m3, SS304, 6.57 kW Continuous sterilizer CdpaCity = 5.20 m3/h Centrifugal compressor DPress = 5.0 bar, SS316, 50.00 kW Cartridge air filter Air inlet, SS316 Housing Agitated fermentor vol. = 35.04 m3, SS316, 78.85 kW Agitated seed fermentor vol. = 3.50 m3, SS316, 7.88 kW Agitated seed fermentor vol. = 0.35 m3, SS316, 0.79 kW Cartridge air filter exhaust gas, SS316 housing Vertical storage tank vol. = 32.09 m3, SS304, 1.01 bar Disk-stack centrifuge X = 147,718 m, 25.58 kW High pressure homogenizer cap. = 10,102.1 Uh, 224.49 kW Disk-stack centrifuge 8 = 304,878 m2, 34.18 kW Disk-stack centrifuge 8 = 46,727 m2, 16.14 kW Agitated reactor vol. = 7.13 m3, SS316, 3.03 kW Membrane diafilter area = 63.4 m, 12.68 kW Dead-end filter area = 5.0 rn, pore size = 0.45 Frn Agitated reactor vol. = 12.26 m3, SS316, 5.21 kW Agitated reactor vol. = 8.75 m3, SS316, 3.72 kW Membrane diafilter area = 69.8 ni, 13.95 kW Ion exchange uni t L = 0.25 m, D = 0.97 m Agitated reactor vol. = 72.90 m, SS316, 30.98 kW Membrane diafilter area = 74.5 m, 14.90 kW HIC column unit L = 0.25 m, D = 0.96 m Agitated reactor vol. = 3.46 m, SS316, 1.47 kW Membrane diafilter arca = 32.7 m, 6.53 kW Ion exchange unit L = 0.25 m, D = 0.98 m Membrane diafilter area = 25.5 m, 5.10 kW RP-HPLC unit L = 0.25 m, D = 0.97 m Membrane diafilter area = 37.1 m, 7.43 kW Gel filtration unit L = 0.50 m, D = 0.97 m Membrane ultrafiltcr area = 44.7 m, 8.97 kW Crystallizer

66,000

249,000

152,000

11,000

535,000

186,000

77,000

12,000

46,000

277,000

68,000

282,000

183,000

83,000

76,000

7000

109,000

92,000

79,000

268,000

327,000

81,000

267,000

57,000

54,000

270,000

46,000

268,000

58,000

270,000

65,000

33,000

66,000

249,000

152,000

11,000

535,000

186,000

77,000

12,000

46,000

277,000

136,000

564,000

183,000

83,000

228,000

7000

109,000

92,000

158,000

2,144,000

327,000

1,053,000

2,136,000

57,000

54,000

2,160,000

46,000

1,072,000

58,000

2,700,000

65,000

33,000

534 BIOTECHNOLOGY AND BIOENGINEERING, VOL. 48, NO. 5, DECEMBER 5, 1995

Table 11. Continued

Quantity/ stand-by Description

Unit cost cost ($) ($1

1/0 v-101 vol. = 0.8 m, SS316, 0.5 kW

110 CF-104 Basket centrifuge 27,000 27,000 area = 1 m, SS316, 0.00 kW Cost of unlisted equipment 3,776,000 20.0% of total

Total cquipment purchase cost 18,879,000

proinsulin is initiated by the addition of trypsin, which cleaves at the carboxyterminal of internal lysine and argi- nine residues, and carboxypeptidase B, which removes ter- minal amino acids.6 Carboxypeptidase is added to a final concentration of 4 mg/L and trypsin to a final concentration of 1 mg/L. The reaction takes place at 10C for 12 h reach- ing a conversion yield of 95%.

Final Purification Steps

A purification process based on multimodal chromatogra- phy, which exploits differences in molecular charge, size, and hydrophobicity, is typically used to isolate the human insulin. . The assumed values of performance and oper- ating variables were based on information available in the literature as we11 as our experience from similar processes.2 A detailed description of the various steps follows.

The enzymatic conversion solution is exchanged with WFI and concentrated by a factor of 1 .5 in a diafilter (DF- 104). An ion- exchange column ((2-103) is used to purify the insulin solution. The primary impurities removed at this stage include the C-peptide, the enzymes trypsin and carboxypeptidase, and the remaining amount of unfolded proinsulin (S-S03-),. The following operating assumptions were made: (1) total resin binding capacity is 20 mgimL; (2) the eluant volume is equal to 8 CVs and it is an 1 I .5% wlw solution of NaCl in WFI; (3) the total volume of the solutions for column wash, equilibration, regeneration, and storage is 8 CVs; and (4) the protein of interest is recovered in I CV of eluant buffer with a recovery yield of 95%.

The ion-exchange eluant solution is exchanged with WFI in a diafilter (DF-105) and is concentrated by a factor of 2.0. A recovery yield of 98% was assumed for this step.

The purification of the insulin solution proceeds with a reversed-phase high-pressure liquid-chromatography (RP- HPLC) step (C-104). Detailed information on the use of RP-HPLC for insulin purification is available in the litera- ture. The primary impurities removed at this stage include structurally similar insulinlike components. Analytical stud- ies with a variety of reversed-phase systems, have shown that an acidic mobile phase can provide excellent resolu- tion. Minor modifications in the insulin molecule resulting

in monodesamido (A-21) formation, or derivatization of amines via carbamoylation or formylation, result in insulin derivatives which have significantly increased retention. Derivatives of this nature are typical of the kind of insulin- like components that are found in the charge stream going into the reversed-phase purification.

The use of an acidic mobile phase results in elution of all the derivatives after the insulin peak, whereas the use of mildly alkaline pH results in derivatives eluted on either side of the parent insulin peak. An ideal pH for insulin purification is in the region of 3.0 to 4.0, because this pH range is far enough below the isoelectric pH of 5.4 to pro- vide for good insulin solubility. An eluant buffer with an acetic acid concentration of 0.25 M meets all the operational criteria. It is compatible with the chromatography and pro- vides good insulin solubility.

A 90% insulin yield was assumed in the RP-HPLC step with the following operating conditions: (1) total resin bind- ing capacity is 15 mg/mL; (2) the column height is 25 cm; (3) the eluant volume is equal to 6 CVs and its composition is 25% w/w acetonitrile, 1.5% wlw acetic acid, and 73.5% w/w WFI; (4) the total volume of the solutions for column wash, equilibration, regeneration, and storage is 6 CVs; and ( 5 ) the protein of interest is recovered in 1 CV of eluant buffer with a recovery yield of 90%.

The RP-HPLC buffer is exchanged with WFI and con- centrated by a factor of 2.0 in a diafilter (DF-106) that has a product recovery yield of 98%. Purification is completed by a gel filtration unit (C-105). The following operating assumptions were made: (1) the sample volume is equal to 5% of the column volume; (2) the eluant volume is equal to 4 CVs; (3) the total volume of the solutions for column wash, depyrogenation, stripping, and storage is 6 CVs; and (4) the protein of interest is recovered in 0.4 CVs of eluant buffer with a recovery yield of 90%. The mobile phase is a solution of acetic acid.

The ultrafilter (UF-101) is used to concentrate the puri- fied insulin solution. Concentration is followed by zinc crystallization by adding small amounts of zinc chloride.

The crystals are recovered with a basket centrifuge and freeze dried. The purity of the crystallized end product is between 99.5% and 99.9%, measured by analytical high pressure liquid chromatography (HPLC).

PETRIDES, SAPIDOU, AND CALANDRANIS: BIOSYNTHETIC HUMAN INSULIN PRODUCTION 535

Table 111. Fixed capital estimate summary (1994 prices)

A. Total plant direct cost (TPDC) (physical I . Equipment purchase cost allocated

to this product Total

2. Installation 3. Process piping 4. Instrumentation 5. Insulation 6. Electrical 7. Buildings 8. Yard improvement 9. Auxiliary facilities

cost)

(PC)

(0.51 x PC) (0.35 x PC) (0.50 x PC) (0.03 X PC) (0.10 x PC) (0.85 x PC) (0.15 X PC) (0.60 X PC)

$ 18,879,000 18,879,000 9,540,000 6,608,000 9,439,000

566,000 1,888,000

16,047,000 2,832,000

11,327,000

B. Total plant indirect cost (TPIC) 10. Engineering (0.25 X TPDC) 11. Construction (0.35 X TPDC)

TPDC = 77,126,000

19,281,000 26,994,000

TPIC = 46,27.5,000

C. Total plant cost (TPDC + TPIC) TPC = 123,401,000 12. Contractors fee (0.05 X TPC) 6,170,000 13. Contingency (0.10 x TPC) 12,340,000

D. Direct fixed capital (DFC) TPC + 12 + 13 X(12 + 13) = 18,510,000

141,911,008

The Process Simulation Environment

The process analysis and economic evaluation of the large scale human insulin production was done using the Apple Macintosh version of BioPro Designer@ from Intelligen, Inc. (Scotch Plains, NJ). BioPro Designer is a process sim- ulator that facilitates design, analysis, and evaluation of integrated biochemical, pharmaceutical, and food pro- cesses. Detailed information on the features, capabilities, and limitations of BioPro Designer can be found in the literature. 1 3 9 1 4

RESULTS AND DISCUSSION

Material Balances and Environmental Impact Assessment

Figure 3 shows a summary of the overall material balances per batch. The overall product recovery yield is around 30%. The materials for equipment cleaning (primarily so- lutions of water) and maintenance were not considered in the material balances. Huge amounts of waste are generated for the production of 15 kg of purified insulin per batch. The total waste-to-product ratio is 45,520: 1, while the or- ganic waste-to-product ratio is 3 100: 1.

In the base case, it was assumed that this waste is treated and disposed. However, opportunities may exist for recy- cling some chemicals (for in-process use) and recovering others (for off-site use). For instance, formic acid (HCOOH), acetonitrile, and urea are good candidates for recycling or recovery. Formic acid is used in relatively large

quantities (7.4 tonslbatch) in the CNBr cleavage step (R- 103) and it is removed using a rotary vacuum evaporator (M-104) along with small quantities of CNBr, H,O, and urea. The recovered formic acid can be readily purified by distillation and recycled in the process. A total of 1.8 tons/ batch of urea are used in the solubilization of inclusion bodies (R-102) and a larger amounts of 17 tonsibatch is

Table IV. Raw materials (1994 prices).

Unit cost Annual amount cost Component Wkg) (kg) Wyr.)

H2O

NH, Glucose

Acetic acid Triton-X 100 CNBr Na,SO, Na,O,S, Guanadine-HC1 HCOOH NaCl Acetonitrile NaOH Urea MrEtOH Enzymes Salts NH,HCO, EDTA

Total

5.000e - 02 58,965,136.00 6.000e ~ 01 381,108.25 7.000e - 01 31,758.96 2.500e + 00 280,837.91 1.500e + 00 3695.28 2.000e + 00 1988.16 4.000e ~ 01 19,377.60 6.000e - 01 9681.36 1.000e + 00 333,455.03 1.600e + 00 739,166.69 1.230e + 00 386,505.84 3.000e + 00 657,844.00 3.500e + 00 151.571.02 1.520e + 00 1,895,738.12 8.000e - 01 84,534.72 1.000e + 05 I .92 1.000e + 00 29,566.08 1.000e + 00 2297.52 1.850e + 01 8 1,825.84

2,948,000 229,000 22,000

702,000 6000 4000 8000 6000

333,000 I , 183,000

475,000 1,974,000

530,000 2,882,000

68,000 192,000 30,000

2000 I ,5 14,000

13,107,000

536 BIOTECHNOLOGY AND BIOENGINEERING, VOL. 48, NO. 5, DECEMBER 5, 1995

Table V. Various consumables (1 994 prices).

Unit cost Annual amount Cost Equipment ($imz) (mZ) W r . )

Membrane or filter cloth DF-101 DF-I01 DF-102

DF- 104 DF- I05 DF- I06 UF-I01

DF- 103

Subtotal

200.00 180.00 200.00 200.00 200.00 200.00 200.00 200.00

57.05 125.34 41.86

774.55 13.07 10. I9 14.85 17.87

11,000 23,000

8000 155,000

3000 2000 3000 4000

208,000

Unit cost Annual amount cost Equipment ( $ W (L) ($iyr.)

Chromatography resins C-101 300.00 1 1,735.52 3,52 I ,000 c-102 300.00 5779.58 1,734,000 C-103 300.00 30 12.95 904,000 C- 104 300.00 293 1.46 879,000 C- 105 200.00 4950.13 990,000

Subtotal 8,028,000

Total 8,236,000

used in the first chromatography column (C-101) to purify proinsulin (S-SO3), before its refolding. Approximately 90% of the urea appears in just two waste streams (S 128 and S143). It is unlikely that urea can be purified economically

Table VI. Annual operating cost (1994 prices)

for in-process recycling. However, it can be shipped off-site for further processing and utilization as a nitrogen fertilizer.

Approximately 6.6 tondbatch of acetonitrile is used in the reversed-phase HPLC column ((2-104) and most of it ends up in the waste stream of the column (S 168) along with 7.9 tons of H,O, 2.8 tons of acetic acid, and small amounts of NaCl and other impurities. It is unlikely that acetonitrile can be recovered economically to meet the high purity spec- ifications for a process step so close to the end of the pu- rification train. However, there may be a market for off-site use. Table I shows the assumed values of unit cost of treat- ment/disposal for the various waste chemicals.'

Process Scheduling

Figure 4 shows the results of process scheduling in the form of a Gantt chart generated by BioPro Designer. The total plant batch time is around 180 h . The plant operates around the clock for 300 days per year producing 1500 kg of insulin in 100 batches. In other words, the effective plant batch time (the time between initiation of consecutive batches) is 3 days. This is possible because the cycle time of the lim- iting step (the fermentor in this case) is only 30 h, or 1.25 days.

Process Economics

The tables are excerpts from the economic evaluation re- ports generated by BioPro Designer. Table I1 shows a list of

1 . DFC-dependent items (DFC = $141,911,008) Depreciation Maintenance material (summed over all units) Insurance (0.01 x DFC) Local taxes (0.02 X DFC) Factory expense (0.05 X DFC)

$ 13,482,000 2,246,000 1,419,000 2,838,000 7,096,000

2. Labor-dependent items a. Operating labor (40,553 h X 20.0 $ih) b. Maintenence labor (summed over all units) c. Fringe benefits [0.40 X (a + b)] d. Supervision [0.20 X (a + b)] e. Operating supplies (0.10 X a) f . Laboratory (0.30 X a)

3. Administration and overhead expense

4. Raw materials 5. Other consumables 6 . Utilities 7. Waste treatment 8. Running R&D 9. Running royalties 10. Sales cost

[0.6 X (a + b + c)]

Total annual operating cost Including depreciation Excluding depreciation

27,08 1,000

81 1,000 749,000 624,000 312,000 8 1,000

243,000

2,820,000

1,123,OO 13,107,000 8,236,000

127,000 3,702,000

0 0 0

56,196,000 42,7 14,000

PETRIDES, SAPIDOU, AND CALANDRANIS: BIOSYNTHETIC HUMAN INSULIN PRODUCTION 537

Waste Treatment/Disposal 6.6% titilities

Administration Laltor-L)e endent 5 . 0 2 2.0%

Figure 5. Operating cost breakdown ($56.2 millioniyear)

equipment with some descriptive information and the pur- chase cost for each piece. For the plant of the base case (1500 kg/year of BHI), the total equipment purchase cost is around $18.9 million. It is assumed that the cost of unlisted equipment (various tanks, pumps, and other auxiliary units) is 20% of the total equipment cost. The chromatography columns are the most expensive pieces of equipment. Table 111 shows a summary of the fixed capital investment which is around $14 1 million.

Table IV shows a breakdown of the raw materials cost. The cost of water (with an assumed unit cost of $.05/kg) is the most significant item, followed by the cost of urea., acetonitrile, and formic acid. Highly purified water (dis- tilled in the upstream section and primary recovery stages and WFI in the high resolution purification stages) is used throughout the processes. Urea is used for the solubilization of inclusion bodies (R-102) and in the eluant and washing solutions of the first chromatography column (C- 101) prior

Upstream Section 3J% Primary Recovery

Final Purification 30.7 %

to refolding of proinsulin (S-S03-),. Acetonitrile is used in the eluant solution of the reversed-phase HPLC column (C- 104), whereas formic acid is used in the CNBr cleavage step (R-103). Table V shows a breakdown of the various con- sumables. Substantial cost is associated with the chroma- tography column resins which must be replaced on a regular basis. Table I shows a breakdown of the waste treatment/ disposal costs. Substantial costs are associated with urea, formic acid, guanidine hydrochloride, and acetonitrile.*

Table VI shows a summary of the annual operating cost, and Figure 5 shows a breakdown of the same cost. Clearly, the direct-fixed-capital (DFC)-dependent cost is the most important item accounting for almost 50% of the total op- erating cost. The cost of raw materials lies in the second position accounting for 23.3% of the total cost followed by the cost of various consumables (14.6% of the total cost). Depreciation was calculated over a 10-year period assuming a 5% salvage value for the entire plant.

Figure 6a shows a breakdown of the operating cost on a per-process section and equipment category basis. Almost 60% of the total operating cost is associated with the chem- ical modification and intermediate purification steps (in- cluding steps from solubilization to enzymatic conversion). The final purification steps account for 30% of the total cost. Overall, the cost of product recovery and purification is almost 97% of the total cost. Figure 6b shows the cost breakdown per equipment type. The cost of chromatogra- phy is almost 70% of the total operating cost. This is fre- quently the case for protein-based biopharmaceuticals that are required in very high purity.

Even though the cost of waste treatment and disposal is relatively low (6.6%) compared to the total, it is expected to rise more rapidly (due to stricter future environmental reg- ulations) than the other costs, unless pollution prevention and waste minimization options are considered. Because it

Other 11.6%

Filters 8.2%

Reactors 10%

Chemical Modifications and Intermediate Purification Steps

58.6%

Figure 6. Operating cost per process section and equipment category

Chromatography 70.2 %

538 BIOTECHNOLOGY A N D BIOENGINEERING, VOL. 48, NO. 5, DECEMBER 5, 1995

Table VII. Profitability analysis (I994 prices)

A. Direct fixed capital B . Working capital

D. Total investment (A + B + C) E. Production rate (kgiyr.)

F. Production cost ($/kg)

G. Selling price ($/kg)

H. Revenue ($/yr.)

I. Annual operating cost J . Gross profit (H - I) K. Taxes (40%) L. Net Profit (J - K + depreciation)

c. Stdfl-Up COSt

Insulin

Insulin

Insulin

Insulin

Gross margin Return on investment Payback time (yr.)

$ 141,911,008 873,000

7,096,000 149,880,000

1500

37.464

110,000

165,000,000 56,196,000

108,804,000 43,5 22,000 78,764,000

0.66

1.90 53%

and Table VIII shows the cash flow analysis. For a plant of this capacity (1500 kgiyear), the manufacturing cost of BHI is around $37.5/g. To achieve an internal rate of return (IRR) after taxes of 30% (an appropriate level of IRR for a high value and risk product such as BHI), a selling price of $1 lO/g is required. The unit cost of $37.5/g is four- to five-fold higher compared with costs of other recombinant E . coli proteins that do not require refolding and/or are not manufactured for pharmaceutical use. l 3

An analysis of Novo Nordisks process (secreted proin- sulin) would most likely result in a more balanced distribu- tion of costs between the upstream and downstream sections because fewer recovery and purification steps are required. Furthermore, in that case, the rate of proinsulin secretion by the yeast cells is the critical parameter that determines pro- cess economics.

Sensitivity Analysis

is costly and time consuming to make major process mod- ifications after a pharmaceutical manufacturing process has received approval by the Food and Drug Administration (FDA), it is very important to consider environmental and safety issues at the early stages of process development.

Table VII shows the results of the profitability analysis

After a model for the entire process is developed on the computer, tools like BioPro Designer can be used to ask and readily answer what i f questions and carry out a sensi- tivity analysis with respect to key design variables. In this example, the effect of three parameters on the production cost and the profitability of the project was examined.

Table VIII. Cash flow analysis

Capital Sales Operating Gross Loan Taxable Taxes Net Net Year outlay revenues cost profit payment Depreciationh income 40 .08 Profit Cash-flow

I - 42,573 2 - 56,764 3 - 43,446 4 0 5 0 6 0 7 0 8 0 9 0

10 0 I 1 0 12 0 13 0 14 0 15 7969

0 0

49,500 99,000

132,000 132,000

1 48,500 148,500 1 48,500 148,500 148,500 148,500 148,500 148,500

I 48,500

0 0

23,120 31,517 37.116 37,116 39,915 39,915 39,915 39,915 39,915 39,915 39,915 39,915 39,915

0 0

26,380 67,483 94,884 94,884

108,585 108,585 108,585 108,585 108,585 108,585 108,585 108,585 108,585

0 0

10,277 10,277 10,277 10,277 10,277 10,277 10,046 10,046 10,046 10,046

0 0 0

0 0

13,482 13,482 13,482 13,482 13,482 13,482 13,482 13,482 13,482 13,482

0 0 0

0 0

2622 43,724 71,126 71,126 84,827 84,827 85,057 85,057 85,057 85,057

108,585 108,585 108,585

0 0

1049 17,490 28,450 28,450 33,93 I 33,931 34,023 34,023 34,023 34,023 43,434 43,434 43,434

0 0

15,055 39,716 56,157 56,157 64,371 64,377 64,516 64,516 64,516 64,516 65,151 65,151 65,151

- 42,573 - 56,764 -28,392

39,7 16 56,157 56,157 64,377 64,377 64,516 64,516 64,516 64,516 65,151 65,151 73,120

IRR before taxes = 0.416 IRR after taxes = 0.305

Interest = 0.080 0.100 0.120 Net present value = 269088 218,628 176,927

Direct fixed Working Up front Up front capital capital R&D royalties

Amount 141,911 873 0 Equity (%) 60.0 0.0 0.0 Debt (9%) 40.0 100.0 100.0 Interest (%) 12.0 15.0 12.0 Loan Time (yr.) 10.0 6.0 6.0

0 50.0 50.0 15.0 6.0

Amounts are given in thousands of U.S. dollars. bDepreciation method = straight line.

PETRIDES, SAPIDOU, AND CALANDRANIS: BIOSYNTHETIC HUMAN INSULIN PRODUCTION 539

80 I 1

35 - 70 -

h . M 5 60- 8

2

k

&

50-

.- * V

4 4 0 - 0

30 -

z o ! . , . , . , . , . , . , . , . l 0 10 20 30 40 50 60 70 80

Inclusion bodies in biomass (%)

Figure 7. Production cost as a function of inclusion bodies in biomass.

Product Cost As a Function of Inclusion Bodies in Biomass

In the base case, it was assumed that inclusion bodies con- stitute 20% of dry cell biomass and that the fusion Trp-LE'- Met-proinsulin protein constitutes 80% of inclusion body mass. In the literature, values of up to 70% have been reported for composition of biomass in inclusion bodies. Figure 7 shows the production cost of human insulin as a function of the content of biomass in inclusion bodies. The cost is quite sensitive for values below 20% (a content of 5% would almost double the cost of human insulin) but not very sensitive for values greater than 20% (an increase to 50% would only reduce the cost by 15%). This can be explained by the fact that most of the production cost is associated with the recovery and purification of the product. During research and development, information of this type can be used to prioritize future work. Under limited re-

70 I

s s - . E c1 50-

u 4 s - C

40- V a 'F) 3 5 -

a 30- 25 -

2

base case

100 batcheslyear

20 0 2000 4000 6000 8000 10000 12000

Production Rate (kglyear)

Figure 8. capacity.

Production cost as a function of process scheduling and plant

Figure 9. and development expenditures.

Internal rate of return (IRR) a5 a function of up-front research

sources, for instance, it would be wiser to focus research and development efforts on the purification section of the process and not on fermentation.

Product Cost As a Function of Plant Throughput Process Scheduling The plant batch time of the base case is around 180 h and the product is manufactured in 100 batches per year. How- ever, the cycle time of the limiting process step, which is the fermentor (R-101) in the base case, is only 30 h. Theo- retically, this means that a new batch may start every 30 h resulting in up to 240 batches per year (assuming plant operation of 300 days per year). A higher number of batches per year usually leads to higher labor cost (more operators are needed around the clock) and to lower capital cost (due to better utilization of equipment).

The results of the analysis are shown in Figure 8. The same figure also shows the product cost as a function of production rate (also known as the effect of the economy of scale). As can be seen, a cost reduction of approximately 25% can be realized if the number of batches can be in- creased from 100 to 200. On the other hand, if the plant operates at 50 batchedyear, the product cost increases by almost 57% (from $37/g to $58/g). Figure 8 also shows that the production cost is quite sensitive to production rate for rates of less than 1000 kgiyear, but almost insensitive for rates greater than 3000 kg/year. This can be explained by the fact that, for production rates greater than 3000 kg/year, all expensive equipment (primarily chromatography col- umns, centrifuges, and membrane filters) are required in multiple pieces operating in parallel, thus eliminating the opportunity for savings.

internal Rate of Return As a Function of Up-Front R&D There is a high cost and, consequently, a high risk associ- ated with the development and commercialization of phar-

540 BIOTECHNOLOGY AND BIOENGINEERING, VOL. 48, NO. 5, DECEMBER 5, 1995

niaceuticals. Values of up to $350 million (Business Week, October 17, 1994, p. 204) have been reported for the av- erage cost of a successful drug substance. If this cost is amortized over the life of a product, then it may have a major impact on the profitability of the product. Figure 9 shows the results of such an analysis. In the base case, a selling price of $1 1O/g of insulin was assumed that led to an internal rate of return (IRR) of 30% (which is high for the standards of commodity chemicals, but average or low for high risk products). Figure 9 shows the IRR as a function of up-front research and development expenditures that are amortized over the project life of the product. Expenditures of just $75 million would decrease the profitability of an insulin plant (operating at 1500 kg/year) to levels of com- modity chemicals (around 15%).

CONCLUSIONS

In this article, we have synthesized and analyzed a flow- sheet for the production of BHI based on bits of information scattered in the literature and put together using engineering judgment and previous design experience. The work was facilitated using a computer-aided bioprocess design tool that runs on personal computers. The analysis has clearly shown that most of the cost for manufacturing high-value biopharmaceuticals that are expressed as fusion proteins, such as BHI, is associated with the recovery and purifica- tion of the product. The large number of separation steps required to reach high product purity lead to low recovery yield and high waste-to-product ratios (50,000: 1 in the base case).

Opportunities for process optimization and environmen- tal impact minimization must be considered at the early stages of process development before the process is frozen due to regulatory reasons. Based on the sensitivity analysis, we have found that the production cost is quite sensitive to process scheduling. The production cost is sensitive to plant capacity only at relatively low production rates (less than 1000 kgiyear). Finally, the profitability of a project involv- ing the development and commercialization of a pharma- ceutical substance is a strong function of up-front research and development expenditures.

The authors greatly appreciate the technical literature provided by Dr. Kevin Murphy from Pfizer Central Research. They also

thank Moshe van Berlo for his initial work on thc subject matter during his summer internship at NJIT.

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PETRIDES, SAPIDOU, AND CALANDRANIS: BIOSYNTHETIC HUMAN INSULIN PRODUCTION 541