Scale‐up economics for cultured meat

Abstract This analysis examines the potential of “cultured meat” products made from edible animal cell culture to measurably displace the global consumption of conventional meat. Recognizing that the scalability of such products must in turn depend on the scale and process intensity of animal cell production, this study draws on technoeconomic analysis perspectives in industrial fermentation and upstream biopharmaceuticals to assess the extent to which animal cell culture could be scaled like a fermentation process. Low growth rate, metabolic inefficiency, catabolite inhibition, and shear‐induced cell damage will all limit practical bioreactor volume and attainable cell density. Equipment and facilities with adequate microbial contamination safeguards have high capital costs. The projected costs of suitably pure amino acids and protein growth factors are also high. The replacement of amino‐acid media with plant protein hydrolysates is discussed and requires further study. Capital‐ and operating‐cost analyses of conceptual cell‐mass production facilities indicate economics that would likely preclude the affordability of their products as food. The analysis concludes that metabolic efficiency enhancements and the development of low‐cost media from plant hydrolysates are both necessary but insufficient conditions for displacement of conventional meat by cultured meat.

indicated in the figure, there is an expectation that this step would be carried out in large stainless-steel tanks, as in a large fermentation plant.
Precedent for this concept comes from the biopharmaceutical industry, where therapeutic proteins are produced in suspension cultures of recombinant mammalian cell lines in stainless-steel bioreactors up to ∼20 m 3 . Scaling animal cell culture like a fermentation process, however, presents several technical and economic challenges. Animal cells proliferate much more slowly than microbial cells. Metabolically unregulated cells in vitro tend to exhibit inefficiencies that cause them to produce growth-inhibiting catabolites such as lactate and ammonia. Mass-transfer limitations are expected in large bioreactors, where gas sparging and agitation are limited by the potential for shear-induced damage to animal cells, which lack a rigid cell wall. The capital costs of equipment and facilities with sterility safeguards adequate to avoid microbial contamination are likely to be high. Formulations of amino acids and protein micronutrients (growth factors) suitable for cell-culture media are not currently produced at scales consistent with food production and are also understood to be rather expensive.
These technical and economic aspects are explored in a recent assessment of cultured meat's potential to measurably displace the human consumption of conventional meat (Humbird, 2020). Its methods and findings are summarized in the present article. Basic industrial fermentation and bioreactor design rules are used to establish a stoichiometry of mammalian cell growth and attainable cell density as a function of bioreactor size. Cost projections for media components and sterile bioprocessing equipment are developed. These insights are used to develop production cost estimates for conceptual fed-batch and perfusion facilities that produce bulk animal cell mass. Suspension culture and con- This analysis considers an abstract mammalian cell line adapted for suspension culture. Cells are assumed to be spherical with 70% intracellular water and a hydrated mass of 3000 pg. Cell growth proceeds with a maximum growth rate µ max = 0.029/h, equivalent to a doubling time of 24 h. A CHON formula for animal dry cell mass (DCM a ) of CH 1.68 O 0.34 N 0.21 derives from an average composition of 15% lipid, 10% carbohydrate, 5% RNA/DNA, and 70% protein (Alberts, 2002). Sulfur, phosphorous, and metals are ignored and the CHON formula for protein (CH 1.57 O 0.31 N 0.28 ) derives from an average amino acid profile of 207 cellular proteins (Xie & Zhou, 2005).
An anabolic growth reaction for DCM a is constructed from the macromolecular composition: Lipid and carbohydrate are synthesized from glucose (Glc); nucleotides are synthesized from glucose and one of the N atoms on glutamine (Gln), rejecting glutamate (Glu); and protein is synthesized from its individual amino acid residues (Xie & Wang, 1994). Guan and Kemp (1999) This reaction is a superposition of respiration, glycolysis, and glutamine catabolism, and it can be generalized with two degrees of freedom, for example, lactate/glucose ratio (Lac/Glc) and glutamine/glucose ratio (Gln/Glc). As written, these ratios are relatively high (Lac/Glc = 1.77 and Gln/Glc = 0.13), indicating significant metabolic inefficiency consistent with the so-called Warburg effect, which is frequently observed in rapidly proliferating animal cells (Hosios et al., 2016). Catabolism proceeds F I G U R E 1 Conceptual cultured meat production process at a rate that meets an observed heat dissipation or "metabolic power." West et al. (2002) showed that in vitro metabolic power (2) Given the metabolic inefficiency inherent to Reaction 2, it will stand in for an unoptimized "wild-type" cell line. It will be demonstrated below, however, that its Lac and NH 3 generation rates preclude it from reaching an economically high cell density. Growth optimization strategies include: extensive characterization; selection for lactate reuptake; transfection of glutamine synthetase enzyme; and feedback control of glucose and pH (Freund & Croughan, 2018;Pereira et al., 2018). To permit a 20 m 3 fed batch carried out with this metabolism to remain under likely inhibition limits of Lac and NH 3 (to be discussed shortly), the maximum inefficiencies are rather Lac/Glc = 0.50 and Gln/Glc = 0.025. Reaction 3 therefore represents this "metabolically enhanced" cell line: At food scale, plant protein hydrolysates may be more costeffective and sustainable than amino acids produced individually by fermentation. Figure 2 compares the amino-acid profile implied in  catabolite accumulation rates, and other factors. As these limits are exceeded, the growth rate will drop precipitously due to inhibition.
At the already-low growth rate considered here, any inhibition is economically unacceptable with respect to the accumulation of bulk cell mass. In general, it will be more cost-effective to stop the batch and begin a new one at the uninhibited growth rate.

| Viscosity
A practical maximum cell density in suspension culture occurs at a volume fraction ф∼0.25; above this limit, viscosity increases sharply as cellcell collisions become more frequent (Iordan, 2008). For 3,000 pg cells with a diameter of 17.7 μm, the absolute maximum attainable cell density is thus 86 × 10 6 /ml or 258 g/L wet. Note that for smaller cells like CHO, the maximum number density can be significantly higher (Clincke et al., 2013), but the maximum mass density is the same.

| O 2 mass transfer
The volumetric oxygen transfer rate (OTR) is the product of a mass transfer coefficient k L a and a driving force: the deviation of the  To determine k L a, the correlation of Xing et al. (2009) was developed for CHO cell-culture media at 37°C: STR design rules for animal cell culture limit the gas sparge rate to u s = 0.006 m/s, which is equivalent to 0.1 vvm in a 20 m 3 bioreactor (Ozturk, 1996). Agitation is limited to a power input that creates turbulent eddies λ K on the length scale of a single cell (∼20 µm): where ν = η/ρ is the kinematic viscosity of the medium fluid and the factor of 50 corrects for the power input close to the impeller relative to the bulk P/V (Nienow, 2010 transfer rate is approximately equal to the OTR. The liquid concentration of CO 2 (typically measured as pCO 2 , or the CO 2 partial pressure in equilibrium with the liquid phase) is therefore a function of the bioreactor sparge rate. In CHO culture, inhibition is noted when pCO 2 falls outside a range of 40-100 mbar (Gray et al., 1996).
With sparging fixed at u s = 0.6 cm/s, the cell density must be limited such that the CO 2 evolution rate does not cause pCO 2 > 100 mbar.

| Mixing
Mixing time in the STR can be estimated with Equation 8 (Nienow, 2014). Here, T is the bioreactor diameter, D is the impeller diameter, and H L is the liquid height. Mixing time should be less than 1/k L a to ensure that dissolved O 2 is quickly transported away from the bubble (Van't Riet & Van der Lans, 2011).

| Catabolite inhibition
In fed-batch cell culture for biopharmaceuticals production, the accumulation of toxic and growth-inhibiting catabolites is a far more frequently encountered limit than the physical limits discussed so far.
Inhibiting concentrations of 2-10 mM NH 3 have been reported for mammalian cells; lactate inhibition is an order of magnitude higher (Xie & Zhou, 2005). For modeling purposes, the limits of 5 mM NH 3 and 50 mM lactate are considered here. These simulations can be extended to find the maximum fedbatch cell density as constrained by O 2 , CO 2 , NH 3 , and mixing. As shown in Figure 4b, NH 3 inhibition limits the cell density to 110 g/L in bioreactors <20 m 3 . At 20 m 3 , the NH 3 -limited density is coincident with the pCO 2 -limited density (the catabolic parameters of Reaction 3 were selected to cause this coincidence). Bioreactors >20 m 3 have a lower maximum cell density due to CO 2 inhibition.
While NH 3 inhibition can be mitigated to a degree with enhanced metabolic efficiency, CO 2 inhibition cannot. Regardless of the metabolic parameters, the respiratory quotient (CO 2 /O 2 ) of the growth reaction will remain ∼1 and the CO 2 stripping rate will remain equal to (or less than) the O 2 transfer rate. The only way to circumvent CO 2 inhibition in a sparged bioreactor is to sparge harder, possibly to the point of cell death (Al-Rubeai et al., 1995). This may preclude the scale-up of animal cell culture into extremely large bioreactors.
CO 2 and NH 3 inhibition can also be mitigated, to a degree, with perfusion. In perfusion culture, the contents of the bioreactor are continuously cycled through a cell retention device, which continuously removes extracellular products and inhibitors and generally permits higher cell densities than fed-batch cultures. At steady state, cells are bled from the bioreactor to maintain growth rate; in principle, this cell bleed could be a harvest stream for bulk cell culture. Perfusion culture volumes are limited by the capacity of the cell retention device. The alternating tangential-flow (ATF) filter is commonly used in high cell-density applications and the largest available (e.g., the Xcell ATF 10) can run at perfusion rates up to 1000 L/d (Pollock et al., 2013). In dual-ATF configurations, 1 m 3 bioreactors can thus be perfused at up to 2 reactor volumes per day (RV/d), and 2 m 3 bioreactors at up to 1 RV/d. Figure 4c presents curves of attainable cell density as a function of perfusion rate. The wild-type Reaction 2 generates 2 mmol NH 3 /mol DCM a -h. If NH 3 is removed via the perfusate stream at a steady-state concentration of 5 mmol/L, then at a perfusion rate of 2.0/d it can be computed that inhibition is reached at a cell density of only 20 g/L (6.8 × 10 6 /ml).
With the more efficient Reaction 3, an O 2 -limited cell density of 195 g/L (65 × 10 6 /ml) can be achieved at a perfusion rate of 1.0/d.

| Capital costs
The capital cost of a conceptual bulk animal cell-culture process is developed from the bare-equipment costs of its most important items. From this purchased equipment cost, a total capital investment (TCI) is obtained through the application of cost escalation factors, which are understood to be rather high for biopharmaceutical cell-culture processes (Petrides, 2015). Compared to existing biopharmaceutical processes, future bulk cell-culture processes for cultured meat would probably have similar safety and sterility requirements to protect the culture from being overcome by con- Estimated bare equipment (vessel and agitator) and total system costs for bioreactors of 1-200 m 3 are further shown in Table 1,

| Amino acids
Each amino acid is currently produced at some commercial scale ranging from thousands of kTA for animal-feed supplements to <1 kTA for aminos with primarily pharmaceutical uses. Figure 5a presents available price-volume data for individual amino acids individual amino acid, the market volume of a novel "cultured meat-grade" formulation is probably a more reliable predictor of price at scale than any specific details of its manufacturing process. These will certainly influence price, but not by multiple orders of magnitude. Prices for individual amino acids at scale are therefore correlated to estimated production volume by the following equation:

| Plant protein hydrolysate
Soybean hydrolysate was discussed above as a potential alternative source of amino acids. In animal-free media formulations for cell culture, plant hydrolysates are sometimes used supplementally in otherwise chemically defined media (Babcock et al., 2007). Single amino acids are not generally produced from hydrolysates because posthydrolysis isolation is cost-prohibitive.
Instead, whole hydrolysates would have to be designed to provide all amino acids in the proper ratios (with the probable exception of glutamine, which decomposes easily). Current soybean meal pricing for animal feed is about $0.33/kg (USDA-IL, 2020). Subtilisin protease enzyme at $15/kg (see Figure 5b) and 2% loading would further add $0.30/kg of meal processed. If meal were 48% protein, protein 88% soluble, and hydrolysis conversion 80%, then a formulation of mixed amino acids from hydrolysis could cost as little as $1.60/kg, plus processing costs. This analysis considers a price of $2/kg mixed amino acids.

| Protein micronutrients
Protein growth factors are provided in media to regulate growth and metabolism. Animal-free sources of these proteins include highly processed extracts of plant proteins and recombinant proteins made by fermentation. Commercial recombinant proteins (therapeutics, industrial enzymes) have a price-volume relationship similar to amino acids, as presented in Figure 5b and regressed in Equation 12 (Arbige, 1989;Gotham et al., 2018;Kelley, 2009).

| Fixed costs
Facility overhead is scaled to CAPEX and is here taken to comprise 4% TCI/y for maintenance (including CIP) and 5% TCI/y for insurance. Labor is quantified in terms of operator attention per batch, based on a series of task/time assumptions. A salary of $50,000/y for regular FTEs is taken from the U.S. Bureau of Labor Statistics-Chemical Plant Operator; the supervisor's salary is 140% of this. A labor burden of 100% is added to the total labor cost.

| Fed-batch case study
A process flow diagram (PFD) of a conceptual fed-batch cellculture process is given in Figure 6a. For the detailed example given below, the model cell-culture facility is designed with 24 × 20 m 3 production bioreactors and produces 6.8 kTA of wet cell mass. Cells are propagated from the lab through a seed train to the production bioreactors. Upon harvest, the cell mass is dewatered to 20% solids in a disk-stack centrifuge. Two large media tanks attached to HTST sterilizers are shared between bioreactors to provide pre-inoculation fill. A smaller, dedicated tank attached to sterile retention filters contains the pre-mixed media that will be added during the batch.
The fed-batch simulation in Figure 4a is used to size equipment and compute media and utilities usages. Bioreactor costs are estimated with Equation 9, while costs for the remaining equipment and buildings are estimated as described above. A Class 8 clean room is selected for cellculture areas and Class 6 for laboratory areas. The sum of equipment and buildings gives a TDC of $94 M. Indirect costs are factored from the TDC to give a TCI of $328 M. As an annual capital charge, this TCI is equivalent to $48 M/y, or ∼$12/kg of wet cell mass. At a total production volume of 100 kTA, macronutrients (amino acids) contribute an additional $19/kg and micronutrients (growth factors) $3/kg. Consumables (filter membranes), utilities (clean room power) and labor (95 total FTE) contribute a combined $3/kg. These capital and operating costs are summarized in Table 2. The overall cost of production estimated for a fed-batch cell-culture process is $37/kg wet cell mass.  Figure 7b indicates that 24 production bioreactors are optimal in a single facility; the production cost increases at >24 bioreactors because the clean room area grows faster than the process volume it contains. Figure 7c presents the sensitivity associated with the production bioreactor size, indicating an optimal volume of 50 m 3 . At larger volume, the reduction in final cell density (due to pCO 2 limitations) outweighs the cost benefits of larger reactors.

| Perfusion case study
A PFD of a conceptual cell-culture process based on perfusion technology is given in Figure 6b. The 2 m 3 production bioreactors are inoculated by expanding cells from the lab through 125 and 500 L seed bioreactors in fed-batch mode. Upon inoculation, the production bioreactor also operates in fed-batch mode up to 120 g/L. At this density, perfusion begins at a rate of 1.0/d (requiring 2× ATF 10 filters) and the retained cell density rises to 195 g/L per Figure 4c. Each production bioreactor operates with continuous cell harvest except during 10 scheduled turnarounds (72 h each) per year. The average time at steady state is ∼91%, and 9 kg/h of wet cell mass is harvested from each bioreactor. To produce 6.9 kTA (to match the fed-batch process above), 96 bioreactors are required. Equipment and building costs are estimated as described previously. Bulk amino-acid and glucose tanks are shared for initial filling of all bioreactors, and each production bioreactor gets two dedicated tanks to hold 24 h of makeup media each. A cost for the ATF 10 perfusion devices is taken from Pollock et al. (2013). The TCI developed in Table 2 is $663 M, giving an annual capital charge of $97 M/y or ∼$23/kg wet cell mass. Macro-and micronutrient costs are similar to the fed-batch process, while consumables costs are significantly higher due to the replace-ment of ATF membranes (20 per bioreactor per year at $16k each). The perfusion process is also slightly more labor-intensive than the fed-batch process (132 total FTE) and its overall cost of production is estimated as $51/kg wet cell mass. The limited volume, relatively high bioreactor direct costs, and the CAPEX and consumables associated with the perfusion device present significant disadvantages. Nutrient costs can be minimized by assuming a much larger global production volume, as shown in Figure 8a, or reduced $16/kg by substituting low-cost hydrolysate. Figure 8b presents sensitivity to the facility production rate, noting a very weak minimum at 3.5 kTA. The issue with divergent clean room cost is also present for perfusion, but to a much smaller degree. In a 2 m 3 bioreactor, a retained cell density of 195 g/L is attained at a perfusion rate of 1.0/d, which requires dual ATFs. With a single ATF and a perfusion rate of 0.5/d, 140 g/L of cells can be retained. As shown in Figure 8c, the capital and consumables costs associated with the second ATF offset the economic benefit of higher density, such that the production cost at 195 g/L is hardly any better than at 140 g/L.

| CONCLUDING DISCUSSION
To reach a market of 100 kTA, or ten million consumers consuming 10 kg/y, it must be assumed that cultured meat has at least attained the price-acceptance status of a reasonably affordable "sometimes" food. To assert a threshold on the subjective metric of affordability, this analysis submits a target of~$25/kg of wet animal cell matter produced in a bulk growth step. After further processing, packaging, distribution, and profit, unstructured products made 100% from bulk cell mass at $25/kg might be expected to reach a minimum of $50/kg at the supermarket: The price of a premium cut of meat, paid instead for a mincemeat or nugget-style product. Above this cost, the displacement of conventional meat by cell culture may arguably be measurable but increasingly less significant.
Although both estimates detailed above exceed this threshold, a fed-batch process could potentially be brought under $25/kg with low-cost hydrolysate media. The same is not true of the perfusion process, which has capital costs and capital-dependent fixed costs that are well above the target. Hydrolysates appropriate for whole, unsupplemented cell-culture media do not exist today and the assertions of their ultimate suitability and price are somewhat speculative. Further recall that both processes were examined with a cellular metabolism significantly enhanced relative to a wild-type cell line, implying extensive characterization, process development, and metabolic engineering. From the modeling above, it can be concluded that metabolic efficiency and low-cost hydrolysate media development can both be taken as necessary but insufficient conditions of affordability. Capital cost reduction is a secondary condition at best.
T A B L E 2 TEA estimates for the production of bulk cell mass by fed batch or perfusion