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494 24 Fixed Film Stationary Bed and Fluidized Bed Reactors

Symbols and Abbreviations

Symbol Name unit

a geometrical factor ci concentration of the substrate i mol mP 3cib substrate concentration at the biofilmliquid interface mol mP 3cip substrate concentration at the biofilmparticle interface mol mP 3cis substrate concentration in bulk liquid mol mP 3n expansion index D diffusion coefficient m2 sP 1D eff effective diffusion coefficient m2 sP 1

dR reactor diameter mdP particle diameter mdV/S Sauter diameter m

j molar substrate flux mol mP 2 sP 1k0 kinetic constant (zero order) mol mP 3 sP 1k1 kinetic constant (1st order) sP 1k e mass transfer coefficient m sP 1r reaction rate mol mP 3 sP 1Re Reynolds number Re P particle Reynolds number Re T particle terminal Reynolds number

q r amount of particles Sc Schmidt number Sh Sherwood number t time s

g gravitational acceleration m sP 2C D drag coefficientu superficial upflow velocity m sP 1uT terminal settling velocity m sP 1uT* terminal settling velocity in a particle swarm m sP 1

X biomass concentration kg mP 3 xb bioparticle radius m xc substrate concentration depth m xp particle radius m xs radial distance of cS ` 0 mY yield coefficient BV, COD volumetric loading rate kgBV, inert volumetric loading rate of non-biodegradable solids kg voidage sphericity X biomass retention time d Thiele modulus 1m 1st (reaction order) modified Thiele modulus density kg mP 3

P particle density kg mP 3

L liquid density kg mP 3 effectiveness factor dynamic viscosity kg mP 1 sP 1

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hydrates such as sugar, starch, cellulose, andhemicellulose, protein and fat, which readilyundergo bacterial degradation to fatty acids,mainly acetic, propionic, butyric, and lacticacid. The majority of installations are in thepotato, starch, and sugar industries, in fruit,vegetable and meat processing, in cheese,yeast, alcohol, citric acid, pectin manufactur-ing, and in the paper and pulp industries. Theconcentrations of substrates are typically inthe range of 550 kg(COD) mP 3 , which are di-luted by recirculation (loop reactor) down toless than 2 kg(COD) mP 3 .

In some countries like China systems for thetreatment of domestic sewage also play a major role (YI-ZHANG and LI-BIN, 1988). Fur-thermore, considerable efforts were under-taken in order also to treat inhibitory or toxicsubstances. Thus cyanide, formaldehyde, am-monium, nickel, and sulfide containing waste-water were investigated,and it has been shownthat methanogens can accommodate even torather high concentrations of such toxicants,depending on the retention time (PARKIN andSPEECE, 1983). Furthermore, organochlorine

compounds in kraft bleaching effluents and inpesticide containing water could be treated tothe stage of mineralization by adapted bio-films, including chloroform, chlorophenols,chlorocatechols, and similar compounds aswell as chlorinated resin acids (SALKINOJA-SA-LONEN et al., 1983).These positive results wereobtained on the laboratory scale.The degrada-tion of furfural in sulfite evaporator conden-sate can proceed to a conversion of about 90%(NEY et al., 1989).

2 Basic PrinciplesReactors are tubes with fixed bed internals

or fluidized suspended particles, which serveas a support for biomass immobilization. Thedimensions range from 10500 m3 with a ratioof height to diameter from 15. In general anexternal loop recycles a part of the effluent tothe inlet, where mixing with the wastewaterprovides for its dilution to non-inhibitory sub-strate concentrations and pH. In a few casestapered beds have been used; most of the

fluidized beds are provided with a settlingzone with a larger diameter at the top of thereactor.

2.1 Biofilm FormationThe basis for the use of packed bed and

fluidized bed systems is the immobilization of bacteria on solid surfaces. Many species of bac-teria (and also other microorganisms) have theability for adhering to supporting matrices.

While in the aerobic wastewater treatmentimmobilized bacteria have been used since thebeginning of this century, the application of those systems to the anaerobic wastewatertreatment is relatively new.

The fundamentals of bacterial adhesion andgrowth on solid surfaces are discussed inChapter 4, this volume, thus only some aspectsconcerning anaerobic fixed films will be con-sidered here.The preconditioning of solid sur-faces is influenced by both environmental con-ditions (e.g., pH,T ) and by the surface itself (e.g.,hydrophobicity, surface charge) (GERSON

and ZAJIC, 1979). With the addition of cationicpolymers (STRONACH et al., 1987; DIAZ-BAEZ,1988) or slime-producing bacteria (DIAZ-BAEZ, 1988) the initial anaerobic biofilm at-tachment could be improved, but the biofilmdevelopment was worse than in systems with-out these components. JRDENING (1987) re-ported a positive effect with the supplementa-tion of calcium.

The primary adhesion of cells to the surfaceis due to hydrogen bonds,van der Waals forcesand/or electrostatic interactions (DANIELS,1980). This reversible form of adhesion canbecome irreversible by the production of exo-polymeric substances (EPS), which act as aglue (FLEMMING, 1991). Experiments on thefirst steps of the formation of anaerobic bio-films gave different results: while LAUWERS etal. (1990) and SANCHEZ et al. (1996) found fac-ultative anaerobic bacteria as primary coloniz-ers, SREEKRISHNAN et al. (1991) observed thatthe biofilm formation was initiated by methan-ogenic bacteria.

After a lag phase, which seems to be neces-sary for an adaptation of the microorganismsto the new environment, exponential growthof bacteria begins. The growth rate is mainly


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2 Basic Principles 497

determined by substrate transport and tem-perature (HEIJNEN et al., 1986; CHARACKLIS,1990).

2.2 Biofilm CharacteristicsIn view of the wide range of possible biofilm

compositions it is obvious that biofilm thick-ness does not correspond to the activity of thebiocatalyst. HOEHN (1970) reported that thehighest biofilm density can be found, if thebiofilm thickness corresponds to the activebiofilm thickness, i.e., the substrate-penetratedpart of the biofilm.

2.3 Kinetics and Mass TransferThe reaction kinetics for any process change

with immobilization of the catalyst. In biofilmreactors the following mass transfer processeshave to be considered:

(1) transport of substrate from the fluid to

the surface of the support through theboundary layer (external mass trans-fer),

(2) transport of substrate from the surfaceinto the pores of the biocatalyst,

(3) reaction,

(4) (5) transport of products in the oppo-site direction of steps 1 and 2.

The transport of substrates and productsthrough the reactor is related to the hydro-dynamic characteristics of the system and ingeneral is much faster than steps 24. Masstransfer is mostly reduced by diffusion limita-tion.

Depending on the mode of limitation onecan distinguish between film diffusion andpore diffusion and combined limited systems.Fig. 2 illustrates schematically the resultingsubstrate profiles.

For the description of the activity changesby immobilization an effectiveness factor isusually used, defined as

p p f (S, Sh, ) (1)

2.3.1 External Mass TransferPassing a solid surface, the fluid characteris-

tics change from turbulent to laminar flow andproduce a boundary layer around the surface.The flux j 1 through the boundary layer is equalto a mass transfer coefficient and the concen-

observed reaction ratereaction rate at bulk

liquid conditions

Fig. 2. Substrate concentration profilesat an immobilized biocatalyst surface,(1) reaction rate-controlled system;(2) combination of reaction rate and diffu-sion controlled system; (3) diffusion-con-trolled system.

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tration gradient from the outer shell to theparticle surface:

j 1 p k e (cis P cib ) (2)

An analytical solution fork e can only be givenfor the ideal case of a single particle at infinitedilution. The mass transfer coefficient is givenby:

k e p 2 (3)

For more complex systems, such as packed orfluidized bed reactors, the external mass trans-fer is usually described by dimensionless anal-ysis which leads to the correlation character-ized by the Sherwood number:

Sh p f (Re, Sc)p (4)

Re p (5)

Sc p (6)

MULCAHY and LA MOTTA (1978) calculatedthe external mass transfer for a fluidized bio-film with a correlation given by SNOWDEN andTURNER (1967):

Sh p Re 0.5 Sc0.33 (7)

2.3.2 Internal Mass TransferThe diffusional transport of substrate in the

biofilm may reduce the reaction rate. Forsteady state conditions the net diffusion rate isequal to the reaction rate:Writing in a generalform the resulting mass balance gives a secondorder differential equation

Deff

(c

)P r p 0 (8)

where D eff is the effective pore diffusion co-effficient,ci the concentration of the substrate

d c i

d x

a

x

d 2 ci

d x2

0.81

D

d p u

k e d pD

Dd p

i, x the position in the porous particle,r thereaction rate, and a a geometrical factor: for aplate ap 0, for a cylinderap 1, and for a sphereap 2.

The diffusion coeffficient in the biofilm (Deff )may be smaller than in water. KITSOS et al.(1992) found by comparison of acetate diffu-sion with lithium diffusion a value of 7%, relat-ed to the diffusivity in water (6.610P 10 m2 sP 1).OZTURK et al. (1989) calculated an effectivediffusion coeffficient of 1.710P 10 m2 sP 1 frommeasurements with inactive anaerobic bio-films. These values are low in comparison tothose determined for glucose or oxygen inaerobic biofilms, which are nearly the same asfor water (ONUMA et al., 1985;HORN and HEM-PEL,1996).KITSOS et al. (1992) explain this dis-agreement with principal differences betweenaerobic and anaerobic biofilms with regard tothe symbiosis between the bacterial groups inanaerobic biofilms.

Two boundary conditions can be defined atthe support interface and at the interfacebetween the biofilm and the boundary layer.

p 0 xp xp (9)

cip cis xp xb (10)

If a zero order reaction is assumed, the inte-gration of Eq. (5) with the boundary condi-tions yields an expression given by SHIEH andKEENAN (1986), i.e.

( )3 P 1.5( )2 c ( P )p 0 (11)where the Thiele modulus is defined as

p xb s (12)Hence, the effectiveness factor can be calculat-ed for cS ` 0.

For first order reactions SHIEH et al. (1982)show a good agreement between and theThiele modulus given by

p P (13)13 21mcoth(3 1m )

1m

k 0D eff ci

3 2

12

xc xb

xc xb

d c id x


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2 Basic Principles 499

with a modified 1st order Thiele modulus 1m

1m p ( k 1 ci)0.5

(14)

Since anaerobic reactors are generally treatedat substrate concentrations which are high inrelation to the substrate affinity constant, zeroorder kinetics are the useful way for calculat-ing diffusion limitation.

2.4 Support Characteristics

Many different support materials were test-ed for the application in fixed film stationaryand fluidized bed reactors. The major factorsfor bacterial attachment and growth related toboth systems are roughness and porosity. Theprocess configurations will be discussed separ-ately, because of differences in the relative im-portance of these factors and because addi-tional factors may apply to only one of thesystems.

2.4.1 Stationary Fixed FilmReactors

Supports in fixed bed reactors must meetspecific requirements for scale-up: Biomasstends to accumulate to a major part as sus-pended particles in the voids of the fixed bed.Biogas must separate from the fluid phase, andits transport must be possible through the full-scale reactor up to several meters and gashold-up of up to 3%, without major pressuredrop. Supports with small dimensions (rangeof a few mm or below) are, therefore, not suit-ed for this reactor type. Most supports testedwith success up to the pilot or industrial scaleoffer suffficient hydrodynamic radii (in gener-al more than 20 mm) and a very high void vol-ume (over 70%, mostly more than 90%). As aconsequence the support surface per volume israther low, the concentration of biomass in thesurface fixed film is low as well, and the sus-pended biomass in the void volume contrib-utes considerably to the activity (HENZE andHARREMOES , 1983; WEILAND and WULFERT,1986).The maximal volumetric load is distinct-

x3b P x3p x 2b

ly lower in general as compared to UASB, ex-panded, or fluidized bed reactors.

Supports utilized in most applications aretypically internals, such as Raschig or Pall rings,Berl or Intalox saddles, plastic cylinders, clayblocks or potter clay of dimensions in the rangeof typically 2060 mm (HENZE and HARRE -MOES, 1983; YOUNG and DAHAB, 1983; WEI-LAND et al., 1988). Trends in application favorsupports with a void volume of over 90% and asurface area in the range of 100300 m2 mP 3(BISCHOFSBERGER , 1993; AUSTERMANN-HAUNet al., 1993).

2.4.2 Fluidized Bed ReactorsThe choice of support material determines

the process engineering much more than forpacked bed reactors.This is due to the fluidiza-tion characteristics which depend on the den-sity and the diameter of the support.

For the calculation of the fluidization behav-ior, considering only ideally spherical particleswithout a biolayer,one starts with the terminal

settling velocity for a single particle at infinitedilution:

uT p s (15)where g is the gravitational acceleration. Theterminal settling velocityuT depends on theparticle diameter d p , the drag coefficientC D ,and the difference in the densities of the parti-cle p and the liquid L . C D correlates with theparticles Reynolds numberRe p

Re p p (16)

Equations describing the relation betweenC Dand Re p can be found in the literature. BIRD etal. (1960) gave a generally accepted formulafor the intermediate region of Reynolds num-bers (1~ Re P ~ 50) with

C D

p (17)

From these equations the single particle set-tling velocity can be calculated by iteration.

18.5Re P0. 6

uT d p L

4 g d p p P L3C D L

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However, for the calculation of the fluidizedbed expansion additional effects of the reactorwall, the characteristics of flow and of adjacentparticles have to be considered. Usually thefluidization is then described with empiricalequations.The correlation used most frequent-ly for fluidized beds is given by RICHARDSONand ZAKI (1954):

up uT* n (18)

where u is the superficial liquid velocity,uT* thesettling velocity of the particle swarm andequal to uT 10P d P /d R , the bed voidage and anexpansion indexn.n is given as follows, provid-ed that the particle diameter is much smallerthan that of the bed:

np 4.65 Re T ^ 0.2np 4.35Re TP 0.03 0.2 Re T ^ 1np 4.45Re TP 0.1 1^ Re T ^ 500np 2.39 500 Re T ^ 7,000

For anaerobic fluidized bedsn can normally becalculated forRe T in the range of 1500.

Most of the materials (e.g., granular activat-ed carbon, pumice, sepiolite) used as supportsfor anaerobic fluidized beds are not idealspheres and show a particle diameter distribu-tion. For that the Sauter diameter dV/S shouldbe used for the above calculations, defined as

dV/S p (19)

where q r is the amount of particles (of volumeor surface fraction) with the diameter x. Asphericity factor can be determined by mi-croscopical comparison of particle shape withmodel geometrical figures given by RITTEN-HOUSE (1943). From these values the voidagecan be described with a correlation of WENand XU (1966).

p 11 (20)

Even the volume contraction has to be consid-ered for the calculation of , if a mixture of dif-ferent particle sizes is used.A detailed calcula-tion procedure is given by OUCHIJAMA and TA-NAKA (1981).

1P 3

x q r (x)dx q r (x)dx

The easiest way for determining the fluidiza-tion behavior is to make experiments in labscale reactors. There one has to consider thatfor obtaining representative data it is neces-sary to adjust a ratio of 100 as a minimum forthe reactor and the particle diameter. How-ever, the growing biofilm and possible pre-cipitations of inert solids may cause significantchanges in the fluidization behavior duringwork.Therefore, it is very important to use realwastewater and to control the fluidization un-til a dynamic steady state is reached.

3 Reactor DesignParameters

The development of any anaerobic systemneeds the evaluation of some optimal condi-tions concerning several topics. For fixed filmsystems this includes especially the choice of the support, the reactor geometry, the start-up

procedure and the handling of excess sludge orinert support.

3.1 Scale-UpConcepts for scale-up have been summar-

ized by KOSSEN and OOSTERHUIS (1985),where dimensional analysis and rules of thumbmay be mentioned, since they provide guid-ance and recourse to practical experience.Fluid flow and fluidization can be treated withthe aid of Reynolds, Peclet, and Froude num-bers in order to estimate regimes appropriatefor technical-scale operation (MSCHE, 1998).However,a rational design seems very difficultdue to the high complexity of the systems.Therefore, empirical rules are mostly used inpractice for the design of technical reactors(HENZE and HARREMOES , 1983).

The parameter considered most importantis the load with biodegradable organics interms of COD. This must be correlated to theactive biomass in the reactor. So a load of 11.5 kg (COD) kgP 1 (VSS) dP 1 is consideredas the upper limit for stable operation, where


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3 Reactor Design Parameters 501

the following correlation can be used for guid-ance (HENZE and HARREMOES , 1983):

BV, COD p (21)

A range of further parameters of high signifi-cance should be taken into account: geometri-cal dimensions and H DP 1; recirculation rate,determining the substrate dilution and pH(and their gradient); residence time and distri-bution; mixing behavior; flow rate, pressuredrop, and energy requirements; fluidizationand bed expansion.

The recirculation rate, inlet substrate con-centration, and pH and its gradient are corre-lated to each other, and they are highly impor-tant aspects since the stability of the stationaryoperation greatly depends on it, as subse-quently discussed (BURKHARDT and JRDE -NING,1994;MSCHE, 1998).

An example for modeling of an industrialfluidized bed reactor as a guide to scale-up andoptimization of its operation has been present-ed by SCHWARZ et al. (1996, in press). The

model comprises those aspects which weremost sensitive for the results: material balanceequations for substrates and products in thegas and liquid phase, kinetics of the biologicaldegradation, mass transfer between gas and

X / X P BV, inertY

liquid phase, chemical equilibria as well asconvection and dispersion (Fig. 3).

Maximum individual reaction rates for thedifferent substrate components turned out tobe most sensitive for substrate and productconcentrations as well as for pH.With scale-upthe ratio of feed and recirculation increases,since the flow rate is limited by the carriersettling velocity. Axial gradients of pH valueswere calculated and turned out to be most sen-sitive to the load (Fig. 3) and the buffer capac-ity.Thus at high loading rates dynamic changesin the feed composition could lead to a pHdrop in the lower part of the reactor and to abreakdown of the system. Backmixing turnedout to be of minor importance in this system(with 500 m3 , 18 m height of the fluidized bed).A distribution of substrate feeding at two posi-tions, the bottom and the medium height,turned out to overcome these problems. Thecorresponding increase of the superficial up-flow velocity restricts this possibility only forhighly concentrated wastewaters. Otherwise,problems caused by a higher segregation of thebed would occur. For pumice as support mate-

rial an increase of the superficial upflow veloc-ity of 10% in the upper half of the fluidized beddid not cause any problems (MSCHE, 1998).

Furthermore, the buffer capacity was signif-icantly influenced by CO2 and its equilibrium

Fig. 3. Finite volume in a fluidizedbed reactor and considered fluxes(SCHWARZ et al., 1996).

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concentration as a function of pressure and of ions such as calcium present in the wastewater.

3.2 SupportA great number of support materials have

been investigated in laboratory scale reactorsfor the use in packed or fluidized bed systems(HENZE and HARREMOES , 1983). Despite that,the number of supports which are used evenfor technical-scale systems is rather low. Cer-tainly, one big problem for the realization of new processes is due to the difficulties of coop-eration with manufacturers. But often, theproblems result from shortcomings in the stud-ies with respect to requirements of large-scaleapplication. For a successful application of anymaterial as support in fixed film stationary orfluidized bed reactors on the technical scalethe following general requirements should befulfilled:

(1) availability of the material in big quan-tities ( 100 m3),

(2) low cost of the material (related to theachieveable performance; it should bein general lower than 150 $ mP 3),

(3) inert behavior (mechanically and mi-crobially stable) without toxic effectsand easy disposal,

(4) low pressure drop (low energy demandfor mixing or fluidization).

3.2.1 Stationary Bed ReactorsA selection of supports, which seem to have

been applied successfully on the pilot and in-dustrial scale is summarized in Tab. 1; exam-ples are shown in Fig. 4.

The biofilm thickness is of limited signifi-cance in fixed bed reactors since it makes upthe minor part of biomass. Data published aremostly in the range of 14 mm, the upper limitrelating to the inflow zone (top of a reactor op-erated downflow) (SWITZENBAUM, 1983; VANDEN BERG and KENNEDY, 1983; ANDREWS,1988, p.790/794).

More important is the biomass concentra-tion in the reactor which obviously is distribut-ed into two fractions, one immobilized in thebiofilm, the other being suspended in the voidvolume of the reactor (HALL, 1982; WEILANDand WULFERT , 1986). The organic dry biomassin pilot reactors is mostly in the range of 515 kg mP 3 (HENZE and HARREMOES , 1983).Gradients are observed,with the maximal con-centration near the reactor inlet [e.g., about15kgmP 3 in the lower and 4 kg mP 3 in the

upper part for upstream operation (WEILANDand WULFERT, 1986)]. Gradients were alsofound by HALL (1982), depending on the flowdirection: 6 fixed and 9 suspended for up-stream, and 9 fixed and 4 suspended for down-stream operation (all in kg mP 3 organic drymatter). For an industrial plant in starch pro-cessing a biomass concentration of 20 kg mP 3


Tab. 1. Typical Supports for Fixed Beds

Support Diameter Range Surface Bed Equivalent ReferencePorosity Pore

[mm] [m2 mP 3 ] Diameter

Raschig rings 10.16 4549 0.760.78 CARRONDO et al. (1983)Pall rings 90.90 102 0.95 20 YOUNG and DAHAB (1983)

corrugatedmodular blocks 98 0.95 46

Pall rings 25 215 SCHULTES (1998)Hiflow 90 90 65 0.965 ` 30 mm WEILAND et al. (1988)

Plasdek C. 10

modular blocks 148 0.96 ` 30 mm WEILAND and WULFERT(1986)Flocor R corrugated rings 320 97 ` 30 mmCeramic Raschig rings 25 190 0.74 ANDREWS (1988)

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was reported (SCHRAEWER , 1988). A highbiomass concentration was reported for a spe-cial support, sinter glass fillings (of high price)up to 65 kg mP 3 in lab-scale experiments, of

which 31 kg mP

3 were found to be immobiliz-ed within the inner pores of the support (NEYet al., 1989).

Fig. 4ac. Typical support materials foranaerobic stationary fixed film reactors:a Hiflow 90 (uncovered and biofilm-covered);b Flocor R ; c Plasdek C . 10(WEILAND and WULFERT, 1986).

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3.2.2 Fluidized Bed ReactorsThe energy demand for the fluidization of

the support is often discussed to be very highin contrast to other anaerobic techniques. Ingeneral, much higher volumetric flow rateshave to be raised in comparison to a CSTR.Nevertheless, the energy demand is relativelylow because only the additional pressure dropof the support has to be overcome.Hence,withrespect to the higher loading rates, the overallenergy demand is in the same range or lowerthan in CSTRs, depending on the support den-sity. Tab. 2 shows some materials which aretested to be used as support for anaerobicfluidized bed systems.All materials have parti-cle diameters significantly lower than 1 mm.By this, the surface for colonization can be in-creased, the superficial upflow velocity can bereduced, and diffusion limitation will play norole, even for porous materials.

Porous materials show the advantage of low-er superficial upflow velocities than non-por-ous materials. Furthermore, biomass gradientsmainly occur in reactors filled with non-porous

materials (ANDERSON et al., 1990; JRDENING,1987). FRANKLIN et al. (1992) reported that asand fluidized bed reactor contained in the bot-tom part only uncovered bare sand. He ex-plains this with the extreme shear forces, whichhave a much bigger influence on bacteria on

the surface. In contrast bacteria in pores areprotected against shear forces and, therefore,such pronounced gradients are not known forthe use of porous materials. In Fig. 5 some por-ous and non-porous supports are shown.

3.3 WastewaterWastewater should be acidified to a high de-

gree (` 80%, related to COD). Otherwise,acidifying bacteria could lower the pH, over-grow the methanogenic biofilm and, hence, re-duce the methanogenic activity. Thus 2-stagesystems are considered superior, since the per-formance in terms of stability and space-time-yield will be superior to 1-stage systems. Theload of reactors with volatile fatty acids (in asystem with a separate acidification reactor)can be higher by factors of 45, compared tofeeding with complex substrates (HENZE andHARREMOES , 1983).

Furthermore, inhibitory substances like sul-fur compounds may play a major role, as inyeast processing (FRIEDMANN and MRKL,

1994). It is also essential that results concern-ing load refer to an average of a stable, contin-uous process rather than to a singular maxi-mum. Even results obtained in the laboratoryreactors differ in general from pilot plant andfull-scale reactors operating at the factory site


Tab. 2. Support Materials for Anaerobic Fluidized Bed Systems

Support Diameter Density Surface Porosity Upflow Biomass ReferenceVelocity[10P 3 m] [kg mP 3 ] [m2 mP 3 ] [m hP 1 ] [kg mP 3 ]

Sand 0.5 2,540 7,100a 0.41 30 420 ANDERSON et al. (1990)Sepiolite 0.53 1,980 20,300b 32 BALAGUER et al. (1992)GAC 0.6 34 CHEN et al. (1995)Biomass granules 26.5 FRANKLIN et al. (1992)Sand 0.10.3 2,600 16 40 HEIJNEN (1985)Biolite 0.30.5 2,000 510 3090 EHLINGER (1994),

HOLST et al. (1997)Pumice 0.250.5 1,950 2.2106 0.85 10 JRDENING (1996),

JRDENING and KSTER(1997)

a Calculated from the given data with the assumption of total sphericity.b Calculated with data given in SANCHEZ et al. (1994).

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with variations in substrate quality and con-centrations and further fluctations.

3.3.1 SolidsStationary Fixed FilmReactors

Suspended solids and even suspended bio-mass may cause clogging in the reactor; this canbe reinforced by extracellular polysaccharidessecreted by acidogenic bacteria (EHLINGER etal., 1987). Therefore, backwash and excesssludge removal must be provided for in the re-actor design. Gas-phase desorption and trans-port through the reactor must be possible.Fixedbed systems are not feasable for wastewaterwith high solids content or components whichtend to precipitate, such as calcium ions.

3.3.2 SolidsFluidized BedReactors

Solids from the wastewater may cause clog-ging, especially at the entrance region of thereactor. HOLST et al. (1997) recommend solidconcentrations lower than 0.5 kg mP 3 with re-spect to problems of clogging in the distribu-tion system, while MSCHE (1998) reports thateven solid concentrations up to 1.7 kg mP 3 didnot cause any problems.This difference can beexplained by differences in the composition of the solids as well as by the construction of thereactor inlet.

Some inorganic compounds, such as calciumcarbonate or ammonium magnesium phos-phate, will be precipitated in the reactor main-ly onto the support, when the actual concen-

Fig. 5ad. Support materials for anaerobic fluidized bed reactors (a and b uncovered and biofilm coveredsand, c and d uncovered and biofilm covered pumice particles).

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trations are beyond the equilibrium.At higherconcentrations of precipitated solids on thesupport diffusion limitation occurs. In suchcases it is necessary to provide the possibilityfor removing this material, and substituting itby uncovered new support. Sand and othermaterials with a high settling velocity in rela-tion to the support have either to be removedbefore entering the reactor or must be re-moved by a device at the bottom of the reactor(as described in Sect. 3.4.2.1).

3.4 Reactor Geometryand Technological Aspects

General aspects of reactor design are dealtwith in Chapters 17 and 18, this volume.

3.4.1 Fixed Bed ReactorsUpflow reactors tend to be favored,because

they allow clumps of biomass to be retained inthe filter by gravity, and the start-up may be

shorter (e.g., 34 months compared to 4 6months for downflow) (ANDREWS, 1988; WEI-LAND et al., 1988). The height is limited by thegradients of the biomass and reaction rate.Thus the fixed bed height was chosen not toexceed 7 m, with an overall reactor height of 12 m, a diameter of 14.5 m for a 1,800 m3 reac-tor, working with distillery effluents at a loadof 4 kg mP 3 dP 1 and 90% conversion.The sub-strate inlet was introduced by 6 inlet devices inorder to obtain a distribution at the reactor in-let with a recycle ratio of 510. The purifiedwater was collected by 12 tubes at the top of the reactor (WEILAND et. al., 1988). Thissystem exhibited good process stability evenfor changes in substrate composition due to di-verse raw materials used.

Fixed bed reactors do not require majorspecific design considerations. The ratio of height and diameter is usually in the range of 12. The inlet must provide equal distributionof the wastewater by means of distribution de-vices. This is in general a system of tubes withnozzles, about one for each 510 m2 (LET-TINGA et al., 1983). The fluid flow should ingeneral be about 1 m hP 1 , up to a maximum of 2 m hP 1 (AUSTERMANN-HAUN et al., 1993).

3.4.2 Fluidized Bed ReactorsFluidized bed reactors are tall in general

and high in relation to agitated tanks or sta-tionary bed reactors. While the heightdiame-ter ratio for laboratory and pilot scale reactorshas a range from 525, the range for technicalplants only varies from 25.Fig. 6 shows a tech-nical plant with 500 m3 volume. The ratio of height and diameter should not be too highwith respect to axial concentration gradientswhich increase with the height of the reactor.But difficulties concerning a uniform fluidiza-tion of the support increase with increasing re-actor diameter (COUDERC , 1985). Therefore, acompromise for both has to be found. The re-actor volume can be calculated as the ratio of the COD load (in terms of kg dP 1) and the vol-umetric loading rate (kg mP 3 dP 1) can be esti-mated from lab and pilot-scale experiments.


Fig. 6. Technical anaerobic fluidized bed reactor(sugar factory Clauen, Germany).

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