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Production of Single Cell Protein from Natural Gas John Villadsen Center for Biochemical Engineering Technical University of Denmark.

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Presentation on theme: "Production of Single Cell Protein from Natural Gas John Villadsen Center for Biochemical Engineering Technical University of Denmark."— Presentation transcript:

1 Production of Single Cell Protein from Natural Gas John Villadsen Center for Biochemical Engineering Technical University of Denmark

2 Genome of Methylococcus capsulatus

3 The bacteria with membrane bound Methane-monooxidase

4 Dividing M. capsulatus with clearly visible membranes

5 The key-enzyme Methane monooxygenase

6 Capture of CH 4 by Methane monooxygenase

7 Further oxydation of methanol in the organism

8

9 Methane and Oxygen demand for SCP production From 1.25 kg methane one obtains 1 kg biomass *) This corresponds to 1 kg biomass per 1.75 N m 3 methane or Y sx = C-mole biomass per C-mole methane The O 2 demand is (8 – 0.520٠4.20) / 4 = 1.45 mol O 2 per C-mole CH 4 or 2.53 N m 3 O 2 / kg biomass = 3.62 kg O 2 / kg biomass. Stoichiometry of methane conversion to biomass: CH O NH 3 → 0.52 CH 1.8 O 0.5 N CO H 2 O *) Reference : Wendlandt, K.D, Jechorek, M, Brühl, E. ”The influence of Pressure on the growth of Methanotrophic Bacteria” Acta Biotechnol. 13, (1993) and industrial experience: Dansk Bioprotein A/S present.

10 Demand for heat removal The reaction should take place at ≈ 45 o C, the optimal temperature for Methylococcus capsulatus fermentation. Stoichiometry: CH O NH 3 → 0.52 CH 1.8 O 0.5 N CO H 2 O Heat of reaction  460Y so kJ (C-mol carbon source) -1 or Q = 460٠ 1.45 = 667 kJ (mole CH 4 ) -1 = 52 MJ (kg biomass) -1 This is an appreciable heat duty!

11 Demand for O 2 and CH 4 mass transfer The production rate depends on the rates of two separate processes A. The reaction between bacteria and dissolved O 2 + CH 4 B. The rate of mass transfer from gas- to liquid phase. The ”bio-chemical” reaction is limited by NH 3 since we need to keep the NH 3 concentration below about 40 mg L -1 to avoid formation of NO 2 - which is toxic to the bacteria. At 30 mg L -1 the rate is q x = 0.21 X kg m -3 h -1 where X is the biomass concentration in kg m -3. But q o2 = (1.45 / 0.52)(1000 / 24.6) q x = 113 q x mol m -3 h -1 = k l a (c O2 * - c O2 ) where c O2 * and c O2 are respectively the saturation and the actual O 2 concentrations in the liquid.

12 Factors that affect the mass transfer The rate of mass transfer k l a (c O2 * - c O2 ) (and k l a (c CH4 * - c CH4 )) depend on : A.The mass transfer coefficient k l a Maximum achievable k l a ≈ 1200 h -1 B.c O2 * C.c O2 c O2 * is proportional with the partial pressure of O 2 in the gas phase. At 1 atm total pressure and pure O 2 one obtains c O2 * = 0.9 mM (45 o C) c O2 should be above about 20 μM to keep the organism healthy.

13 The switch from bioreaction control to mass transfer control Assume that we wish to have X = 20 kg m -3 (q x = 4.2 kg m -3 h -1 ) q O2 = 113 ٠ 4.2 = 475 mol m -3 h -1 = k l a (c O2 * - 20) mol m -3 h -1 For k l a = 1000 h -1 c O2 * must be > 495 μM to obtain a gas transfer rate that is higher than the rate of the liquid phase reaction 4.2 kg m -3 h -1. For a total pressure of 1 atm and pure O 2 (c O2 * = 900 μM) about 50 % of the oxygen is consumed before O 2 limitation sets in. With O 2 extracted from air (21% O 2, c O2 * = 189 μM) oxygen limitation prevails throughout the reactor. With pure oxygen and 4 atm total pressure (c O2 * = 3600 μM) O 2 limitation occurs only in the last ≈ 14 % of the reactor.

14 Consequences of O 2 limitation The constant production rate q x = 4.2 kg m -3 h -1 can not be maintained The production rate in the last part of the reactor is 1 st order in c O2 * If we wish a high utilization of O 2 (e.g. 95 %) the reactor volume may increase beyond reasonable limits (or q x may decrease to an unacceptably low level).

15 Reactor design A stirred tank reactor is hopeless: We wish the first order conversion of O 2 in the last part of the reactor to proceed in plug-flow mode. In a CSTR c O2 * would be 0.05 of inlet value. The large heat release dictates that external heat exchange is to be used. Liquid and gas is forced through a number of stationary mixer elements at a velocity of ≈ 1 m s -1. Gas is injected through an ejector. Ample allocation of head space assures gas/liquid separation. Holding time for liquid ≈ 5 h and for gas ≈ 60 s. Centrifuges (or drum filters) are used to separate biomass from liquid. Ultrafiltration gives ≈ 20 wt% biomass sludge. Spray drying gives the final powdery product Heat shock treatment (123 o C, 2-5 min) removes nucleic acids and gives a product suitable for direct human consumption.

16 500 L pilot plant loop-fermentor at DTU

17 Design of a 10 m 3 loop reactor

18 A 10 m 3 fermentor

19 250 m 3 reactor (≈ 9000 t year -1 production) in Norway


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