January 2004 • NREL/MP-560-35592 

Biological Water-Gas Shift 

Conversion of Carbon 

Monoxide to Hydrogen 

Milestone Completion Report 

Wade A. Amos 

National Renewable Energy Laboratory 

1617 Cole Boulevard 

Golden, Colorado 80401-3393 

NREL is a U.S. Department of Energy Laboratory

Operated by Midwest Research Institute • Battelle 

Contract No. DE-AC36-99-GO10337 
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Biological Water-Gas Shift Conversion 

of Carbon Monoxide to Hydrogen 

Wade A. Amos 

National Renewable Energy Laboratory 

Milestone Report for the U.S. Department of Energy HFCIT Program

Process Analysis Task 

September 2003 – Final Version 
Table of Contents 

1.0 Introduction................................................................................................................... 1 

2.0 Conventional Catalytic Water-Gas Shift ...................................................................... 1 

3.0 Biological Water-Gas Shift Reaction............................................................................ 2 

4.0 Bio-Energetics of the Biological WGS Reaction.......................................................... 4 

5.0 Biological WGS Kinetics.............................................................................................. 4 

6.0 CO Inhibition ................................................................................................................ 5 

7.0 Mass Transfer Limitations ............................................................................................ 7 

8.0 Pressurized Operation ................................................................................................... 8 

9.0 Reactor Designs .......................................................................................................... 12

10.0 Process Economics.................................................................................................... 14 

10.1 Biomass-Derived Synthesis Gas with a Trickle-Bed Reactor .............................. 15 

10.2 Industrial Collaboration ........................................................................................ 15 

10.3 Drop-in replacement in SMR................................................................................ 16 

10.4 Non-reforming applications.................................................................................. 17 

10.5 Interface to a Gasifier............................................................................................ 19 

10.6 HyCO Plants ......................................................................................................... 19 

11.0 Conclusions............................................................................................................... 20 

12.0 References................................................................................................................. 21 
1.0 Introduction 

The purpose of this report is to summarize the results of research and economic analysis 

on the biological water-gas shift process being studied at the National Renewable Energy 

Laboratory. The report describes some of the technical issues regarding the process, 

addresses some claimed benefits of the process and presents some economic results from 

economic studies using different process configurations. 

While the technical feasibility of the biological water-gas shift process has been 

demonstrated, there are only limited cases where replacing conventional catalytic 

reforming with a biological alternative makes economic sense. 

2.0 Conventional Catalytic Water-Gas Shift 

The water-gas shift (WGS) reaction is used to convert carbon monoxide (CO) to carbon 

dioxide (CO2) and hydrogen (H2) through a reaction with water (H2O). 

CO + H2O Æ CO2 + H2 (1) 

The reaction is exothermic, which means the reaction equilibrium shifts to the right and 

favors the formation of the H2 and CO2 products at lower temperatures. At higher 

temperatures, the equilibrium shifts to the left, limiting complete conversion of CO to H2. 

The reaction is the basis for most of the industrial H2 produced in the world from 

methane (CH4) in natural gas through steam-methane reforming. Methane is first 

reformed to a mixture of CO, CO2 and H2 in the presence of steam over a nickel catalyst. 

A conventional water-gas shift reactor then uses a metallic catalyst in a heterogeneous 

gas-phase reaction with CO and steam. Although the equilibrium favors formation of 

products at lower temperatures, the reaction kinetics are faster at elevated temperatures. 

For this reason, the catalytic water-gas shift reaction is initially carried out in a hightemperature shift (HTS) reactor at 350-370°C. Conversion in the HTS reactor is limited 

by the equilibrium composition at the high temperature. To achieve higher conversions 

of CO to H2, the gas leaving the HTS reactor is cooled to 200-220°C and passed through 

a low-temperature shift (LTS) reactor (Kirk-Othmer, 1995; Ulllman’s, 1989). 

Approximately 90% of the CO is converted to H2 in the first HTS reactor and 90% of the 

remaining CO is converted in the LTS reactor. 

Figure 1 – Catalytic WGS process. 

1
Before pressure-swing adsorption (PSA) purification came into wide use, the preferred 

hydrogen purification method was HTS, LTS, CO2 scrubbing/absorption, followed by a 

methanation reaction to remove any remaining CO: 

CO + 3 H2 Æ CH4 + H2O (2) 

Because the CO consumes part of the H2 product during the methanation reaction, H2 

production systems using methanation steps would include a LTS reactor to convert as 

much CO to H2 as possible before methanation. Otherwise, large amounts of the product 

H2 would be consumed during methanation. 

With PSA, both the CO2 and un-reacted CO are adsorbed, producing hydrogen purities 

over 99.9%. Because the CO is removed to very low levels (less than 10 ppm), the 

methanation step is no longer needed. The CO2 scrubbing step can also be eliminated 

with PSA purification, which simplifies the system and reduces the overall system cost. 

Without methanation, there is less need to remove CO before the PSA, so the LTS reactor 

is sometimes eliminated to reduce capital costs. If desired, the HTS reactor operating 

temperature can be adjusted to balance higher conversion versus a larger reactor size. 

Eliminating the LTS reactor results in slightly lower hydrogen yield, but this is balanced 

against lower capital costs. 

3.0 Biological Water-Gas Shift Reaction 

The organism Rubrivivax gelatinosus CBS is a non-sulfur, purple photosynthetic bacteria 

that is capable of performing the same water-gas shift reaction under anaerobic 

conditions, atmospheric pressure and an ambient temperature of 25°C. Wolfrum (2003) 

has shown the organism can perform the shift reaction at pressures up to 4 atmospheres 

absolute pressure. It is believed that the organisms can be effective at even higher 

pressures with increased cell loading to the reactor, but these tests have not been 

performed. 

The organism uses the biological WGS reaction as a means to obtain energy to maintain 

metabolic processes and grow. In the wild, Rx. gelatinosus CBS can obtain energy 

through photosynthesis, aerobic heterotrophic metabolism (e.g., acetate or malate) and 

anaerobic fermentation pathways. The biological WGS represents an anaerobic, darkphase pathway. This pathway produces much less energy for metabolic activities as 

compared with either the photosynthetic or aerobic pathways. This lower energy 

production results in a slower cellular growth rate which increases the time needed to 

reach steady-state, but reduces the amount of waste cell mass produced by the biological 

WGS process once it is operating. The benefit of the dark-phase reaction is that the 

biological WGS system can be operated in a conventional closed reactor, similar to 

trickling filters used for wastewater treatment—no expensive photo-bioreactor is 

required. 

The energy obtained through the biological WGS is obtained by transferring electrons 

from CO to H2O in the following coupled reactions: 

2
CO + H2O Æ CO2 + 2e- + 2H+ (3) 

H+

 + 2e- Æ H2 (4) 

CO + H2O Æ CO2 + H2 (5) 

This reaction produces 4.46 kcal/mol CO reacted. For comparison, the aerobic 

conversion of 1 mol of CO to CO2 would produce 61.1 kcal/mol CO. 

CO + O2 Æ CO2 (6) 

The aerobic reaction would result in more energy to the organism, resulting in more cell 

growth per mol CO but no H2 would be produced under aerobic conditions. It is 

desirable to minimize cell growth so less excess cell waste is produced, but if the growth 

rate is less than the natural death rate, the reactor operation will not be stable. In any 

biological system, cells are continually dying. Many of the nutrients in the dead cells are 

recycled when the cells lyse or split, but a portion of the cell remains as refractory or inert 

material that cannot be “re-digested” and used for new cells. (This can be thought of as 

the “teeth and bones” of the cells—insoluble minerals that cannot dissolve.) In many 

batch systems, not enough refractory material builds up to affect operation of the system, 

but in continuous systems, the accumulation of this refractory material must be removed. 

Cell growth rates are also important for reactor startup and recovering from process 

upsets. Initial reactors can be “seeded” or inoculated with a sterile culture, but most of 

the cells used to convert the CO to H2 come from new cell growth after the inoculation. 

The higher the growth rate, the faster the reactor will reach full production capacity. 

Upsets in pH or temperature can also result in a loss of biological activity due to damage 

to or death of the cells. Higher growth rates would allow quicker recovery in the reactor. 

One possibility during startup is to provide the Rx. gelatinosus CBS with “food” in the 

form of acetate, malate or a low-cost sugar source (i.e., corn-steep liquor) and provide 

oxygen to temporarily boost the energy production and growth rate of the cells. Once the 

oxygen is removed, there will be some delay while the cells produce new enzymes for 

their anaerobic pathway, but the cells will eventually switch back to using CO for 

generating energy and H2. 

One concern with using this approach to boosting cell growth is that there are many 

organisms that could potentially use acetate and oxygen for growth. If there is any 

contamination, other cells may grow faster than the Rx. gelatinosus CBS and take over 

the reactor. When switched back over to CO, little or no H2 would be produced. There 

are believed to be fewer organisms that can grow in the presence of only CO, so startup 

on CO may be a safer approach, even though it may take more time. 

If reactor upsets can be avoided, it should be possible to run the reactor for months 

without the need to shutdown and restart the reactor. At the same time, the process must 

be designed so if there is any contamination, the system can be quickly sterilized and reinoculated. Ideally, if growth conditions can be determined that favor Rx. gelatinosus 

CBS over most other organisms, contamination would not be a concern because Rx. 

3
gelatinosus CBS would dominate and the other organisms would die off. Currently, 

these growth conditions have not been identified, but laboratory-scale reactors have been 

run for months with little or no loss in H2 production capacity. One other possibility 

would be to transfer the genes responsible for the biological WGS reaction into another 

organism that is easier to culture and control. This molecular engineering approach 

would, however, take significant effort. 

4.0 Bio-Energetics of the Biological WGS Reaction 

Even without knowing the exact metabolic pathways and enzymes involved in the 

biological WGS reaction, the overall free energy change of the chemical reactions can be 

used to determine the maximum amount of energy the Rx. gelatinosus could obtain from 

performing the biological WGS. By assuming an average energy efficiency of 60% for 

biological processes (Gossett, 1995), it is possible to estimate the amount of cells 

produced per g of CO fed to the bacteria. Under anerobic conditions, Rx. gelatinosus 

should produce 1.4 g of cells per 1 mol of CO. When growing aerobically on O2 and 

acetate, the cell yield would be 2.0 g of cell mass per mol of acetate. 

What bio-energtic calculations provide is an estimate of the yield for any proposed 

biological reaction. Bio-energetics also do not provide any information on the kinetic 

rates of biological reactions. Rates depend upon the efficiencies of the enzymes involved 

and reactions may not occur at all if the required enzymes are not present in the 

organism. Bio-energetics can also help determine if reactions are likely to become more 

favorable or less favorable depending on environmental conditions, such as pH, 

temperature or feed concentrations. (When a reaction becomes “more favorable” it 

produces more energy for the organism. Depending on the environmental conditions, 

some reactions may shut down completely because they become thermodynamically 

unfavorable.) 

Calculations were done to determine the effects of CO concentration, CO2 concentration, 

H2 concentration and pH have on the bio-energetics of the biological WGS reaction. The 

effects of pH, CO concentration and CO2 concentration were also confirmed in the 

laboratory. 

5.0 Biological WGS Kinetics 

As mentioned in the previous section, the kinetics of a reaction can only be determined 

experimentally. This is true in catalytic WGS reactions as well. However, the maximum 

reaction rate for the biological WGS reaction can be estimated using assumptions on the 

ability of organisms to transfer electrons to external electron acceptors. Measurements 

show that most organisms transfer electrons to external electron acceptors are a rate of 

approximately 1 mole of electrons per g cell dry weight (cdw) per day (Gossett, 1995; 

Perry, 1971, 1972). In the case of H2, 2 electrons are required per molecule of H2, so 0.5 

mol of H2 can be produced per day per g of viable cells. With a molecular weight of 2 

g/mol, this results in a convenient rule of thumb for biological hydrogen production: 

4
1 g of cells can produce 1 g of H2 per day 

(or 1 kg of cells can produce 1 kg of H2 per day) 

This is based upon a temperature of 25°C and the value can range from 0.5 to 2.0 g H2/g 

cells/day depending upon the organism (Gossett, 1995; Perry 1971, 1972). Most 

biological reactions double their reaction rate for every 10°C increase in reaction 

temperature, but that does not mean increasing the reaction temperature is the answer. 

While different organisms have different temperature optima depending on the 

environment to which they have adapted, most organisms can only operate over a small 

temperature range before its enzymes are denatured and the organism dies. Rx. 

gelatinosus can most likely be used in a reactor operating at up to 37°C, but above this 

temperature, the organism’s water-gas shift activity will drop rapidly and the organism 

will die. If the enzymes specific to the biological WGS reaction are temperature stable, it 

might be possible to transfer the enzymes into an organism isolated from a hightemperature environment, but this requires advanced biological techniques. It would also 

require more expensive, insulated reactors designed to operate above ambient conditions. 

It also appears that the hydrogenase enzyme responsible for H2 production in Rx. 

gelatinosus requires auxiliary genes that assist in synthesizing the enzyme within the cell 

(Wolfrum and Maness, 2003). These genes would need to be present in the new host 

organism’s chromosomal DNA or these auxiliary genes would also need to be transferred 

to the host organism. 

It is important to note that the above kinetic rate is based upon viable cell mass. In 

biological reactors operated over long time periods, the refractory cell material may 

build-up. The observed cell concentration may represent 50% or less viable cell mass 

because of the build-up of dead cells and inert material. 

Laboratory estimates of the actual biological shift rates agree with the above estimate. 

6.0 CO Inhibition 

Another concern with chemical reactions in general is inhibition. For example, the nickel 

used in conventional SMR catalysts is sensitive to sulfur. In biological process, reactants, 

products or contaminants can reduce the growth and/or production rate of organisms, if 

the organisms aren’t simply killed. In studying the biological WGS reaction, some 

potential sources of inhibition are CO and CO2. The CO2 also has an effect on pH, 

especially at higher pressures due to the formation of carbonic acid and its carbonate 

derivatives. 

Studies of the intrinsic rate of the biological WGS reaction show inhibition at high 

dissolved CO levels. Figure 2 shows the trend of biological shift rate versus the liquidphase CO concentration. These results were from shake flask experiments conducted 

with dilute cultures under conditions that approached equilibrium CO concentrations. 

This data was collected under ambient pressure, so the maximum CO concentration 

observed by the cells was the equilibrium concentration of CO at one atmosphere of 

pressure, or approximately 1.0 mmol/L. Increasing the operating pressure increases the 

5
liquid-phase CO concentration. Liquid-phase concentrations greater than 0.15 mmol/L 

result in decreased reaction rate and eventual shutdown of the biological WGS pathway. 

High-pressure shake flask experiments were not performed due to the more complex 

experimental procedures required, however, pressurized experiments were conducted 

using 1-liter continuous reactors and the effects of inhibition were observed (Wolfrum 

and Maness, 2003). The reaction rate in Figure 2 is presented per unit cell mass. Higher 

cell mass in a reactor results in higher potential reaction rates per unit reactor volume. 

1.0 

0.8 

0.6 

0.4 

0.2 

0.0 

0.0 0.2 0.4 0.6 0.8 

Dissolved CO Concentration (mM) 

Figure 2 – Carbon Monoxide Inhibition of the Biological WGS Reaction 

Because one possible application of the biological water-gas shift reaction is the 

conditioning of synthesis gas from a gasifier, there was some concern that higher 

molecular weight contaminants, such as benzene, might adversely affect the cells. To 

alleviate those concerns, a biological reactor was run using quenched, biomass-derived 

synthesis gas produced by NREL’s Thermo-Chemical User Facility. No significant 

adverse effects were seen on the organism’s health or biological WGS activity due to 

acetylene, ethylene, benzene concentrations of 5000 ppmv and toluene concentrations of 

3000 ppmv (Wolfrum and Watt, 2001). 

In additional, a specific refinery gas stream was also tested using Rx. Gelatinosus CBS 

through a Work-for-Others agreement with an industrial partner. No significant effects 

were seen on growth rates of H2 production rates for the gas compositions tested. (The 

specific gas compositions and source of the gas stream are proprietary information.) 

Whatever slight inhibition that was observed was attributed to pH effects in samples 

containing high CO2 levels. 

With regard to pH, biological WGS activity remained high over a range of pH 6.8 to 8. 

Below 6.5 pH, the cells lyse and die. In the trickling filter reactors (described in Section CO Uptake Rate (mmol min-1g-1) 

6
8), biological WGS activity was lost at pH above 8. This may have been due to an 

inability of the organisms to remain attached to the trickling filter packing at high pH. 

The pH was expected to drop with increased pressure due to higher concentrations of 

dissolved CO2, but bubble column reactors (see Section 8) were operated at pressures up 

to 3 atmospheres absolute pressure without any obvious losses of activity. 

No tests were conducted to determine if H2 was inhibitory, but it is known that Rx. 

gelatinosus CBS can uptake H2 to support cell growth, so the main effect may be reduced 

hydrogen yield and not reduced production rates. 

7.0 Mass Transfer Limitations 

In a chemical reactor, there are two limitations on overall reaction rate: the intrinsic 

reaction rate and the mass transfer rate. The intrinsic reaction rate is the maximum 

reaction rate for the chemical reaction for given the concentration of reactants and 

temperature. If mass transfer is not a limiting factor, the overall reaction rate observed 

will very closely equal the intrinsic reaction rate. However, when mass transfer 

limitations is a limiting factor, the reaction can’t proceed faster than the reactants can be 

transported to the reaction site. Even though the reaction may be inherently fast, if the 

reactants can’t be supplied fast enough, the overall reaction will be slow. In the case of 

the biological WGS reaction, the transfer of CO from the bulk gas phase into the liquid is 

a limiting factor for most reactor configurations. The organisms are waiting to consume 

CO once it makes it into solution, but the mass transfer rate into solution is slow. 

The following equation governs mass transfer from the gas phase to the liquid phase: 

N = kLa * (C*

CO,l – CCO,l) (7) 

where N = the mass transfer rate 

kL = the mass transfer rate coefficient 

a = the interface area 

C*

CO,l = the equilibrium CO concentration at the interface of the gas and liquid 

CCO,l = the CO concentration in the liquid. 

For sparingly soluble gases like CO, C*

CO,l is small. The limited solubility of CO is one 

of the factors affecting the slow mass transfer rate. In a catalytic WGS reactor, all the 

reactants are in gaseous state (i.e, the water is steam). The mass transfer from the bulk 

gas into the pores of the catalyst are diffusion limited, but the rate of mass transfer is 

much higher than the rate of mass transfer from a gas into a liquid. 

Looking at Equation 7, we can see that there are several ways to increase the mass 

transfer. First, an increase in kL will improve mass transfer. This is generally 

accomplished by increasing the turbulence, often by increasing the flow rate of the liquid 

and/or gas. The second method of increasing mass transfer is by increasing the surface 

area for mass transfer. In a bubble column, more bubbles and smaller bubble sizes result 

in higher surface area and increased mass transfer. Third, increasing the bulk gas 

7
concentration of the CO in the gas will increase the concentration gradient (the term in 

parenthesis) increasing mass transfer. Lastly, if the CO is consumed almost immediately, 

the concentration of CO in the liquid will be low, improving mass transfer. 

If mass transfer rate is higher than the reaction rate, the CO concentration in solution will 

increase until it reaches the equilibrium concentration at the gas-liquid interface. At that 

point, the concentration gradient will be zero and no gas transfer will take place. Note 

that in Figure 2, if the CO concentration exceeds 0.15 mmol/L, the reaction rate starts 

dropping, causing even higher CO concentrations and more inhibition. 

In the shake flask experiments used to obtain the intrinsic reaction rates, very high shaker 

speeds are used to provide high turbulence and increased gas-liquid interface area to 

improve mass transfer. Cell concentrations in the shake flasks were also kept low so that 

the biological reaction rate was much less than the mass transfer rate when shaking. It is, 

however, very difficult to duplicate these conditions in a commercial-scale reactor. 

8.0 Pressurized Operation 

One method of increasing mass transfer is to pressurize the reactor. Pressurizing the 

reactor increases the equilibrium concentration of CO at the gas-liquid interface, 

increasing mass transfer into the liquid. At ambient pressures, an increase in CO 

concentration in the gas-phase increases the equilibrium concentration of CO and 

increases the mass transfer rate. 

If the reactor shift rate is limited by the mass transfer, operating at increased pressure will 

improve mass transfer and increase the overall reaction rate. However, increasing the 

reactor pressure can increase the mass transfer rate above the intrinsic biological reaction 

rate. When this happens, the liquid-phase CO concentration builds up and eventually 

inhibits the biological reaction, reducing hydrogen production. 

A simple graphical model can be used to compare the intrinsic bioreactor rate with the 

mass transfer limitation for a give reactor. Under steady-state conditions, there will be a 

liquid-phase CO concentration where the rate of biological uptake of CO exactly matches 

the bulk mass transfer rate of CO into the liquid. Figure 3 shows the gas transfer 

operating line superimposed on the bioreactor rate curve for an arbitrary reactor design. 

The point where these two lines cross is the steady-state bioreactor rate and the steadystate liquid-phase CO concentration. In Figure 3, this represents a stable operating point 

because if the liquid-phase CO concentration increases slightly, the mass transfer rate 

will drop slightly (moving right along the sloped line) and the bioreactor reaction rate 

will increase slightly, pulling the CO concentration back down to the steady-state point. 

If the CO concentration drops slightly, the mass transfer rate will increase while the 

bioreactor rate drops slightly. This will force the reactor back to the steady-state CO 

concentration. 

In Figure 4, the pressure has increased to 2 atm. The bioreactor rate curve remains the 

same, but the y-axis intercept (representing the maximum mass transfer rate) has 

8
increased and the x-axis intercept (representing the equilibrium liquid-phase CO 

concentration) has increased. Compared to Figure 3, the steady-state point has moved to 

a higher CO concentration and a higher bioreactor rate. 

Total Pressure = 1 atm, Cell Mass = 1 g/L, kla = 0.05/sec 

0.0020 

0.0015 

0.0010 

React ion 

Tr ansf er 

0.0005 

0.0000 

0.0000 0.0001 0.0002 0.0003 0.0004 0.0005 0.0006 0.0007 0.0008 0.0009 0.0010 

Liquid CO Conc. (mol/L) 

Figure 3 – Steady-State Mass Transfer 

Total Pressure = 2 atm, Cell Mass = 1 g/L, kla = 0.05/sec 

0.0020 

0.0015 

0.0010 Reaction 

Transfer 

0.0005 

0.0000 

0.0000 0.0001 0.0002 0.0003 0.0004 0.0005 0.0006 0.0007 0.0008 0.0009 0.0010 

Liquid CO Conc. (mol/L) 

Figure 4 – Increased Operating Pressure Increases Mass Transfer 

In Figure 5, the pressure has been increased to 3 atm. The mass transfer line now crosses 

the bioreactor rate curve to the right of the peak bioreactor reaction rate. This is an 

unstable operating point because if the CO concentration increases slightly, the reaction Rate (mol/min) Rate (mol/min) 

9
rate drops, which will push the CO concentration even higher and cause the bioreactor 

reaction rate to drop even further, until the biological WGS reaction effectively shuts 

down. 

Total Pressure = 3 atm, Cell Mass = 1 g/L, kla = 0.05/sec 

0.0020 

0.0015 

0.0010 Reaction 

Transfer 

0.0005 

0.0000 

0.0000 0.0001 0.0002 0.0003 0.0004 0.0005 0.0006 0.0007 0.0008 0.0009 0.0010 

Liquid CO Conc. (mol/L) 

Figure 5 – Further Increases in Pressure Result in Unstable Operation 

Keeping in mind that the bioreactor reaction rate is based upon the mass of viable cells, 

increasing the cell mass would allow operation at higher pressures and higher mass 

transfer rates. Figure 6 shows the effect of doubling the cell mass. The mass transfer 

operating line is in the same position as in Figure 5, but the bioreactor rate for the given 

reactor has increased and the steady-state solution is again a stable point on the left side 

of the peak reaction rate. 

Total Pressure = 3 atm, Cell Mass = 2 g/L, kla = 0.05/sec 

0.0020 

0.0015 

0.0010 Reaction 

Transfer 

0.0005 

0.0000 

0.0000 0.0001 0.0002 0.0003 0.0004 0.0005 0.0006 0.0007 0.0008 0.0009 0.0010 

Liquid CO Conc. (mol/L) 

Figure 6 – Stable Operation at 3 atm Due to Increased Cell Mass Rate (mol/min) Rate (mol/min) 

10
A second option that would allow operation at higher pressures would be to decrease the 

kLa value of the reactor, reducing mass transfer. This could be accomplished by 

decreasing the area for mass transfer or reducing turbulence in the reactor. The net effect 

is shown in Figure 7. Comparing Figure 7 with Figure 5, the maximum mass transfer 

(the y-intercept) has decreased, but the equilibrium CO concentration (the x-intercept) 

remains the same. However, increasing the mass of cells in the reactor is a more 

desirable solution because the steady-state bioreactor rate in Figure 6 is almost twice that 

of Figure 7 for the same reactor volume. 

Total Pressure = 3 atm, Cell Mass = 1 g/L, kla = 0.025/sec 

0.0020 

0.0015 

0.0010 Reaction 

Transfer 

0.0005 

0.0000 

0.0000 0.0001 0.0002 0.0003 0.0004 0.0005 0.0006 0.0007 0.0008 0.0009 0.0010 

Liquid CO Conc. (mol/L) 

Figure 7 – Reduction in Mass Transfer Coefficient 

Increasing the cell mass in the reactors has not been demonstrated in the laboratory for 

the biological WGS reaction, but there would be a number of ways of increasing cell 

mass in the reactor, if desired. The net effect is that there are so many cells in the reactor 

that any CO transferred is consumed before it can build up to levels that cause inhibition 

of the biological WGS reaction. If for some reason the operating line moves to the right 

of the peak biological reaction rate, the mass transfer must then be reduced until stable 

operation is re-established to the left of the peak biological reaction rate. This can be 

accomplished by reducing the mass transfer coefficient kla or by decreasing the reactor 

operating pressure. 

To summarize the results of comparing the bioreactor rate with the mass transfer rate, 

increasing the pressure of the reactor will increase the mass transfer rate, while increasing 

the cell loading increases the bioreactor rate. For stable operation, the two operating lines 

must cross to the left of the peak reaction rate. Rate (mol/min) 

11
9.0 Reactor Designs 

Three different reactor designs were tested for biological WGS of CO to H2: a trickling 

filter, a bubble column and a bubble column with gas recycle. Batch shake flasks 

experiments were also conducted for determining intrinsic reaction rates. 

The trickling filter consists of a tubular reactor containing a solid packing support that 

cells can attach to (see Figure 10a). Various packings were tested, including wood, 

activated carbon, lava rock, plastic industrial packing and porous ceramic supports. The 

configuration tested in the lab was a counter-current design with the gas flowing upwards 

through the bed and water flowing down through the packed bed. The water flow rate is 

kept low to prevent flooding in the column. The main purpose of the water was to keep 

the cells moist with the secondary purpose of supplying nutrients for cell growth. While 

increased liquid flow rates might increase the mass transfer slightly, the main advantage 

of the design was increased gas-transfer area with a minimal pressure drop. Another 

possible configuration to increase mass transfer while still avoiding flooding would be a 

co-current down-flow reactor, but this configuration was not tested. 

Two reactor sizes of trickling filter reactors were tested: 1-liter and 5-liter. Figure 8 

shows typical conversions for different reactor packings in the 1-liter reactor (Wolfrum 

and Watt, 2001). Figure 9 compares the conversion of the 1-liter and 5-liter reactors 

under similar operating conditions. This data shows that a 5-fold increase in reactor 

volume with the same residence time yielded the same conversion, confirming that the 1-

liter data could be used to predict the performance of larger reactors. 

100% 

80% 

60% 

40% 

20% 

0% 

Hardwood Support 

Cellulose 

3mm Beads 

6mm Beads 

0.00 0.02 0.04 0.06 0.08 

G/Veb (min-1) 

Figure 8 – Typical Biological WGS Conversion Data % Conversion 

12
1.2 

1.0 

0.8 

0.6 

0.4 

0.2 

0.0 

0 50 100 150 200 250 300 

y = 0.0044x + 0.0102 

R2

 = 0.9839 

y = 0.0046x - 0.0332 

R2

 = 0.9676 

5-L Reactor 

1-L Reactor 

EBCT (min) 

Figure 9 – Trickling Filter Scale-Up 

Trickling filter reactors were operated under pressurized conditions. The higher 

pressures increased the effective the bulk gas CO concentration, increasing the driving 

force for mass transfer. However, at times the increased pressure would increase the CO 

liquid-phase concentration enough to inhibit the biological WGS reaction and shut down 

the reactor until the mass transfer rate was reduced. This result is expected based upon 

the shape of the rate curve in Figure 2. 

The other two reactor configurations tested were bubble column designs (Figures 10b and 

10c). Good mass transfer in a bubble column is achieved through high surface area for 

mass transfer and increased turbulence as the bubbles rise through the column of liquid. 

Ideally, a bubble column should be designed to produce small bubbles, which have high 

surface-to-volume ratios and low rise-velocities though the liquid, resulting in long 

residence times in the reactor. Large bubbles have less area for mass transfer and rise 

very quickly up through the reactor liquid. Producing small bubbles, however, requires 

pushing the gaseous feed through small opening, such as those in a porous frit. This 

means there is a high pressure drop across the reactor. The liquid head in the reactor adds 

an additional pressure drop, which is not present in a trickling filter. 

The short residence time and high pressure drop in the bubble reactor make it impractical 

for commercial-scale biological WGS, but there is a advantage in obtaining reactor 

design information: it is easy to estimate the amount of cell biomass in the reactor. Since 

the reactor is well-mixed, a small amount of liquid can be tested to determine the total 

reactor cell density. In the trickling filters, it was more difficult to determine the amount 

of cellular material attached to the packing without disassembling the reactor. ln(1-%Conv) 

13
Stand-alone bubble columns produce very small single-pass conversions of just a few 

percent for a 1-meter column, making it difficult to accurately calculate conversion rates 

for the reactors. To increase the accuracy of the calculations, but still use the bubble 

column, a gas recycle loop was added. This allowed the simulation of a longer residence 

time in the reactor, resulting in overall conversions of over 50% in reactors with 10:1 

recycle-to-feed ratios. The bubble columns ran more stable than the trickling filters, 

presumably due to the higher cell mass, more liquid available to buffer upsets and lower 

inlet CO concentrations to the reactor. 

Figures 10a, 10b, 10c show the three reactor designs. 

Figure 10 – Three Biological WGS Reactor Designs 

Some factors affecting operation of the reactors are operating pressure, inlet CO 

concentration, recycle-to-feed ratio and cell concentration. Both the trickling filter and 

recycle bubble column reactors were operated under pressure. The recycling bubble 

column reactor was also operated with variable recycle ratios. Gradual increases in the 

recycle-to-feed ratio initially resulted in higher mass transfer rates as the volume/number 

of bubbles increased, but at very high recycle rates, the bubbles combine and the surfaceto-volume ratio drops, hurting mass transfer. 

10.0 Process Economics 

In the course of studying the biological WGS process, several alternative process 

configurations were studied. Several of the processes concentrated on biomass-derived 

synthesis gas, but any stream containing CO would be a possible feed to the biological 

WGS process. 

14
10.1 Biomass-Derived Synthesis Gas with a Trickle-Bed Reactor 

In 2001, a study was done to look at the economics of the biological WGS process using 

a quenched gas leaving a FERCO/BCL indirectly-heated gasifier (Amos, 2001). The 

assumed gas composition was an equal mixture of CO, H2, and CO2. The process layout 

consisted of compression of the cool gas, followed by a water quench before the gas 

entered a counter-current packed-bed trickling-filter biological WGS reactor. The shifted 

gas would then go through a MEA CO2 scrubber before passing through drying beds to 

remove any residual water. A 2.5 MM scfd plant had a capital cost of $17 million and 

resulted in an incremental processing cost $3.40/kg. Adding in the cost of the syngas at 

$5/MM Btu gave a minimum hydrogen selling price of $4.25/kg. This is compared to 

$1.00-1.50/kg for hydrogen via catalytic steam-methane reforming. 

Sensitivity studies conducted included looking at the effects of reaction rate, outlet CO 

concentration, packing cost, labor required and non-pressurized operation. One 

important conclusion of this work was that the trickling filter needed to operate under 

pressure. If the system operated at ambient pressures, the large reactor size would result 

in a minimum hydrogen selling price of over $30/kg. 

10.2 Industrial Collaboration 

NREL conducted a series of laboratory studies and economic analysis through a WorkFor-Others agreement with an industrial partner interested in refinery applications of the 

biological WGS process. One aspect of these studies was the removal and/or recovery of 

CO2 from the process outline in Section 9.1. 

Reducing reactor size is important to the economics of the biological WGS process. One 

way to reduce the size of the biological WGS reactor is to remove the CO2 from the inlet 

gas before it enters the reactor. Because the inlet gas from a FERCO/BCL gasifier can 

contain 33% CO2, this could result in a 33% reduction in reactor volume. The CO2 

resulting from the shift reaction can also be removed from the gas steam after the reactor, 

improving downstream hydrogen purification steps. 

A significant amount of time was spent studying the options for CO2 recovery. The 

process design for biomass-derived gas included MEA scrubbing. One issue with this 

method of CO2 removal is the high natural gas consumption associated with the 

subsequent CO2 stripping column used to regenerate the MEA. The natural gas costs 

significantly added to the operating expenses. To help reduce this cost, an alternative 

stripper design used air stripping to regenerate the MEA and not recover the CO2. Even 

with the lost CO2 revenues, this was a cheaper alternative to consuming natural gas and 

recovering the CO2. 

The next CO2 recovery option examined was to compress and freeze-out the CO2 in the 

inlet stream. The gas was compressed and cooled using a refrigeration system. Solid 

CO2 would deposit on the inside of a set of heat exchangers. The CO2 would then be 

periodically vaporized by passing hot water through the heat exchanger. This would 

allow capture of gaseous CO2, but liquid CO2 brings in considerably higher revenues. 

15
An alternative process consisting of more compression and cryogenic condensers was 

examined. While a stripping column is required for purifying the liquid CO2 before sale, 

low-quality waste heat from the compression intercoolers is sufficient for providing heat 

to the stripping column reboiler. Condensing the CO2 also required operation at higher 

pressures, which increased compression costs. 

The results from the economic studies indicated that the overall process economics 

strongly depended upon how much the CO2 could be sold for. Liquid, food-grade CO2 

can be sold for up to $50/ton in some areas, but the sales are limited by region and are 

seasonal in nature. The CO2 removal did, however, improve the economics of the 

process and reduce the reactor size for the biological WGS. 

10.3 Drop-in replacement in SMR 

One early argument in support of the biological WGS process was as a drop-in 

replacement for conventional high- and low-temperature WGS reactors. However, the 

biological WGS reactor would be significantly larger than a comparable catalytic reactor. 

The larger reactor volume for the biological WGS doesn’t reduce capital or operating 

costs for SMR. 

One argument for biological WGS was that the lower operating temperature (25°C) 

would result in high conversions of CO to H2, possibly eliminating the need for PSA 

purification. While it is true that the chemical equilibrium is shifted to producing almost 

all H2 with little CO at 25°C, the production rate of a full-scale biological WGS reactor 

will be limited by mass transfer, not equilibrium. Mass transfer from the gas phase to the 

liquid phase—where the biological WGS reaction takes place—is a first-order reaction. 

That means for every order of magnitude increase in conversion, the reactor volume 

doubles. To increase the conversion of a reactor achieving 90% conversion to 99% 

conversion, the reactor volume would need to be doubled. To achieve 99.9% conversion, 

the reactor volume would need to be doubled yet again to 4 times the original volume. 

While it is theoretically possible to achieve CO concentrations in the ppm range, the large 

reactor size wouldn’t justify going to such high conversions. A far better method is to get 

a conversion of 90% in the biological WGS, then use a PSA to remove the additional CO. 

Since the CO2 and other possible trace contaminants must be removed to produce pure 

H2, a PSA would be required anyway. 

There were also questions concerning whether there were any heat transfer benefits to 

using a biological WGS reactor operating at 25°C. In the conventional reforming, HTS, 

LTS, PSA design, heat from the reformer is used to preheat the inlet gas to the reformer, 

then steam is generated from the outlet to the HTS. A relatively large amount of lowquality heat can be recovered from the outlet of the LTS reactor—mostly from 

condensing the steam in the gas stream—but since the LTS is operating at 200°C, this is 

not high-quality steam. Some cooling water is also needed to condense as much water as 

possible before entering drying beds prior to the PSA purification. The net effect is 

recovering heat from 350°C down to 25°C before the PSA in multiple steps. With the 

biological WGS process, the gas is still cooled down from 350°C to 25°C before entering 

the biological WGS reactor, but it is done in one step. Because the conventional catalytic 

16
WGS reaction is exothermic, some extra heat is produced and recovered leaving the shift 

reactors. The biological WGS uses some of the exothermic energy of the shift reaction 

for metabolic processes and the remaining heat is lost at ambient temperatures and can’t 

be recovered. The biological WGS reactor actually has worse heat recovery because the 

exothermic heat of reaction of the water-gas shift reaction is given off at 25°C. 

One cost analysis looked at exactly how much of the commercial H2 price is due to the 

capital and catalyst costs associated with the shift reactors. The result was that 

completely eliminating the capital and catalyst replacement costs of the conventional 

catalytic WGS only reduced the hydrogen selling price from a modern SMR plant by 

$0.10/kg of hydrogen. The biological WGS alternative would cost considerably more. 

10.4 Non-reforming applications 

While the costs associated with the conventional catalytic WGS reactor are small, larger 

savings are possible if the reforming and steam production associated with hydrogen 

production can be eliminated. As was mentioned earlier, one difference between 

catalytic and biological water-gas shift is that the water used in the reaction is in liquid 

form for the biological process; for catalytic WGS, high-pressure steam is used. There 

are both capital and operating costs associated with producing this steam. The reformer 

in SMR hydrogen production also represents a significant capital investment for highpressure, high-temperature reactor tubes. If the reformer and the steam production can be 

eliminated, this might make a low-temperature biological WGS reactor attractive. 

Amos et al (2003) conducted a series of process simulations using different feed 

compositions and four different combinations of reforming, HTS, LTS and biological 

WGS. The four process configurations were: 

1. Conventional SMR with natural gas 

2. Reforming, catalytic HTS, PSA 

3. No reforming, catalytic HTS, PSA 

4. No reforming, biological WGS, PSA 

Figure 11 shows these four process configurations. No CO2 recovery was attempted in 

these designs. For conventional SMR, natural gas at $3.50/MM Btu was assumed. For 

the remaining cases, equal concentrations of CO, CO2 and H2 were assumed with 0-25% 

methane. The assumed cost of the synthesis gas was $5.00 MM/Btu. 

17
Case #1: 

Natural 

Gas HTS PSA 

Steam 

Reformer H2 

Case #2: 

Syngas Reformer HTS 

Steam 

PSA H2

Case #3: 

Syngas HTS 

Steam 

PSA H2

Case #4: 

Syngas Bio 

WGS 

Water 

PSA H2

Figure 11 – Alternative Reformer/Shift Processes 

Figure 12 compares the alternative reforming and shift strategies with conventional SMR 

(Case 1). As suspected, at low methane/hydrocarbon concentrations, eliminating the 

reforming and steam production resulted in lower hydrogen costs with a pressurized 

biological WGS process. However, if the methane concentration was over 5%, installing 

the equipment for steam production and reforming resulted in enough of an increase in 

hydrogen yield from the reformed methane to justify the higher capital and operating 

expenses with the reformer and steam production. 

$4 

$3 

$2 

$1 

$0 

0% 5% 10% 15% 20% 25% 

Case #1 Case #2 

Case #3 Case #4 

Methane Concentration (%) 

Figure 12 – Minimum Hydrogen Selling Price for Various Syngas Compositions H2 Selling Price ($/kg) 

18
It is important to note that if one decides to steam reform the inlet stream to boost 

hydrogen yields, the stream leaving the reformer already contains sufficient steam for the 

WGS reaction and is already hot enough for the HTS reactor, so no additional steam and 

no additional heat must be supplied. Because of this, it makes sense to go directly into a 

conventional catalytic reactor, rather than cooling the gas to go into a significantly larger 

and slower biological WGS reactor. The ideal stream for biological WGS would be a 

pressurized CO-containing stream with little or no hydrocarbons that could be reformed 

into CO2 and hydrogen. 

10.5 Interface to a Gasifier 

With respect to coupling a WGS process to a gasifier, for any low-pressure gasifier, in 

order to minimize the size of the reformer and/or shift reactors, the gas would need to be 

pressurized. In order to pressurize the gas in a cost-effective manner, the gas must be 

cooled and any steam condensed, before compressing the gas. That means the gas from 

any low-pressure gasifier will pass through a low-temperature compression step—what 

the gasifier and heat integration looks like before that step is irrelevant. Whether a 

catalytic or biological WGS is used, the gasifier, quench/heat recovery and compression 

sections will most likely be designed the same. 

For a high-pressure gasifier, no compression would be needed—just some variation of 

hot-gas cleanup and possibly pre-reforming or tar cracking at the gasifier outlet 

temperature. In any case, if the gas leaves the gasifier at a high pressure, at a high 

temperature and with steam present, it is the perfect stream to pass into a catalytic 

reformer and/or catalytic shift reactor. Like in the SMR case, it makes no sense to cool 

the gas stream to go through a larger, slower biological WGS reactor if the stream is 

already hot and contains steam. 

10.6 HyCO Plants 

In July 2003, the author toured an Air Liquide HyCO plant. This plant was designed to 

supply CO to a chemical manufacturing site. The CO was the primary product, but some 

of the H2 produced by methane reforming was purified using a PSA and sold into a 

hydrogen pipeline onsite. Because CO was the primary product, there was no need for a 

WGS reactor. In fact, CO2 was recycled back to the reformer to boost CO production. 

However, there was one operational concern with the plant. At times, the CO demand 

would drop, but not completely disappear. The reformer could not operate stably under 

60% capacity. If the CO demand was less than 60% of the plant capacity, the reformer 

must still operate at 60% of capacity and the plant would flare the extra CO. It would not 

make sense to install a catalytic shift reactor for occasional use because it would require a 

lengthy heat up time to reach the proper operating temperature. 

A biological WGS reactor would fit into the process because the pure CO stream is 

pressurized to 120 psig and the shifted H2-containing stream could be fed into the system 

ahead of the existing CO2 scrubber. The CO2 produced could be removed and then the 

hydrogen would be sent through the existing PSA and sold to the pipeline merchant. 

19
When CO demand was over 60%, only a small slipstream would be required to maintain 

the health of the biological WGS reactor. Because the biological WGS reactor operates 

at ambient temperatures, no warm-up would be required. In this HyCO plant design, the 

biological WGS reactor would be the only piece of equipment that would need to be 

installed; all compression, scrubbing and purification equipment is already in place. 

The applicability of this technology to other HyCO plants is unknown. 

11.0 Conclusions 

Based upon the laboratory research completed and the various economic analyses 

mentioned in Section 9, several conclusions can be made about the suitability of the 

biological WGS reaction for hydrogen production. 

1). A biological WGS will always be larger than catalytic WGS due to the lower mass 

transfer and slower kinetics as compared to a catalytic WGS reactor. For this reason, 

using a biological WGS reactor as a drop-in replacement for a WGS reactor won’t 

improve the process economics. 

2). There are not significant heat integration advantages with using a biological WGS 

reactor in place of a conventional catalytic WGS reactor. It would be possible to recover 

approximately the same amount of heat from both systems—the conventional system 

uses multiple steps to recover the heat, while the heat recovery for the biological WGS 

could be done in one step. 

3). If steam reforming is already included as a process step, it makes sense to pass the 

outlet from the reforming to a catalytic WGS reactor because the gas is already hot and 

already contains enough steam for the WGS reaction. 

4). If no steam reforming is required, the capital cost reduction associated with 

eliminating steam production might justify the larger biological WGS reactor. However, 

if significant amounts of reformable hydrocarbons are present (i.e., greater than 5% by 

volume), adding a reforming step to increase hydrogen yields improves the economics. 

5). Pressurized operation of the biological WGS reactor would be a requirement to 

improve mass transfer and reduce reactor size. However, due to CO inhibition, operating 

at higher pressures requires higher cell mass in the biological WGS reactor. Successful 

operation at pressures above 3 atm with elevated cell mass in a biological WGS reactor 

has not been demonstrated. 

6). Recovery of carbon dioxide as a co-product can have a major effect on the process 

economics, depending upon the CO2 recovery process used and the market value of the 

CO2. This is true of both catalytic and biological WGS processes and depends upon the 

inlet gas composition. 

20
7). The best application for the biological WGS process is conversion of a pressurized 

CO-containing stream low in hydrocarbons to hydrogen with CO2 recovery. At least one 

such commercial site has been identified where biological WGS might be applicable, but 

no economic analysis has been performed. 

12.0 References 

Amos, W.A. (October 2001). Economic Analysis of the Biological Water-Gas Shift 

Process for the Production of Hydrogen from Synthesis Gas. Milestone report. 

Available from National Renewable Energy Laboratory, Golden, CO. 

Amos, W.A., Wolfrum, E.J., Watt, A.S. (2003). “Biological H2 Production from 

Synthesis Gas: Preliminary Techno-Economics & Reactor Design Issues.” 25th 

Symposium on Biotechnology for Fuels and Chemicals; May4-7, 2003; 

Breckenridge, Colorado. 

Gossett, J.M. (1995). Bioenergetics and Stoichiometry. Course notes. CEE756 – 

Environmental Engineering Processes II. Ithaca, NY: Cornell University. 

Kirk-Othmer Encyclopedia of Chemical Technology, 4th ed. (1995). Vol. 13. New York: 

John Wiley & Sons; pp. 838-894. 

McCarty, P.L. (1971). “Energetics and Bacterial Growth.” Chpaer 21 in Organic 

Compounds in Aquatic Environments. New York: Marcel Dekker, Inc.; pp. 495-

531. 

McCarty, P.L. (1972). “Energetics of Organic Matter Degredation.” Chapter 5 in Water 

Pollution Microbiology. New York: John Wiley & Sons, Inc. 

Ullmann’s Encyclopedia of Industrial Chemistry, 5th ed. (1989). Vol. A13. New York: 

VCH Publishing; pp. 297-442. 

Wolfrum, E. J., Maness, P. (2003). ”Biological Water Gas Shift.” U.S. DOE Hydrogen, 

Fuel Cell and Infrastructure Technologies Program Review; May 19-22, 2003; 

Berkeley, California. 

Wolfrum, E.J., Watt, A.S. (2001). ”Bioreactor Design Studies for a Novel HydrogenProducing Bacterium.” Proceedings of the 2001 U.S. DOE Hydrogen Program 

Review. NREL/CP-570-30535. Golden, CO: National Renewable Energy 

Laboratory. 

Wolfrum, E.J., Watt, A.S., Huang, J. (2002). “Bioreactor Development for Biological 

Hydrogen Production.” Proceedings of the 2002 U.S. DOE Hydrogen Program 

Review. NREL/CP-610-32405. Golden, CO: National Renewable Energy 

Laboratory. 

21
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Biological Water-Gas Shift Conversion of Carbon Monoxide to Hydrogen: Milestone 

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Wade A. Amos 

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This report summarizes the results of research and economic analysis on a biological water-gas shift process for the 

production of hydrogen. The organism Rubrivivax gelatinosus CBS is a photosynthetic bacteria which can perform the watergas shift reaction under anaerobic conditions. The report describes some of the technical issues regarding the process, 

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