Bioresource Technology 98 (2007) 2641–2651
Two-phase anaerobic digestion for productionof hydrogen–methane mixtures
Michael Cooney a,b,*, Nathan Maynard a, Christopher Cannizzaro b, John Benemann c
a Hawaii Natural Energy Institute, University of Hawaii, 1680 East–West Road POST 109, Honolulu HI 96822, United Statesb Center for Space Research, Massachusetts Institute of Technology, 77 Massachusetts Avenue, Cambridge MA 02139, United States
c Benemann Associates, 3434 Tice Creek Dr. No.1, Walnut Creek, CA 94595, United States
Received in revised form 11 September 2006; accepted 16 September 2006Available online 13 December 2006
Abstract
An anaerobic digestion process to produce hydrogen and methane in two sequential stages was investigated, using two bioreactors of2 and 15 L working volume, respectively. This relative volume ratio (and shorter retention time in the second, CH4-producing reactor)was selected, in part, to test the assumption that separation of phase can enhance metabolism in the second methane producing reactor.The reactor system was seeded with conventional anaerobic digester sludge, fed with a glucose–yeast extract – peptone medium and oper-ated under conditions of relatively low mixing, to simulate full scale operation. A total of nine steady states were investigated, spanning arange of feed concentrations, dilution rates, feed carbon to nitrogen ratios and degree of integration of the two stages. The performanceof this two-stage process and potential practical applications for the production of clean-burning hydrogen–methane mixtures arediscussed.� 2006 Elsevier Ltd. All rights reserved.
Keywords: Biological hydrogen; Methane; Anaerobic digestion; Renewable energy; Acetogen; Methanogen
1. Introduction
Biological hydrogen fuel production is a challengingarea of biotechnology. Photobiological processes, by whichmicroalgae or photosynthetic bacteria produce H2 fromsunlight and either water or organic substrates, respec-tively, have been studied for several decades, but arelimited by many practical and fundamental limitations(Benemann, 1997; Levin et al., 2004; Nath and Das,2004). Dark bacterial hydrogen fermentations or organicsubstrates also have limitations, principally the relativelylow yields of hydrogen obtained until now, with typicallyonly 10–20%, and at most 30%, of the substrate energy
0960-8524/$ - see front matter � 2006 Elsevier Ltd. All rights reserved.
doi:10.1016/j.biortech.2006.09.054
* Corresponding author. Present address: Hawaii Natural EnergyInstitute, University of Hawaii, 1680 East–West Road POST 109,Honolulu HI 96822, United States. Tel.: +1 808 956 7337; fax: +1 808956 2336.
E-mail address: [email protected] (M. Cooney).
converted to H2 fuel (Das and Veziroglu, 2001; Van Ginkeland Logan, 2005). This compares to 80–90% yieldsobtained in commercial ethanol or methane fermentations(Claassen et al., 1999). Higher yields may be achievable, inprinciple, through metabolic engineering (Hallenback andBenemann, 2002; Keasling et al., 1998) but, thus far, nomajor improvements in yields have been reported.
Methane fermentations, also called anaerobic digestion,involve consortia of two major types of bacteria: the so-called acidogenic bacteria that break down the substratesinto mainly H2, acetic acid and CO2, and the methanogenicbacteria, that convert acetic acid, H2 and CO2 to methanegas. A variety of higher organic acids, such as propionic,butyric and lactic, as well as alcohols and ketones, are alsoformed during the breakdown of the organic substrates bythe highly diverse populations generically known as acido-gens. However, in a well operating process, these productsare mostly converted to acetic acid and H2, which, in turn,are then converted to methane gas. The key to this process
2642 M. Cooney et al. / Bioresource Technology 98 (2007) 2641–2651
is that the H2 produced by the acetogenic bacteria isremoved to a very low partial pressure, typically in thenanomolar range, by the methanogenic bacteria, whichallows the otherwise thermodynamically unfavorablemetabolism of the higher organic acids and alcohols to ace-tic acid and H2 (Belaich et al., 1990). The result of thiscomensal relationship is an overall high yield conversionof fermentable substrates to methane fuel with at mosttrace amounts of H2 present in the gas phase (Ferris,1993). Indeed, even a small amount of H2 (>0.1%) in thegas phase indicates a malfunctioning process, due to over-load, toxicity, or other factors unbalancing the comensalrelationships, generally followed by a cessation of CH4-production. However, the two reactions, formation oforganic acids and H2 and methane-production, can, at leastpartially be separated into separate bioreactors in series, inwhich the first, smaller, reactor produces organic acids, H2
and CO2, while the second, much larger, reactor producesCH4 and CO2. Such two-phase anaerobic digestions wereproposed as a way to optimize for the growth of each typeof bacteria in the separate reactors, specifically by growingthe acetogenic bacteria at a lower pH (e.g., 5–6) and shorthydraulic residence time (typically 1–2 days) in the firststage, while the slower growing methanogenic bacteriastage, requiring a more neutral pH, were preferentially cul-tured in the second stage with a much longer hydraulic res-idence time (typically 10–20 days), (Blonskaja et al., 2003;Demirel and Yenigun, 2002; Pohland and Ghosh, 1971).However, as noted above, anaerobic digestion involves acomensal interaction of the two general types of bacterialpopulations, in which the methanogens feed on and effi-ciently remove, the waste products (H2 and acetic acid)of the acidogenic bacteria. Thus separating these two basicprocesses will not generally significantly accelerate orincrease overall methane-production, although it can beof some advantage in making the process more resistantto shock loading.
In earlier work on two-phase anaerobic digestion, thegas, H2 and CO2, produced in the first stage was trans-ferred to the second stage to be converted to CH4, and,indeed, few measurements on gas production from the firstphase are reported in the literature. However, due to theburgeoning interest in H2 fuels and fuel cells, there hasbeen a great deal of research in recent years on dark H2 fer-mentations (Das and Veziroglu, 2001; Hallenback andBenemann, 2002; Levin et al., 2004) which, essentially, cor-respond to the first phase of such a two-phase anaerobicdigestion process. However, as stated above, in all suchstudies, the overall H2 yields are low, only 10–20% of thesubstrate energy being converted to H2 fuel with theremainder converted to organic acids, and other products.A number of authors have proposed converting these by-products to H2 fuel using photosynthetic bacteria, whichcan exhibit high yields, though at very low solar conversionefficiencies (Miyake, 1998), which makes such an approachimpractical. Producing methane gas in a second reactor hasbeen proposed as another option (Benemann, 1998), and
some work on such dual H2 and CH4-production hasappeared recently (Benemann et al., 2004; Kramer andBagley, 2004). However, if H2-production is the objective,it is more direct and plausible to produce methane gas bynormal anaerobic digestion and then convert this fuel toH2 through a conventional reformer process.
A potential near-term practical application of two-phaseanaerobic digestion is the production of H2–CH4 methanemixtures. H2–CH4 mixtures, in the range of 10–30% H2
and 90–70% CH4, on a volumetric basis, are known toburn with much lower NOx emission in internal combus-tion engines, and this allows the use of such fuels in regionswhere NOx emissions are strictly regulated (Bauer and For-est, 2001; Collier et al., 1996). Here we address the co-pro-duction of H2 and CH4 mixtures in a two phase anaerobicdigestion process, using a simulated high carbohydratewastewater with a mixed bacterial population obtainedfrom a conventional anaerobic digester under operatingconditions designed to simulate a scaled-up process.
2. Methods
2.1. Two-stage anaerobic bioreactor system
The two-stage anaerobic bioreactor system (Fig. 1)consisted of a 4.7 L polycarbonate jar with cover(DS5300-9609, Nalgene, Rochester, NY) for the first(‘‘hydrogen-production,’’) reactor and a 18.8 L polycar-bonate jar and cover (DS5300-9212, Nalgene) for the sec-ond (‘‘methane-production’’) reactor. The workingvolumes were 2.0 and 15.0 L, respectively. All connectionswere made with Teflon tubing (890 FEP, Nalgene) andstainless steel or nylon compression fittings (SwagelokCo., Solon, Ohio). The lids for the reactors were sealedby compression against an o’ring using two steel platesplaced above and below the reactor and bolted in placealong four corners. The assembled reactors were indepen-dently tested for gas leakage by introducing nitrogen gasvia a gas sparging port and with an exhaust tube placedat the bottom of a graduated cylinder filled with 45.0 cmof water (corresponding to a pressure head of about0.05 bar or 0.67 psi). After the nitrogen flow was stoppedthe gas level in the submersed exhaust tube remained con-stant for at least 1 h, indicating no significant gas leakage.
The reactors were mixed by placing each reactor ona magnetic stirrer (PC-310, Corning Inc., Corning, NY)and stirring with 1.5 in. magnetic stir bars. When thehydrogen and methane-production reactors were operatedindependently (‘‘non-integrated operation’’), the effluentfrom each reactor was peristaltically pumped to separatewaste bottles. When operated as a two-phase process(‘‘integrated operation’’), the effluent from the first reactorwas pumped into the second reactor. In both cases, peri-staltic pumping of reactor media was through a tube thatwas placed below the surface of the reactor liquid, to avoidthe exchange of gases between the two reactors, foam frac-tionation and related artifacts. The liquid height and thus
L N2
Cryostat
T
pH
T
pH
P
P
Feed Tank
L
Base
Acid
The relative fraction of feed media or effluentdelivered to the second CH4-productionreactor is described in Table 2.
H2 Rxr
CH4 Rxr
Fig. 1. Two-stage anaerobic reactor system design. The first stage is referred to as the H2-production reactor while the second stage is referred to as theCH4-production reactor.
M. Cooney et al. / Bioresource Technology 98 (2007) 2641–2651 2643
reactor volume in both reactors was thus controlled by aconductivity sensor which activated a peristaltic effluentpump when the liquid reached the tip the of the sensor.In all experiments, in the first (hydrogen-production) reac-tors the conductivity sensor was set to a height that main-tained a working volume of 2.0 L while in the second(methane-production) reactors the height was set to aworking volume of 15.0 L.
The feed medium was held in a polyethylene tank with acapacity of 114.0 L, maintained at 5 �C with an internalaluminum coil heat exchanger (EX11, Aquatic Eco-systemsInc., Apopka, FL), connected to an external, setpoint con-trolled cryostat (Ultratemp, 2000, Julabo Labortecknic,Germany) and mixed with a magnetic stir bar. Mixing inboth reactors was deliberately low to better reflect condi-tions of industrial-scale systems, where high mixing ratesnormally used in experimental bioreactors are not applica-ble. The feed tubings were replaced daily or every 2 days, asneeded. Nitrogen gas was bubbled continuously throughthe medium to help maintain an anaerobic environment.The feed media was delivered to the reactors by peristalticpumping through Teflon tubing (101U/R, Watson–Mar-low Ltd., Cornwall, England). The pumps were calibratedand their feed rates periodically verified gravimetrically.
Temperature and pH in both reactors were controlledwith a MicroDCU unit (B. Braun Biotech Inc., Allentown,PA) and continuously logged through a RS232 serial inter-face connected to an external Dell 4100 computer (Dell,Texas USA). Temperature was measured in situ using PT-
100 probes and the pH with gel-based probes (Mettler-Toledo, Greifensee, Switzerland) and maintained at35 �C ± 0.1 �C in both reactors using 200 W and 450 Wcartridge heaters (McMaster-Carr, Los Angeles, CA), inthe first and second reactors, respectively. pH was con-trolled at 5.5 in the first reactor and at 7.0 in the secondreactor through automated addition of sodium hydroxide.The concentration of base was 2.0 M, but was increased to4.0 and 6.0 M during the experiments using the highfeed concentrations. The total base consumed was calcu-lated from the recorded ‘‘on-time’’ of the base pump andthe base consumption rate was calculated by leastsquares regression of the data with respect to time over a60 min s interval.
Data was calculated in 1 or 2 min intervals and loggedusing Lab View. For each steady state, rates (gas evolution,base addition) were averaged over a 24 h period that dem-onstrated consistent data and did not exhibit mechanical orother interruptions (foaming, clogged lines, lack of mixing,or power outages). Steady state was assumed to have beenreached after at least two full residence times (i.e. reactorsvolumes) had passed and the gas evolution rate (estimatedover a 20–24 h period) had steadied to the same value over3 days.
2.2. Media, inoculum, and start-up
The standard media used for the batch and continuousexperiments consisted of (per liter): 10.0 g glucose, 1.5 g
2644 M. Cooney et al. / Bioresource Technology 98 (2007) 2641–2651
KH2PO4, 1.67 Na2HPO4, 0.5 NH4Cl, 0.18 MgCl Æ 6H2O,2.0 g yeast extract, 2.0 g peptone, 0.02 ml antifoam A(Sigma). In some experiments this media was modified byincreasing the feed glucose concentration or by doublingthe concentration of peptone and yeast extract per liter(termed 2N as opposed to 1N for the standard medium)as described. The carbon to nitrogen ratio calculated foreach media formulation was calculated by summing thegrams of carbon and nitrogen (per liter) provided by theglucose, ammonium chloride, peptone, and yeast extract.The contribution of nitrogen and carbon (per liter) frompeptone (15.4% N, 31.5% C) and yeast extract (18% Nand 32% C) was estimated by summing the contributionsof carbon and nitrogen from the components (i.e., aminoacids, casein, etc.) in each as per the compositional analysisprovided by the manufacturers. The C:N mol mol�1 ratioof the standard medium was 7.59.
Anaerobic digester sludge was collected from the HawaiiKai Waste Water Treatment Plant (Hawaii Kai, HI) andstored at 4 �C until 5.0 ml of the well-mixed sludge wasused to innoculate both reactors. No pre-treatment (e.g.,heating to help select for spore-forming anaerobes) wasapplied, to better reflect larger-scale processes where self-selecting bacterial populations would likely dominate aftera relatively short initial period. To establish the methano-genic culture, the second reactor was initially operatedindependently of the first, by directly feeding it standardmedia at increasingly greater flow rates. Once a methanevolume percentage of 50% was achieved, effluent fromthe first hydrogen-production reactor was used to slowlyreplace the second reactor media feed, eventually achievinga fully integrated operation of a two-phase anaerobicdigestion process.
For the batch cultures, each reactor was drained to onefourth its volume and fresh standard media was added,over a 5 min period. The reactors were briefly deaeratedwith N2 gas before being sealed and liquid and gas phasemeasurements initiated, as in the continuous cultivationexperiments.
2.3. Gas evolution rate measurements
The gas production rate from each reactor was mea-sured on-line using high precision gas meters built in-houseconsisting of a pressure transducer, a three-way solenoidvalve, a ballast chamber and custom built circuit boards.The gas meters were calibrated with a peristaltic pump thatdelivered a precise volume of gas over a defined time inter-val measured with a digital flowmeter (model 520, FisherScientific) which gave a linear response for valve actuationrate vs. gas flow rate. The calibrated gas meters were con-nected to the head space of the reactor. The gas meterresponse was not perturbed with regard to the action ofthe liquid level control, confirming the accuracy of thegas flow measurements and the general gas-tightness ofthe bioreactors. Leakages from the first and second reac-tors over a 44 h period was found to be only 2.1 ml h�1
and 3.7 ml h�1, respectively, which did not impact theresults. Repeated calibration curves over the 6-monthperiod of the experiments were stable, reliable and linearupto 600.0 ml h�1. A full description of the gas meter isdescribed elsewhere (Cooney et al., 2006).
2.4. Gas- and liquid-phase metabolite analysis
The gas composition of the reactor headspace wasmeasured using an Agilent Technologies 6890 gas chro-matograph equipped with two columns separated by aswitching valve as designed by the manufacturer. The firstcolumn was a Plot Q polymer column (19091P-Q04, Agi-lent Technologies), to separate CO2 and higher molecularweight compounds, and the second a molecular sieve col-umn (19091P-MS8, Agilent Technologies) to separate thelower molecular weight gases (H2, O2, N2, CH4). Calibra-tion curves generated for these five components were linearand reproducible.
Liquid phase volatile organic acids (VOAs) were ana-lyzed by HPLC analysis (1100 series, Agilent Technolo-gies). Separation was achieved using a 0.1% H2SO4
mobile phase pumped at a rate of 0.8 ml min�1 through aSupelcogel C-610H ion exchange column maintained at30 �C. The injection volume of sample was 10.0 lL. Detec-tion was accomplished using an RI detector maintained at30 �C with a Timberline TL-105 column heater (TimberlineInstruments, Boulder, CO). To identify the retention timesof unknown peaks, over 30 compounds were screened.Based upon these results, a calibration mixture was madeup with eleven metabolites normally present at significantlevels in samples from the two bioreactors.
2.5. Carbon mass balance
A carbon balance was calculated around each reactor(when run independently) by comparing the carbon enter-ing the reactor (i.e., feed glucose, yeast extract, andpeptone) against the carbon leaving the bioreactor(i.e., liquid and gas phase metabolites, biomass, and uncon-sumed peptone and yeast extract). The carbon entering thereactor was calculated using the MW of glucose and anestimated value of 1.27 g C L�1 for YE and peptone addedat 2.0 g yeast extract and 2.0 g peptone per liter (e.g., man-ufacturers information on carbon content of YE and fromthe amino acid(s) composition of peptone). The carbon inthe liquid phase VOAs was calculated from the HPLCmeasurements. The carbon in the dissolved carbon dioxideand methane was calculated using Henry’s law coefficientsof 29.76 and 714.0 atm�1 mol�1, respectively, and by tak-ing into account the carbon dioxide held in the form of car-bonate and bicarbonate (Blanch and Clark, 1997). Thecarbon in the biomass leaving the bioreactor was estimatedby assuming a yield coefficient of 0.2 gram dry weight pergram glucose consumed, and a biomass carbon composi-tion of 50%, considered reasonable as much of the biomasswould be derived from the input yeast and peptone
Tab
le1
Ind
epen
den
t(n
on
-in
tegr
ated
)o
per
atio
no
fth
ere
acto
rsat
vary
ing
dil
uti
on
rate
s
No
n-i
nte
grat
ed
H2R
xrF
eed
(Ld�
1)
D (d�
1)
H2P
R
(mm
ol
L�
1d�
1)
YH
2/G
lu
(mm
ol
mm
ol�
1)
To
tal
gas
(ml
h�
1)
Bas
e
(ml
h�
1)
H2
(%)
CH
4
(%)
Lac
tate
(gL�
1)
Fo
rmat
e
(gL�
1)
Ace
tate
(gL�
1)
Ace
toin
(gL�
1)
Eth
ano
l
(gL�
1)
Bu
tyra
te
(gL�
1)
SS
12.
01.
06.
62±
0.06
0.11
±0.
037
.47
±0.
283.
735
.32
±0.
080.
327.
53±
0.28
0.37
±0.
020.
00±
0.00
0.27
±0.
060.
34±
0.01
0.22
±0.
02
SS
23.
01.
59.
60±
0.16
0.12
±0.
055
.0±
1.34
7.5
35.7
9±
0.64
0.59
8.47
±0.
850.
40±
0.04
0.26
±0.
030.
08±
0.01
0.33
±0.
00.
25±
0.05
SS
34.
02.
015
.30
±1.
180.
16±
0.01
79.8
5±
1.35
5.0
41.1
3±
2.76
0.62
7.22
±0.
130.
47±
0.0
0.26
±0.
020.
49±
0.09
0.24
±0.
030.
33±
0.04
SS
45.
02.
523
.01
±0.
120.
16±
0.0
138.
68±
0.23
4.6
34.9
6±
0.12
0.51
6.78
±0.
620.
43±
0.06
0.19
±0.
030.
23±
0.11
0.38
±0.
130.
75±
0.06
SS
56.
03.
06.
61±
0.39
0.05
±0.
045
.38
±0.
059.
130
.70
±1.
830.
507.
44±
0.62
0.14
±0.
020.
19±
0.06
0.30
±0.
210.
12±
0.01
1.13
±0.
53
CH
4R
xrF
eed
DC
H4P
RY
CH
4/G
luG
asB
ase
H2
CH
4L
acta
teF
orm
ate
Ace
tate
Ace
toin
Eth
ano
lB
uty
rate
SS
34.
00.
272.
72±
0.05
0.21
±0.
010
4.42
±0.
016
.79
±0.
00.
041
.15
±0.
740.
00±
0.0
0.01
±0.
021.
97±
0.04
1.97
±0.
040.
01±
0.0
0.46
±0.
02
SS
45.
00.
334.
36±
0.02
0.22
±0.
017
1.74
±1.
9710
.08
±0.
00.
038
.13
±0.
640.
01±
0.01
0.24
±0.
072.
12±
0.03
1.95
±0.
590.
08±
0.08
0.52
±0.
14
SS
56.
00.
400.
49±
0.01
0.03
±0.
064
.44
±1.
6816
.05
±0.
027
.27
±3.
0812
.01
±0.
160.
00±
0.0
2.99
±0.
261.
74±
0.47
0.47
±0.
041.
38±
0.52
0.88
±0.
31
Sta
nd
ard
med
iaw
asfe
dd
irec
tly
tob
oth
reac
tors
atth
ein
dic
ated
flo
wan
dd
ilu
tio
nra
tes,
and
efflu
ents
fro
mb
oth
sen
td
irec
tly
tow
aste
.H
2P
Ris
the
pro
du
ctio
nra
teo
fh
ydro
gen
gas,
CH
4P
Ris
the
met
han
e-p
rod
uct
ion
rate
,Y
H2/
Glu
and
YC
H4/G
luar
eth
em
ola
ryi
eld
so
fH
2an
dC
H4
togl
uco
se,
resp
ecti
vely
,G
AS
isth
eo
vera
llga
sp
rod
uct
ion
rate
,B
AS
Eis
the
bas
ead
dit
ion
rate
,H
2an
dC
H4
(%)
are
the
hea
dsp
ace
com
po
siti
on
of
H2
or
CH
4,
resp
ecti
vely
(Bal
ance
of
gas
to10
0%is
com
pri
sed
mo
stly
of
CO
2w
ith
the
bal
ance
bei
ng
N2).
On
lysm
all
amo
un
tso
fC
H4
(0.3
–0.6
%)
wer
em
easu
red
inth
ega
sh
ead
spac
eo
fth
eH
2-p
rod
uct
ion
reac
tor.
Th
eC
H4-p
rod
uct
ion
reac
tor
pro
du
ced
no
H2
exce
pt
atth
eh
igh
est
feed
rate
(12.
0%).
No
glu
cose
was
mea
sure
din
any
of
the
efflu
ents
.
M. Cooney et al. / Bioresource Technology 98 (2007) 2641–2651 2645
extracts. The carbon leaving in the form of unconsumedpeptone or yeast extract was calculated assuming that only60% of both media was consumed (40% of each was left).While it is acknowledged that this value is practicallyimpossible to accurately approximate, it neverthelessaccounts for some carbon and the effect of changing thisvalue is discussed below.
2.6. Statistical analysis
Fermentation data presented in Tables 1, 3, and 4 arepresented as averaged values with standard deviations cal-culated in Microsoft Excel over two or three measurementstaken on separate days after steady state was achieved. Inall cases HPLC (e.g., VOAs) and GC (e.g., headspace %H2 and CH4) measurements were taken 2, 3, and even 4days prior to the steady state being declared, with these val-ues used as a baseline to gauge the correctness of the steadystate measurements (i.e., their values were used as a trendto gauge if any one measurement was unrealistic). Theexception is the base consumption rate which is presentedas a single value taken from linear regression fit over a24 h period (R2 values above 0.95). For those cases wherethe headspace H2 and CH4 values were low (e.g., less than1%) compared to the accuracy of the TCD detector, aver-ages and standard deviations are not given. The statisticalrelevance of comparisons between steady states withinTables 1, 3 and 4 (H2 Rxr or CH4 Rxr) were evaluatedusing Tukey’s pairwise comparisons with a family errorrate of 5% (i.e. 95% confidence level) executed throughthe ANOVA one-way analysis function of the statisticalsoftware analysis program MINITAB (SBTI, San MarcosTexas).
3. Results and discussion
3.1. Culture start-up and individual reactor performance with
varying dilution rates
Both reactors were inoculated with digester sludge (seemethods) and were then operated with standard media atdilutions of 1.0 followed by 2.0 L d�1 over a period of 2months, to autoselect the appropriate cultures, with thefirst, 2 L hydrogen-production, reactor maintained at pH5.5 and the second, 15 L methane, reactor maintained atpH 7.0. After a relatively long lag time, both reactors estab-lished cultures that exhibited the desired headspace gascomposition (e.g., �35% H2 in the hydrogen-productionreactor and 54% CH4 in the methane-production reactor,with the remainder mostly CO2 in both cases). The resultsindicate that these reactors operated with standard media,at these pH levels and dilutions, select for bacterial culturesthat are mainly either H2 or CH4 producers.
The effect of dilution rate at constant substrate concen-tration with standard medium was first investigated foreach reactor operated independently. The media supplyrate was increased in steps and the results are summarized
2646 M. Cooney et al. / Bioresource Technology 98 (2007) 2641–2651
in Table 1. Results for the methane-production reactor areonly reported for dilution rates at 0.27 h�1 and abovebecause of the time it took for this system to achieve stableand steady operations. Both reactors were quite stable andable to withstand occasional disturbances, such as downtime due to power outages, reductions in the feed ratedue to crimping of the feed lines in the peristaltic pumps,or fluctuations in temperature or mixing, when the mag-netic stir bars were temporarily destabilized.
From Table 1, the H2-production rate increased withfeed rate until 5.0 L d�1, achieving a total gas evolutionrate of 139 ml h�1, of which 35% were H2 and the balancemostly CO2, with only trace amounts of CH4. As the feedrate was increased to 6 L d�1, total gas and H2-productiondeclined to values that were statistically below all steadystates except that for 2 L d�1 (i.e., SS1, Tukey’s pairwisecomparison, family error rate of 5%), presumably due towash-out of the bacterial population (Van Ginkel andLogan, 2005). Although the glucose provided was com-pletely consumed, for all dilutions the calculated H2 yieldsfrom glucose were below 1.0 mol mol�1, an observationassumed due to domination by lactic acid producing bacte-ria, as lactate was the main metabolite at each steady state(Table 1, H2Rxr). Their dominance and low H2 yields weredue to the use of non-heat treated inoculum, and the lowmixing in the reactor, both of which favored lactic acidbacteria, lack of mixing by increasing the liquid-side partialpressure of hydrogen, both reflecting the conditionsexpected in process scale-up. Although heat treatment willkill most facultative bacteria and select for spore formingclostridia, which allows experimental production of hydro-gen gas at higher yields, this would be only a transientphenomenon, as contaminants entering with the feed orotherwise would rapidly dominate such reactors, parti-cularly when operating at short retention times of 1 dayor less. Indeed, even in these experiments the feed tank,although refrigerated, exhibited visual growth of bacteria,presumably lactic acid bacteria, as did the media feed lines,which had to be cleaned and replaced about every 2 days.
The methane-production rate in the CH4-productionreactor similarly peaked at a feed rate of 5.0 L d�1 with astatistically significant gas evolution rate of 171.74 ml h�1
Table 2
Carbon mass balance for non-integrated H2- and CH4-production reactors, respectively
H2Rxr In Liquid phase – out
Feed
(g C d�1)
VOA’s
(g C d�1)
YE + Pep
(g C d�1)
CO2
(g C d�1)
SS1 10.83 6.53 1.01 0.06
SS2 15.46 12.16 1.52 0.08
SS3 19.17 13.97 2.02 0.09
SS4 26.79 17.18 2.53 0.14
SS5 31.69 22.67 3.04 0.21
CH4Rxr In Liquid phase – out
SS3 19.17 4.25 2.02 3.13
SS4 26.79 6.09 2.53 3.40
SS5 31.69 11.94 3.04 0.78
and methane head space composition of 38.13% (Table 1,Tukey’s pairwise comparison, family error rate of 5%).As the feed rate was increased to 6 L d�1 the gas produc-tion rate decreased to a value that was significantly lowerthan SS3 and SS4 (Table 1, Tukey’s pairwise comparison,family error rate of 5%). The highest methane to glucoseyields of 0.21 and 0.22 were found at the lower feed ratesof 4.0 and 5.0 L d�1, with a statistically significant dropto 0.03 at the highest feed rate of 6.0 L d�1 (Table 1,Tukey’s pairwise comparisons, family error rate of 5%).The dominant liquid phase metabolites at lower feed rateswere acetate and acetone (�1.97 and 2.0 g L�1) with for-mat, acetate, ethanol, butyrate approaching or surpassinglevels above 1.0 g L�1 at the higher feed rate of 6.0 L d�1.In the cases of formate, ethanol, and butyrate their valuesat 6.0 L d�1 were statistically higher than their levels at thelower steady states (Table 1, Tukey’s pairwise comparison,family error rate of 5%). In the case of acetate its value at6.0 L d�1 was statistically lower that its value at 5 L d�1
but only statistically indistinguishable from 4 L d�1 (Table1, Tukey’s pairwise comparison, family error rate of 5%).No lactate was present at measurable levels. The low meth-ane yield based on glucose is likely due to both the rela-tively short retention times of these reactors and the highbiomass yield coefficient possible through use of the yeastand peptone hydrolyzates.
A carbon balance around the H2-producing reactorclosed to an average of 1% over all steady states (e.g.,between 93% and 110%) but only to an average of 70%over all steady states for the CH4-producing reactor, leav-ing a significant fraction of the carbon unaccounted for(Table 2). The HPLC chromatograms of the CH4-produc-tion reactor effluent did not reveal the emergence of anynew and unaccounted peaks that could explain the missingcarbon. One possible source could be in the estimatedpercentage of feed peptone and yeast extract that wasconsumed. If the percent of the carbon in the peptoneand yeast extract that been consumed in both reactors isreduced to 20%, the carbon mass balance averaged overall steady states increases to 11% excess carbon for theH2-production reactor while the deficiet for the CH4-pro-duction reactor reduces to 20%. Although reasonable as a
Gas phase – out C balance
CH4
(g C d�1)
Biomass
(g C d�1)
CH4
(g C d�1)
CO2
(g C d�1)
Co/Ci
(%)
0.00 2.08 0.00 0.27 92
0.00 2.92 0.00 0.40 111
0.00 3.53 0.01 0.52 105
0.00 5.12 0.01 1.02 97
0.00 6.02 0.00 0.36 102
Gas phase – out C balance
0.14 3.53 0.49 0.7 74
0.16 5.12 0.75 1.21 75
0.01 6.00 0.09 0.45 72
Tab
le3
Tan
dem
(in
tegr
ated
)o
per
atio
no
fth
ere
acto
rsat
vary
ing
dil
uti
on
rate
san
dco
nst
ant
feed
com
po
siti
on
Inte
grat
ed
H2R
xrF
eed
Effl
uen
tD
H2P
RY
H2/G
lT
ota
lga
sB
ase
H2
CH
4L
acta
teF
orm
ate
Ace
tate
Ace
toin
Eth
ano
lB
uty
rate
(Ld�
1)
(Ld�
1)
(d�
1)
(mm
ol
L�
1d�
1)
(mm
ol
mm
ol�
1)
(ml
h�
1)
(ml
h�
1)
(%)
(%)
(gL�
1)
(gL�
1)
(gL�
1)
(gL�
1)
(gL�
1)
(gL�
1)
SS
12.
0N
A1.
011
.78
±09
30.
20±
0.02
65.8
0±
2.59
4.13
37.6
9±
1.48
0.22
6.72
±0.
090.
67±
0.05
0.00
±0.
00.
20±
0.0
0.25
±0.
00.
27±
0.02
SS
23.
0N
A1.
515
.11
±0.
070.
19±
0.01
91.0
8±
1.13
8.13
34.6
5±
0.17
0.64
6.15
±0.
090.
91±
0.01
0.40
±0.
010.
18±
0.06
0.22
±0.
010.
29±
0.0
SS
34.
0N
A2.
012
.34
±0.
710.
13±
0.01
70.2
5±
3.25
5.95
37.0
±0.
41.
16.
29±
0.23
0.86
±0.
130.
43±
0.02
0.16
±0.
050.
16±
0.0
0.32
±0.
05
SS
45.
0N
A2.
531
.76
±0.
028
0.22
±0.
021
7.95
±1.
655.
2130
.7±
0.04
0.55
5.78
±0.
361.
34±
0.07
0.5
±0.
090.
24±
0.14
0.48
±0.
170.
41±
0.04
SS
56.
0N
A3.
016
.9±
0.09
0.12
±0.
093
.77
±0.
3112
.64
37.9
7±
0.07
0.5
6.57
±1.
890.
62±
0.32
0.14
±0.
050.
14±
0.02
0.15
±0.
010.
63±
0.35
CH
4R
xrF
eed
Effl
uen
tD
H2P
RY
H2/G
lT
ota
lga
sB
ase
H2
CH
4L
acta
teF
orm
ate
Ace
tate
Ace
toin
Eth
ano
lB
uty
rate
SS
12.
02.
00.
274.
37±
0.45
0.57
±0.
0611
2.49
±10
.45
6.79
0.0
61.1
3±
0.38
0.34
±0.
010.
36±
0.05
0.00
±0.
00.
55±
0.01
0.47
±0.
021.
18±
0.04
SS
21.
03.
00.
274.
17±
0.0
0.39
±0.
013
8.25
±0.
858.
250.
047
.5±
0.0
0.00
±0.
00.
29±
0.06
1.00
±0.
00.
09±
0.09
0.03
±0.
030.
97±
0.01
SS
30.
04.
00.
273.
58±
0.02
0.27
±0.
011
2.18
±2.
273.
270.
050
.39
±1.
290.
00±
0.0
0.07
±0.
030.
75±
0.01
0.01
±0.
00.
38±
0.0
0.79
±0.
01
SS
40.
05.
00.
335.
44±
0.06
0.27
±0.
015
0.99
±0.
934.
130.
054
.07
±0.
30.
01±
0.01
0.02
±0.
010.
89±
0.1
0.13
±0.
130.
19±
0.12
0.95
±0.
08
SS
50.
06.
00.
404.
98±
0.14
0.27
±0.
0015
8.57
±3.
035.
362.
16±
0.23
49.2
3±
0.04
0.00
±0.
00.
01±
0.01
0.79
±0.
10.
3±
0.09
0.11
±0.
01.
09±
0.13
Th
eeffl
uen
tfr
om
the
H2-p
rod
uct
ion
reac
tor
was
dir
ecte
din
toth
ese
con
dC
H4-p
rod
uct
ion
reac
tor,
inam
ou
nts
assh
ow
n.H
ead
ings
asd
escr
ibed
inT
able
1.M
eth
ane
con
cen
trat
ion
inth
ega
sp
rod
uce
db
yH
2-p
rod
uct
ion
reac
tor
was
vari
edfr
om
less
than
1%to
abo
ut
2%,
bu
tH
2w
asab
sen
tfr
om
the
gas
ph
ase
of
the
CH
4-p
rod
uct
ion
reac
tor
exce
pt
for
2.2%
inth
eh
igh
est
load
ing
rate
.
M. Cooney et al. / Bioresource Technology 98 (2007) 2641–2651 2647
guide, the carbon mass balance is limited for processesusing complex media because of the difficulty in analyzingcomplex carbon sources. Future work could use total car-bon analyzers as a means to track carbon flows but thistechnique offers little insight into the metabolic flowsof nutrients and products. Nevertheless, the reasonableaccounting of carbon in the H2-production reactor did giveconfidence in those measurements conducted and presentedin this work, and suggests the missing carbon in the meth-ane reactor is a significant observation and not an artifactof poor measurements.
3.2. Two-phase integrated bioreactor operations – dilution,
feed strength, and C:N ratios
In the integrated, two-phase process, the effluent fromthe first H2-production reactor was fed into the secondCH4-production reactor. The results from the steady statemeasurements of these experiments are presented in Table3, with the head space gas composition data for the entiresix months of operations for the two reactors shown inFig. 2a and 2b. To help acclimate the methane producingculture to the new feed, initially only a percentage of theeffluent from the H2-production reactor was used, withthe remainder made up by standard media (steady statesSS1 and SS2, Table 3). After that, the two reactors werefully integrated (steady states SS3, SS4 and SS5 in Table3). For the H2-production reactor the highest yields (0.2,0.19, and 0.22 mmol mmol�1) were found at feed rates of2.0, 3.0 and 5.0 L d�1 with H2 headspace compositionsranging between 30.7% and 37.7% and the balance mostlyCO2). Statistically significant lower rates were found atfeed rates of 4.0 and 6.0 L d�1 (Table 3, Tukey’s pairwisecomparison, family error rate of 5%) although the head-space H2 was statistically equivalent to the other steadystates at 37% and the balance mostly CO2). The maximumstatistically significant gas production rate was217.95 ml h�1 at a feed rate of 5 L d�1 (Table 3, Tukey’spairwise comparison, family error rate of 5%). For theCH4-production reactor the statistically highest yield wasfound at the lowest feed rate of 2 L d�1 with statisticallylower (but statistically indistinguishable among themselves)yields at the higher feed rates of 4.0, 5.0, and 6.0 L d�1
(Table 3, Tukey’s pairwise comparison, family error rateof 5%). These results demonstrate that two-phase opera-tion can be achieved, with dominance of hydrogen-produc-tion and methane-production in their respective reactors,albeit at relatively low yields (based on glucose input).
It is observed that the hydrogen yields were generallyfound to be higher in the integrated process as opposedto the non-integrated process. At first glance this shouldnot happen since both reactors were operated similarlyand unaffected by the process of integration (e.g., the reac-tion of glucose fermentation identical across both pro-cesses). Inspection of the data in Tables 1 and 3 revealsthat the source of this difference is due to statisticallyhigher gas production rates (ml h�1) at each dilution rate.
H2 rxr
0
20
40
60
80
100
120
0 1000 2000 3000 4000 5000
Time (h)
Hea
d s
pac
e (
%) CO2
H2
O2
N2
CH4
sum
CH4rxr
0
20
40
60
80
100
120
0 1000 2000 3000 4000 5000
Time (h)
Hea
d s
pac
e (%
) CO2
H2
O2
N2
CH4
sum
a
b
Fig. 2. Head space gas composition for the integrated reactors over 187days. Legend: a, first hydrogen producing reactor; b, second methaneproducing reactor.
2648 M. Cooney et al. / Bioresource Technology 98 (2007) 2641–2651
Otherwise, the physiology of the culture remained statisti-cally similar (i.e., in terms of head space composition ofH2 and metabolite profiles at each steady state). In reality,however, the two cultures were distinctly different in termsof age. In our work, the non-integrated cultures wereapproximately two and half months older than the inte-grated cultures. Specifically, we elected to first establishthe cultures in the non-integrated process to demonstrateproof of principle regarding culture selection for H2 andCH4 production, respectively, in the two bioreactors (i.e.,while operated under different selection pressures of pHand dilution rate) before inoculating the integrated bio-reactors. In both systems we found that the older H2 pro-ducing cultures tended to produce lower overall gasproduction rates (ml h�1), an observation we attribute tothe effects of culture selection and growth of biofilm. Thatsaid, we suggest that future work in this area should con-sider overall gas production rates as a function of cultureage, specifically when grown in continuous culture underselective pressures of pH and dilution rate, and that adetailed analysis of species composition and extent of bio-film be measured and directly correlated.
The liquid phase metabolites, dominated by lactate, fedfrom the H2-production reactor to the CH4-productionreactor, were mostly consumed with only small amountsof mostly indistinguishable acetate, acetoin, and butyrateunconsumed (Table 3, Tukey’s pairwise comparison, fam-
ily error rate of 5%). The methane to glucose yields wereagain relatively low at 0.27 mol mol�1 for a feed rate of5 L d�1, although this was slightly higher than for whenfed glucose directly (�0.22 mol mol�1, see CH4Rxr, Table1). As with the H2-production reactor, methane yieldsand gas production rates were higher for the integrated sys-tem relative to the independently operated reactors. Meth-ane yields on glucose were highest for the mixed feed(i.e. 0.57 mol mol�1 for SS1 and 0.39 mol mol�1 for SS2)although this could be attributed to the lower feed rates(1 and 2 L d�1). However, even with this increase, overallmethane yields were still low. Although not shown, themethane to lactate yields were roughly a bit above halfthe values calculated for glucose (calculations not shown),which is expected given two acetates are produced per moleof glucose.
The response of this two-phase integrated system toincreased feed strength (substrate concentrations) andchanging C:N ratio was investigated at a constant feed rateof 5 L d�1 (Table 4). In the H2-production reactor, thehydrogen per glucose yields were lowest (Table 4, Tukey’spairwise comparison, family error rate of 5%) at glucosefeed levels of 20 and 40 g L�1 and with the feed nitrogenremaining at its original level of 1N (C:N of 13.35 and24.88 mol mol�1). As the nitrogen was doubled from 1N–2N, however, the yields increased to their statistically high-est levels of 0.26 and 0.34 mmol mol�1 (Table 4, Tukey’spairwise comparison, family error rate of 5%). The hydro-gen-production rate remained statistically indistinguishable(Table 4, Tukey’s pairwise comparison, family error rate of5%) as the glucose feed level was increased from 20 to40 g L�1 without an concomitant increase in the nitrogenfeed concentration (i.e. the C:N ratio increased from13.35 to 24.88 mol mol�1). However, as the nitrogen feedwas doubled (1N–2N) at a glucose feed rate of 40 g L�1,the hydrogen-production rate doubled and rose yet againwhen the glucose concentration was raised from 40 to60 g L�1. The gas production rates statistically decreasedfrom 224 to 135 ml h�1 when the glucose was doubled from20 to 40 g L�1 at the lowest fixed feed nitrogen concentra-tion of 1N, but increased significantly to 353 when the feednitrogen was increased to 2N at 40 g L�1 (Table 4, Tukey’spair wise comparison, family error rate of 5%). Theseresults suggest that the gas production rates are generallyhigher at lower C:N ratios, and when H2 and CO2 arethe dominant headspace gases. Lactate, which was presentat each steady state, did not increase much when the glu-cose was doubled without a similar increase in nitrogen(SS6 to SS7: glucose 40 g L�1, 1N, C:N 24.88 mol mol�1)but doubled when the nitrogen was likewise doubled (SS7to SS8: glucose 40 g L�1, 2N, C:N 13.35 mol mol�1,Tukey’s pairwise comparison, family error rate of 5%).All other metabolites were generally low except for acetatewhich was significantly highest at SS7 (glucose 40 g L�1,C:N 24.88) and butyrate which generally increased witheach increase in glucose or nitrogen (Table 4, Tukey’spairwise comparison, family error rate of 5%).
Tab
le4
Tan
dem
(in
tegr
ated
)o
per
atio
no
fth
ere
acto
rsat
afi
xed
dil
uti
on
rate
and
incr
easi
ng
feed
con
cen
trat
ion
Inte
grat
ed
H2R
xrF
eed
C:N
H2P
RY
H2/G
luG
asB
ase
H2
CH
4L
acta
teF
orm
ate
Ace
tate
Ace
toin
Eth
ano
lB
uty
rate
(gg�
1)
(mm
ol
L�
1d�
1)
(mm
ol
mm
ol�
1)
(ml
h�
1)
ml
h�
1(%
)(%
)(g
L�
1)
(gL�
1)
(gL�
1)
(gL�
1)
(gL�
1)
(gL�
1)
SS
613
.35
29.9
1±
0.46
0.08
±0.
022
4.28
±0.
095.
7327
.74
±0.
930.
412
.57
±0.
060.
51±
0.04
0.65
±0.
210.
77±
0.33
0.63
±0.
.15
2.78
±0.
08
SS
724
.88
23.1
9±
1.54
0.04
±0.
013
5.13
±1.
2215
.17
36.1
8±
2.74
0.0
15.5
±2.
280.
31±
0.04
10.4
9±
0.45
0.31
±0.
030.
49±
0.06
4.5
±1.
34
SS
813
.35
57.8
±6.
270.
26±
0.03
353.
25±
3.93
15.1
834
.46
±3.
360.
030
.16
±1.
050.
42±
0.04
2.82
±2.
551.
17±
0.55
0.41
±0.
116.
38±
0.51
SS
919
.11
84.5
7±
9.54
0.34
±0.
0444
6.06
±14
.41
12.7
139
.9±
3.22
0.0
31.3
2±
2.59
0.40
±0.
03.
17±
1.42
0.60
±0.
420.
47±
0.11
6.39
±1.
23
CH
4R
xrF
eed
C:N
CH
4P
RY
CH
4/G
luG
asB
ase
H2
CH
4L
acta
teF
orm
ate
Ace
tate
Ace
toin
Eth
ano
lB
uty
rate
SS
6N
A5.
98±
0.14
0.12
±0.
025
9.47
±6.
144.
271.
87±
0.09
35.7
2±
1.0
0.00
±0.
00.
4±
0.57
2.24
±0.
490.
25±
0.13
0.21
±0.
062.
76±
0.31
SS
7N
A3.
69±
0.35
0.05
±0.
027
7.3
±9.
5415
.17
3.98
±2.
620
.48
±1.
90.
07±
0.03
1.2
±0.
217.
66±
0.33
0.14
±0.
150.
32±
0.28
6.13
±0.
01
SS
8N
A12
.95
±0.
370.
06±
0.0
867.
47±
15.0
8n
m�
0.9
±1.
2723
.1±
1.8
0.36
±0.
140.
09±
0.12
2.99
±0.
20.
14±
0.18
0.04
±0.
473.
4±
0.19
SS
9N
A5.
01±
2.05
0.02
±0.
0187
1.1
±2.
266.
756.
51±
6.44
9.05
±3.
648.
65±
0.47
1.76
±2.
495.
03±
0.06
0.19
±0.
080.
65±
0.24
5.66
±1.
02
Th
eH
2-p
rod
uct
ion
reac
tor
was
fed
feed
med
iaat
ara
teo
f5.
0l
d�
1w
ith
stan
dar
dm
edia
wit
hgl
uco
sean
d/o
rn
itro
gen
feed
con
cen
trat
ion
sm
od
ified
asfo
llo
ws:
SS
6:gl
uco
se20
gL�
1,
1N,
C:N
13.3
5;S
S7:
glu
cose
40g
L�
1,
1N,
C:N
24.8
8;S
S8:
glu
cose
40g
L�
1,
2N,
C:N
13.3
5;S
S9:
glu
cose
60g
L�
1,
2N,
C:N
19.1
1.T
he
efflu
ent
fro
mth
eH
2-p
rod
uct
ion
reac
tor
was
fed
ino
the
CH
4-p
rod
uct
ion
reac
tor.
Th
ed
ilu
tio
nra
tew
as2.
5d�
1fo
rth
eH
2-p
rod
uct
ion
reac
tor
and
0.33
d�
1fo
rth
eC
H4-p
rod
uct
ion
reac
tor.
Th
eco
nce
ntr
atio
ns
of
bas
eu
sed
for
pH
con
tro
lw
ere:
2.0
Mfo
rS
S6,
4.0
Mfo
rS
S7,
and
6.0
Mfo
rS
S8
and
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ter
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ere
sid
ual
glu
cose
con
cen
trat
ion
fro
mth
em
easu
red
feed
con
cen
trat
ion
.
M. Cooney et al. / Bioresource Technology 98 (2007) 2641–2651 2649
In the CH4-production reactor the methane to glucoseyield was statistically highest at the lowest feed glucoseand nitrogen feed concentrations (SS6, glucose 20 g L�1,nitrogen 1N). Thereafter, the yields decreased to andremained at indistinguishably low levels (�0.05 mmolmmol�1) with each subsequent increase in feed glucoseor nitrogen (Table 4, Tukey’s pairwise comparison, familyerror rate of 5%). The overall gas production rate statis-tically increased when the glucose was increased from 20to 40 g L�1 at 1N, and then dramatically tripled from277 to 867 ml h�1 when the nitrogen was doubled from1N to 2N with the feed glucose at 40 gl�1 (Table 4,Tukey’s pair wise comparison, family error rate of 5%).The following increase in feed glucose from 40 to60 g L�1 did not significantly increase the total gas pro-duction rate (Tukey’s pairwise comparison, family errorrate of 5%). The methane-production rates were generallyindistinguishable at �4 to 5 mmol L�1 d�1 except at SS8(feed glucose of 40 g L�1 and nitrogen at 2N) where itsvalue was significantly higher at 13 mmol L�1 d�1. Thehead space methane composition declined when the glu-cose was doubled from 20 to 40 and from 40 to 60 with-out a change in nitrogen (SS6 to SS7, and SS8 to SS9),suggesting that increasing the feed C:N ratio depressesmethane-production and stimulates hydrogen-production(substantial amounts of hydrogen and lactate appearedat the highest organic loading (SS9: glucose 60 g L�1,2N, C:N 19.11 mol mol�1)). Lactate levels were otherwisenot statistically significant although acetate and butyratelevels were present at reasonable concentrations (e.g.,between 2.5 and 6.0 g L�1) at all steady states. Increasesin glucose tended to destabilize the two phase system, ascould be seen from the emergence of H2 in the methanereactor head space, but which could generally be arrestedby the concomitant increase in the nitrogen source (e.g.,peptone and yeast extract). However, once a glucose feedload of 60 g L�1 was reached, the system began to crashwithout opportunity for increased nitrogen to stabilizethe system. This result suggests that although the two-phase system can be used to select for CH4 producers inthe second reactor, even though the H2 producers arepresent at sufficiently high organic loadings, can quicklydominate the culture.
This result is emphasized in Fig. 3 which presents thebatch growth of the methane producers after two thirdsof the CH4-production reactor media was emptied andreplaced with standard media that had been refrigeratedovernight. Prior to this experiment, the CH4-productionreactor had been run at conditions that had optimizedfor methane producers and possessed the expected headspace gas composition profile (e.g., CH4 � 50% with thebalance as CO2). The results clearly show that despite theinitial domination of methane producers, in the presenceof excess organic feed (e.g. glucose) the background hydro-gen producers quickly dominate, with the majority of headspace gas being composed of hydrogen and carbon dioxide,with the balance being methane.
0
2
4
6
8
0 5 10 15 20 25 30Time (h)
Co
nce
ntr
atio
n g
L-1
0.00
1.00
2.00
3.00
4.00
gd
w L
-105
101520253035
0 5 10 15 20 25 30
Time (h)
Gas
evo
luti
on
ml h
-1
a
b
Fig. 3. Biomass production and lactate and glucose consumption forbatch growth of methane producers (methanogens) in batch culture.(a): open squares, biomass; open circles, glucose; and solid line, lactate.(b): open squares, H2; open circles, CH4; open diamonds CO2.
2650 M. Cooney et al. / Bioresource Technology 98 (2007) 2641–2651
4. Conclusions
We have shown that the selection pressure of pH anddilution rate is sufficient to select for acetogenic and meth-anogenic bacteria in their respective first and secondphases. The percentage of hydrogen in the head space ofthe first reactor was between 30% and 40% when fed stan-dard media, regardless of varying dilution rates. The sameresults in terms of H2-production were obtained, for bothindividual vs. tandem operations, as expected since thesedid not differ for the first stage operations, demonstratingthe results reinforce the reproducibility across all experi-ments.
The predominant product of glucose metabolism in theH2-production reactor was lactate, regardless of the dilu-tion rate. Dominance of lactic acid bacteria in this reactor,accounted for the relatively low H2 yields and rates obser-ver, compared to what has been observed with clostridiadominated cultures. In some sense lactate (and acetoin)are stealing electrons away from H2-production, and thatredirecting metabolism away from these end productswould theoretically improve H2-production. Some insightto this can be seen by inspecting the hydrogen yield onthe basis of glucose that was converted to products otherthan lactate and acetoin. For Table 1 these yields variedfrom 0.6 to 2.0 mmol mmol�1 (calculations not shown).The lactate was effectively consumed in the second reactor,for most dilutions and organic loadings, but overall yieldsof methane were relatively low. Composition of liquidphase metabolites were fairly constant across all dilutionrates, suggesting that the methanotrophs in the secondreactor acted like a typical anaerobic digestor. However,
the integrated two-phase system, at all dilution rates,improved yields and CH4-production rates. What two-stage integration did not provide, however, was an acceler-ated metabolism in the second reactor that would supportthe use of shorter retention times. It was hoped that such aseparation would permit the use of smaller second stagereactors and shorten the overall processing time, both ofwhich would improve the cost of application in industry.These results, however, indicate that this advantage wasnot achieved with respect to methane-production in thesecond reactor, regardless of the presence of lactate fromthe first reactor.
In the non-integrated system a sharp decline in methane-production in the CH4-production reactor was observedwhen the feed rate was raised from 5 to 6 L d�1. Thisdecline was accompanied by a increase in hydrogen-pro-duction, suggesting that at higher dilution rates, and inthe presence of glucose feed, the hydrogen producing bac-teria begin to dominate the slower growing methanogens.By contrast, in the integrated two-phase system, the steadydiet of low energy organic acids from the first hydrogenproducing reactor maintained a selective pressure in favorof the methanotrophs. This was also seen at the highestdilution rates, where in the non-integrated system fedglucose, methane yields dropped dramatically vs. constantmethane yields in the integrated system.
The effect of increasing the carbon and/or nitrogen con-centrations in the feed media had the effect of increasingthe overall gas production rates in both reactors. Howeverincreasing glucose relative to nitrogen (e.g., increasing theC:N ratio) generally had the effect of destabilizing the sec-ond CH4-production reactor. The general effect was to firstincrease the biomass concentration in the first reactor,which resulted in more acetogenic bacteria pumped intothe second CH4-production reactor, along with a greaterconcentration of lactate. Consequently, acetogens in thesecond CH4-production reactor eventually dominate theslower growing methanogens. This left the two-phase sys-tem somewhat vulnerable to disturbances. Typical distur-bances could include a brief over pumping of media fromthe first H2-production reactor into the second due to prob-lems with the level control (due to foaming), pH control inthe second CH4-production reactor due to periodic loss ofmixing (i.e. if the stirrer bar was pushed to the side) orappearance of foaming.
In summary, two-phase operation for the production ofH2 and CH4 mixtures can be achieved with appropriateapplication of selection pressure through pH and dilutionrate, and at low organic loadings. The process promiseslow H2 and CH4 yields (relative to glucose) at industrialscale unless energy efficient techniques are found to inhibitthe introduction of lactic acid producing bacteria in thefirst H2-production reactor, to provide good mixing, andto decrease the liquid side partial pressure of H2. Whilethese items should be the focus of future research, a fewcomments are provided. While increasing the gas transfercoefficient (e.g., through increased mixing) to reduce the
M. Cooney et al. / Bioresource Technology 98 (2007) 2641–2651 2651
liquid-side hydrogen partial pressure in the reactors couldalso be advantageous, it was not applied in this workbecause its application is not practical at industry scaledue to its high energy requirements. Displacement of lacticacid producing bacteria will require additional processmethodologies such as sterilization of the feed or the useof repetitive batch innoculations with desired H2 produc-ers, something that will need to be addressed in industrialapplication. In our application of this strategy, however,we found that, at pH 5.5, the lactic acid bacteria still dom-inated and lactic acid production proceeded in tandem withglucose utilization and biomass production, and withoutsignificant H2- or CO2-production, and that after theglucose was fully consumed, the H2 (and CO2)-productionwas initiated and continued for some time This suggests aneed to re-innoculate with dense cultures of the preferredH2 producing bacteria along with some heat treatment ofthe pre-existing culture to reduce dominance by lactic acidbacteria.
Acknowledgements
Financial support was provided by the ‘‘Energy Inven-tions Small Grants Program’’ of the California EnergyCommission (Grant Number 5295 A/02-09). The authorswould also like to acknowledge David Harris and JamesJolly from the SOEST Engineering Support Facility forthe design of the gas measurement system.
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