BIOLOGICAL FLUIDIZED BED DENITRIFICATION

BIOLOGICAL FLUIDIZED BED DENITRIFICATION

BIOLOGICAL FLUIDIZED BED DENITRIFICATION OF \-/ASTEWATER by Joseph P. Stephenson, B.A.S. A Projet Report Submitted to the Shool of Graduate Studies in Partial Fulfillment of the Requirements for the Degree. Master of Engineering MMaster University Marh 1978

MASTER OF ENGINEERING (1977) (Chemial Engineering) MMaster University Hamilton, Ontario TITLE AUTHOR SUPERVISOR Biologial Fluidized Bed Denitrifiation of Wastewater Joseph P. Stephenson, B.A.S. (University of Waterloo) Dr. K.L. Murphy NUMBER OF PAGES vi i ' 159

ABSTRACT A half-order kineti model (8-48 mg N03+NOz-N/~), oupled with a temperature dependeny desribed by the Arrhenius relationship (4-27 C), adequately desribed biologial denitrifiation of muniipal wastewater in a pilot sale fluidized bed reator. Biofilm support media (ativated arbon or sand) and hydrauli flux (0.25-1.7 m 3 /m 2 min) were not found to be signifiant fators in ontrolling denitrifiation rate within the reator. Control of biofilm thikness on the support media was essential for satisfatory operation of the proess; exess thikness ontributed to elutriation of media and attahed biofilm. Under similar influent wastewater onditions, the fluidized bed proess was apable of equivalent N0 3+N0z-N removal in about one-tenth of the time neessary in a suspended growth or a rotating biologial ontator (RBC) proess. Temperature dependeny of the N0 3 +N0 2 -N removal rate appeared to be less than the dependeny in a suspended growth or a RBC proess, but similar to the dependeny observed in a paked olumn. i i

ACKNOWLEDGEMENTS sinerely aknowledge the supervision and interest of Dr. K.L. Murphy during the experimental program and the preparation of this report. Many with whom I h~ve been fortunate to work have also ontributed to the ompletion of this projet. Dr. B.E. Jank, Head, Biolog~al Proesses Unit, Wastewater Tehnology Centre, provided enouragement throughout the program. Mr. K. Conn, Head, Analytial Servies Setion, and his staff gave exellent support to the requirements of the projet. Mr. D. Ide assisted in the pilot plant operation, sampling, and data ompilation. I also thank my olleagues at the WTC who helped in many ways with questions and problems during the study. The Environmen~al Protetion Servie, Department of Fisheries and the Environment was generous in its support to me and to the projet throughout its duration. I most espeially thank Andrea, my wife, for her patiene, understanding, and enouragement through the entire projet. i i i

ABSTRACT ACKNOWLEDGEMENTS Ll ST OF TABLES LIST OF FIGURES INTRODUCTION TABLE OF CONTENTS Page i i i i i vi vii LITERATURE REVIEW Mirobiologial Aspets Kinetis and Diffusion Temperature Fluidization Bed Expansion and Porosity Fluid Phase Mass Transfer Considerations Flow Patterns in Liquid Fluidized Beds EQUIPMENT AND PROCEDURES Pilot Plant Faility Media Charateristis Hydrauli Flux Sampling and Analysis Fluidized Bed Expansion - Control RESULTS AND DISCUSSION Flow Model Model for Denitrifiation Kinetis Methanol Requirement Solids Prodution Comparison of Fluidized Bed, RBC, and Suspended Growth Reator Denitrifiation Rates Operational Disadvantages 4 4 7 11 13 14 17 20 22 22 26 28 29 30 31 31 36 49 51 CONCLUSIONS RECOMMENDATIONS iv

Page NOMENCLATURE REFERENCES APPENDICES 66 69 76 v

Table Ll ST OF TABLES Ativated Carbon and Sand Charateristis 26 2 Flow Model Experimental Conditions and Results 3 Model for Denitrifiation 4 Disrimination Between Models for Sand Medium 5 Disrimination Between Models for Sand Medium 6 Disrimination Between Models for Ativated Carbon Medium 7 Disrimination Between Models for Pooled Data 8 Solids Inventory, Sand Medium 9 Solids Inventory, Sand Medium 10 Comparison of Fluidized Bed, RBC, and Suspended Growth Reator Denitrifiation Rates 1 1 Comparison of Reation Times from Jeris et at. (1977) with Results from This Work 31 36 37 40 43 45 52 53 57 59 vi

Figure Ll ST OF FIGURES Correlation Between K~p and.nre for Partiulate Fluidization 16 2 Pilot Plant layout 23 3 Traer Response Pattern Using Ativated Carbon Medium Traer Response Pattern Using Sand Medium 5 Arrhenius Temperature Dependeny for Sand Medium ('Vl.O m 3 /m 2 min) 6 Arrhenius Temperature Dependeny for Sand Medium ('Vl.7 m 3 /m 2 min) 7 Arrhenius Temperature Dependeny for Ativated Carbon Medium 32 33 39 42 44 8 Arrhenius Temperature Dependeny for Pooled Data 46 9 Joint Confidene Regions for Arrhenius Parameter Estimates 47 to Relationship of TOC Removal to N0 3 +N0z-N Removal 50 11 Distribution of Volatile Solids Prodution Data 54 12 Dimensionless Relationship Between Conversion and Time for Various Reation Orders 58 vii

INTRODUCTION Man s impat on the balane of nitrogen ompounds in the environment is diffiult to estimate aurately, but it is lear that the impat is signifiant (ILEWPB, 1969; Delwihe, 1970). Considerable attention has been devoted to ontrolling wastewater disharges ontaining nitrogenous ompounds to the environment during the last few years. Nitrifiation for the oxidation of ammonia in wastewater treatment results in the prodution of nitrate and nitrite nitrogen; these produts remain a soure of nutrients in eutrophiation and an be toxi to humans and wildlife. The Canadian Publi Health Assoiation (1969) has set standards at less than 10 mg/i for N03+N02-N in drinking water to avoid infant methaemoglobinemia (oxygen starvation). In areas of water reuse or industrial ontamination, the aumulation of N03+N02-N in drinking water soures may pose a problem. Brown and Mleay (1975) examined the toxiity of nitrite to rainbow trout (Salmo g~dn~), and observed 50% mortality in 96 hours (LC 50) at a onentration of 0.23 mg N02-N/i; they attributed the toxiity to methaemoglobinemia. This projet was aimed at obtaining design information for a new wastewater denitrifiation proess to ombat these problems - the biologial fluidized bed. Suspended growth and supported film biologial proesses have been identified for denitrifiation (Christensen and Harremoes, 1975). Jewell and Cummings (1975) reviewed suspended growth and supported film denitrifiation proess literature and indiated that hydrauli reten

2 tion time was an order of magnitude less for the biofilm proesses (7 min to 8 h versus 2 h to 7 d). Two problems with paked bed biofilm reators assoiated with the higher rates, however, are shortiruiting and high head losses (English et at., 1974; Murphy et al., 1977). Weber (1966) and Weber et at. (1970) reognized the advantages to be gained by using a fluidized granular ativated arbon adsorption proess over a paked bed proess in the presene of suspended solids. The only neessity for bakwashing their expanded bed ontators was to eliminate biofilm buildup on the partile surfae and exess bed expansion owing to the dereased density of the biofilm oated partiles. Beer ( 1970), Beer et at. (1972), and Jeris et at. (1974) used a granular ativated arbon fluidized bed for biologial denitrifiation of wastewater, and were enouraged by high volumetri denitrifiation rates of 2.86 to 4.71 mg N0 3 +N0 2 -N/t:min. Jeris and Owens (1975) observed average rates of 3.72 mg N0 3 +N0 2 -N/t min using sand instead of ativated arbon medium in a biologial flui~ized bed; peak rates of 14.2 mg/t min were reported. Jeris et at. (1977) have also suggested that using fluidized bed tehnology the need for seondary larifiation may be eliminated. In studies with high nitrate wastewaters, Franis and Malone (1975) obtained maximum denitrifiation rates between 13.2 and 19.0 mg N0 3 -N/t min in a benh sale tapered biologial fluidized bed reator using anthraite oat (2-3 mm effetive size) as a media; for the same type of apparatus and wastewater, data from Sott and Hanher (1976) gave a rate of

3 6.42 mg N0 3 -N/~ min based on total vessel retention time. In addition to these rapid denitrifiation rates, these workers have attributed several distint advantages to the biologial fluidized bed denitrifiation onept: high exposed liquid/solid surfae area, onstant head loss, absene of shortiruiting, and potential for elimination of seondary larifiation for sol ids separation and reyle. This pilot sale study was performed at the Wastewater Tehnology Centre, Canada Centre for Inland Waters, Burlington, Ontario during 1975 and 1976 to ontribute information for design of the biologial fluidized bed denitrifiation reator. The projet was aimed at i) defining the denitrifiation of the proess; rate-temperature response ii) defining the response to hydrauli loading and to N0 3 +NOz-N onentration; iii) obtaining a reator flow model; and iv) omparing denitrifiation proesses. rates to rates for other

LITERATURE REVIEW Mirobiologial Aspets This disussion is limited to a brief review of aspets of mirobiologial denitrifiation signifiant to this work. Painter (1970), Payne (1973), Sutton et at. (1974), and Christensen and Harremoes (1975) have published omprehensive reviews onerning mirobiologial denitrifiation. A large number of bateria in sewage are apable of denitrifiation; these bateria are faultative and heterotrophi. In dissimilatory nitrate redution, the N03 ion replaes 02 in the role of an eletron aeptor in the mirobiologial oxidation of a redued oumpound, e.g. organi arbon. Anoxi onditions must prevail for the nitrate redution to be energetially feasible. Depending on ph, gaseous nitrogen end-produts are formed and released to the atmosphere. Requa and Shroeder (1973), Moore and Shroeder 0971), and Stensel et at. (1973) onsidered a single step dissimilative denitrifiation reation to be an adequate representation beause of low N02-N buildup. MCarty et at. (1969), having seleted methanol as the most desirable eletron donor, desribed the stepwise dissimilation reation in the presene of miroorganisms as follows: N03 + 1/3 CH30H = N02 + 1/3 C02 + 2/3 H20 (1) N02 + 1/2 CH30H = 1/2 N2 + 1/2 C02 + 1/2 H20 + OH (2) and overall, - NO; + 5/6 CH30H = 1/2 N2 + 5/6 C02 + 7/6 H20 + OH. (3) 4

5 The above fails to aount for the synthesis of biomass and its organi arbon requirement, whih MCarty et al. (1969) experimentally evaluated in bath and semi-ontinuous benh sale suspended growth experiments. After inluding deoxygenation, they arrived at ~he following semi-empirial relationship for methanol (CH 30H) onsumption: where all terms are in onentration units. (4) For biomass synthesis, using the empirial formula CsH102N for ells, the relationship (5) was obtained. Tamblyn and Sword (1969) onfirmed MCarty's formulation for methanol requirement in fixed-film anaerobi filters using a variety of media for denitrifiation of agriultural drainage water with about 20 mg/ N03-N. Based on large sale (1140 m 3 /d) pilot data, English et al. (1974) arrived at a slightly different formulation for methanol requirement: (6) They onluded that arbon for synthesis was not entirely obtained from the methanol, but from other soures suh as organis in the effluent although the mehanism was unknown. Through onsiderations of ell yields and negleting N02-N transformations, Stensel et al. (1973) onluded that methanol requirements ould be estimated by: ~CH30H = 2.3~N03-N + l.o~oz. (7) Other arbon soures have been identified (Christensen and Harremoes, 1975; Monteith, 1978) but methanol remains popular owing to its widespread availability, low ost, and reported low yield.

6 Jeris and Owens (1975) and Jeris ~a. (1977) have reported that a CH30H/N0 3 -N mass ratio of 2.9 to 3.0 is suffiient to ahieve essentially omplete denitrifiation in a biologial fluidized-bed proess. Christensen and Harremoes (19i5) have summarized the methanol onsumption in experiments by four other authors using fixedfilm reators, and onluded that a CH30H/N0 3 -N ratio greater than 3.0 is neessary for more than 90% denitrifiation. It appears that essentially omplete denitrifiation is ahieved beyond this ratio. Sutton ~ al. (1974), Wilson (1975), and Soyupak (1976) have reported satisfatory denitrifiation rates, not limited by available arbon as eletron donor, in their experimental work with paked olumns and a submerged rotating biologial ontator (RBC) operated at the Wastewater Tehnology Centre with CH 3 0H/N0 3 -N ratios greater than 2. 7.

7 Kinetis and Diffusion Denitrifiation rates have been desribed using the Monod relationship for substrate removal: {8) where: C is the substrate onentration, Cb is the biomass onentration, t is the time, k m Kk is the maximum substrate removal rate, and is the value of C at whih k = km/2. Many workers (Moore and Shroeder, 1971; Requa and Shroeder, 1973; Stensel eta!., 1973; Dawson and Murphy, 1972; Sutton eta!., 1974; Sutton et at., 1977) have shown that the Monod relationship an be simplified to a zero-order expression with respet to N0 3-N onentration in suspended growth proesses, (dc/dt)/cb = -k (9) m This simplifiation has been determined to hold for N03-N onentrations to less than about 1 mg/~. Murphy et at. {1977) reported that nitrate removal in paked olumns and a submerged RBC with fixed biofilms was also adequately desribed by zero-order reation kinetis. Others (Smith et at., 1972; Parker et al., 1976; Jeris et at., 1974) have shown that olumnar nitrate reations were not linear (zero-order), and derived first-order kineti models from their data. Parker et al. (1976), and Jeris et at. (1974) attributed the observation of first-order behaviour to nonuni

8 form biomass distributions in their reators. Christensen and Harremoes (1975), Harremoes (1976), and Harremoes and Riemer (1975) have explained non zero-order behaviour in the ontext of diffusion limitations in fixed-film reators for their own and other authors data. Christensen and Harremoes (1975), and Harremoes (1976), demonstrated that Sutton s (1973) paked olumn data showing zero-order denitrifiation kinetis ould also be desribed by a zero-order biohemial reation influened by biofilm pore diffusion (or homogeneous biofilm diffusion). The result of their analysis of biofilm diffusion with simultaneous reation is a bulk halforder model through the reator for a range l~[n03-n]<40 mg/2. For the half-order model to hold, the substrate must be onsumed before it has diffused through the omplete biofilm thikness. For omplete biofilm penetration or for nitrate-n greater than 40 mg/2 the reation would be zero-order. For first-order kinetis with or without omplete penetration a first-order model would apply. The diffusion theory with partly effetive pores predits that substrate removal is independent of biomass onentration, but dependent on exposed surfae area in the reator. Other investigators (Williamson and MCarty, 1976a; Williamson and MCarty, 1976b; Mueller, 1976; LaMotta, 1976a; LaMotta, 1976b; Famularo et al., 1976) have also examined the role of diffusion from the bulk liquid to the biofilm surfae and within the biofilm. These authors have onluded that intrinsi zero-order reations are transformed to a bulk half-order model in partially penetrated biofilms, and remain zero-order reations in ompletely penetrated biofilms. LaMotta (1976a)

9 stated that the observed reation order of an n-th order reation signifiantly affeted by pore diffusion would be: Observed order = (n + 1)/2 (10) where: n is the intrinsi reation order. Using gluose removal data from Kornegay and Andrews (1969), he demonstrated that the half-order onept was a satisfatory explanation for biofilms deeper than some limiting thikness typially on the order of 10 to 100 ~m; beyond this 1imiting thikness, removal rate was independent of film thikness and biomass onentration. Jennings (1975) examined the kinetis of aerobi gluose removal in a granular ativated arbon fluidized bed assuming zero-order and first-order kinetis, and onluded that there was a limiting biolayer thikness beyond whih removal was a funtion of exposed surfae area, not biomass onentration within the expanded bed. This suggested that mass transfer (diffusion) was important, but Jennings did not attempt to fit half-order kinetis to this data. Ottengraf (1977) pointed out the omission of the diffusion dependeny in his review of Jenning's results. Williamson and MCarty (1976a) have built a mehanisti model desribing substrate uptake in deep biofilms uniting diffusion through a stagnant liquid surfae layer, biofilm diffusion, and reation. For substrate onentrations muh greater than the Monod half-veloity oeffiient, they indiate that substrate biofilm flux is a funtion of substrate onentration to the one-half power. One ondition for their model to hold was that either the eletron donor or the eletron aeptor must approah a near zero onentration within the biofilm and be the limiting speies throughout the region of reation. William

10 son and MCarty (1976a) and Harremoes and Riemer (1975) have also evaluated eletron donor requirements for the denitrifiation reation in biofilms, and suggested that the donor may be flux 1imiting if only added in stoihiometri proportion to the nitrate. For methanol, Williamson and MCarty (1976a) predited that feed CH30H onentration should be five times greater than feed N0 3-N to avoid methanol flux limitation. Harremoes and Riemer (1975) suggested that methanol as eletron donor for denitrifiation might be flux limiting at the inlet of biofilm reators if added in insuffiient quantities and, to avoid methanol limitation, a ratio CH 30H/N0 3-N greater than 3.6 was neessary. Mueller (1976) and Famularo et af. (1976) have quoted analyses by Mueller in whih the half-order model was shown useful in desribing biofilm reations for aerobi systems with partial penetration of the films (anoxi ore model). It is apparent from the works disussed above that diffusion may exert a signifiant role in kineti haraterization of biofilm reations.

1J Temperature Temperature is an important influene on the speifi substrate removal rate onstant, k. To inlude the effet of temperature in kineti rate equations, the speifi rate onstant, k, for any order reation has been expressed by the empirial Arrhenius relationship: k = 1: -l:ie/(rt) A exp ( 11 ) where A * is the preexponential, frequeny fator,!:ie is the ativation energy, R is the universal gas onstant, and T is the absolute temperature. Himmelblau (1970) reparameterized the Arrhenius relationship to re.j. due orrelation between An and!:ie during parameter estimation: k = A exp-l:ie(l/t - 1/To)/R (l 2 ) * -l:ie/rto A = A exp ( 13) where To is the entral absolute temperature, i.e. mean experimental temperature. Christensen and Harremoes (1975) tabulated a list of workers who have suessfully fit the Arrhenius temperature relationship to the denitrifiation reation. In partiular, Dawson and Murphy (1972) used a dominant ulture of P-6e.udomon.a.6 de.n.ilij.m.a.n.-6 and showed that bath denitrifiation ould be desribed by the Arrhenius relationship. In pilot sale studies, Sutton (1973) and Soyupak (1976) suessfully fit the reparameterized Arrhenius model to zero-order denitrifiation rate data in fixed film reators over temperature ranges of approxi

12 mately 5 to 25 C. Parker et al. (1976) modelled first-order denitrifiation rate data for benh sale paked beds and found the Arrhenius equation to be adequate in the range 5 to 20 C. The ativation energy, ~E. an be used as a measure of whether the reation is biohemially rate ontrolled or diffusion ontrolled (Bush, 1971). Bush indiated that diffusion proesses exert inreasing ontrol as ~E dereases from a reation ontrolled value. For diffusion ontrol, data from Bush (1971), and Johnstone and Thring (1957) indiate that ~E is less than about 5000 al/g-mol. Thereation ontrolled value of ~E is usually in the range 11,000 to 20,000 al/g-mol. Haug and MCarty (1971) predited that deep biofilm systems would give less temperature response beause of diffusion limitations within the film. Johnstone and Thring (1957) stated that a mixed hemial-diffusion ontrolled proess ould exist, even with an ativation energy whih denotes hemial ontrolled resistane.

13 Fluidization Fluidization is a unit proess in whih a fluid, suh as water or air, is passed upward through a bed of partiles with a veloity suffiient to expand the bed. As fluid veloity inreases from the initial fluidization veloity, the bed expands and bed porosity inreases. Below the terminal settling veloity of individual partiles, the partiles are retained in the bed, but are in onstant, random motion. Fluidization in water and wastewater treatment systerns has, in the past, been assoiated with sedimentation and filter bakwashing (Amirtharajah and Cleasby, 1972; Garside and Al-Dibouni, 1977; Weber, 1972). Fundamentals of fluidization have been doumented by Leva (1959), Zenz and Othmer (1960), and Kunii and Levenspiel (1969). In hemial proess engineering many advantages have been ited (Leva, 1959; Kunii and Levenspiel, 1969) for fluidized bed ontators: i) high heat and mass transfer rates ompared to fixed surfae ontators, ii) relatively uniform solids distributions, iii) less pressure drop than paked bed systems, iv) easy transport of solids to and from the reator, beause fluidized solids behave muh like a fluid, v) isothermal reator onditions, vi) vii) suited to large sale operations, and high surfae area per unit reator volume. Disadvantages that should be onsidered are solids elutriation, equipment wear from solids abrasion, pressure drop equal to the mass of the media, and ontribution to bakmixing from solids irulation.

14 Bed Expansion and Porosity Bed height and porosity (voidage) in a fluidized bed are diretly related through the definition: { 14) Inreases in fluid veloity ause the bed height and the porosity or void spae, between the medium partiles to inrease simultaneously. Minimum porosity ours for the bed in the paked state; for granular materials like sand and ativated arbon with partile size of about 1 mm the minimum porosity is _near a value of 0.4. Fluidization and expansion from the minimum porosity our when the pressure drop through the bed equals the gravitational fore exerted on the partiles in the fluid. After inipient fluidization, there is negligible hange in the bed pressure drop, ~p, for a partiulately fluidized bed, and the relationship {Leva, 1959): ~p' = L(g'lg )(p p - p)(l -E) (15) an be used to predit head loss through the bed. Other definitions an also be found in the literature. Leva (1959) desribed several orrelations for prediting bed expansion. In this projet, an effort was made to use the orrelation by Wilhelm and Kwauk (1948) to estimate hydrauli flux settings for speified degrees of bed expansion. Their orrelation was developed using several media inluding sand, with air and water as fluid. They orrelated their data by plotting the produt of the drag oeffiient, 0, times the square of the partile Reynolds number, N~e' against the partile Reynoldi number, NRe (i.e., 0 N~e versus NRe) for values of porosity ranging from 0.40

15 to 1.0 (Wilhelm and Kwauk, 1948; Leva, 1959). N~e has been de 0 noted by the dimensionless quantity, K~p Figure 1 is a plot of their orrelation where: K~p = d~ pfg'(ps-pf)/(2~2) {16) NRe = dpvpf/~ (I 7) With Figure 1 only the partile and fluid harateristis are required. To obtain the fluidization veloity, v, for a given expanded bed porosity,, K~p an be alulated from known partile and fluid harateristis. With K~P and E speified, the orresponding NRe may be obtained from the absia of Figure 1. The required veloity is alulated by rearranging equation (17), (18)

16 E= 40%50 60 70 8090100!I!Ill/ /IIIII., FIGURE 1. CORRELATION BETWEEN K 6 p AND NRe FOR PARTICULATE FLUIDIZATION (Wilhelm and Kwauk, 1948)

17 Fluid Phase Mass Transfer Considerations One reported advantage gained by using fluidized beds over fixed or paked beds is the high mass transfer rate between fluid and solid. The mass transfer oeffiient for a fluidized bed is shown to be less than the oeffiient for a fixed bed based on observations summarized by Kunii and Levenspiel (1969) and Leva (1959). The large surfae area exposed for transfer is the fator responsible for the high mass transfer rate. For Reynolds numbers in a range somewhat higher(>50) than those experiened in this work, Mullin and Treleaven (1962) onluded that intensity of turbulene was related to mass transfer rates in multipartiulate systems. They indiate that intensity of turbulene inreases with dereasing bed voidage and with inreasing Reynolds number. Dwivedi and Upadhyay (1977) reviewed and reanalyzed mass transfer data in fixed and fluidized beds, and onluded that the mass transfer fator was inversely proportional to bed voidage over a broad range of Reynolds number. These observations suggest that a fluidized bed reation, 1imited by fluid-partile mass transfer, should be performed with a bed at minimum pratial voidage and maximum partile Reynolds number. In fixed-film biologial treatment systems many investigators have demonstrated that fluid-solid mass transfer was a signifiant fator in treatment effiieny. Hartmann (1967) used a tubular, benh sale, fixed-film reator and syntheti substrate (Liebig's meat extrat) to study the effet of fluid veloity on reation rate; he onluded that inreasing turbulene with inreasing veloity was respon

18 sible for higher substrate removal rates. Gulevih et al. (19G8) also showed that inreased rotational disk veloity inreased gluose substrate flux to the disk surfae in an aerated benh sale reator, and onluded that external diffusion dependeny must be onsidered in wastewater treatment design. Williamson and MCarty (l97ga, 197Gb) developed a biofilm model with diffusion and reation onsiderations; an external stagnant liquid layer was inluded in the model beause they hypothesized that turbulene and negligible liquid layers were rarely ahieved in biofilm reators. They stated that 1iquid films greater than 30 ~m an be expeted, and they alulated a minimum stagnant 1iquid layer thikness of 56 ~m at a biofilm surfae in their experimental apparatus. The 1iquid film thikness was a funtion of relative veloity between the liquid and biofilm surfae. In developing a mathematial model for biologial ativity in expanded bed adsorption, Jennings (1975) stated that a fluid boundary layer is inonsequential for zero-order reation kinetis. If pore diffusion is an important variable, then this layer may have a signifiant influene on the driving fore for pore diffusion even if zeroorder reation kinetis predominate. In the extreme, fluid boundarylayer diffusion may be the rate ontrolling step, and lead to a bulk first-order model (LaMotta, 197Gb). LaMotta (197Gb) evaluated substrate diffusion to a biofilm surfae and found that fluid veloity was an important variable in the observed uptake rate. Below a ritial veloity, he onluded that external diffusional resistanes influened apparent kineti measurements and, in ~he diffusion ontrolled

19 regime, led to a bulk first-order model. Harremoes and Riemer (1975), studying denitrifiation in granular media paked bed reators, found that the effet of flow variations (laminar flow) on liquid film diffusion was unimportant beause the 1iquid film had no effet on removal rate. They attributed the absene of an effet to the tortuosity and mixing in the flow paths. Beause hydrauli flux simultaneously influenes bed expansion and Reynolds number, both of whih are important mass transfer variables, an effort was made to examine the influene of hydrauli flux on the overall removal of N0 3 +N02-N in a portion of this fluidized bed study.

20 Flow Patterns in Liquid Fluidized Beds For hemial and physial proesses whose rates are onentration dependent, the residene time distribution of the reatants must be onsidered. Many investigators have delineated the hydraulies of fluidized beds; in general, liquid fluidized beds have been desribed using a dispersion oeffii~nt for axial flow (Leva, 1959; Kunii and Levenspiel, 1969; Bruinzee1 et al., 1962; Kramers et al., 1962; Potter, 1971). Leva (1959) suggested that for mass transfer studies axial fluid mixing is negligible. He also presented a dimensional orrelation for dispersion in liquid fluidized systems: where: DL MS = 0.33 v,0. p (19) DL is the diffusivity oeffiient, ft 2 /h, v is the superfiial fluid veloity, ft/h, and D is the partile diameter, ft. p Substitution of representative values of D (~lmm) and v (~1m/min) p for the fluidized bed used in this study indiated only a small amount of dispersion (D/~L < 10-3 ) (Levenspiel, 1962). Plug flow onditions adequately desribe the hydraulis of the system under these irumstanes. Kunii and Levenspiel (1969) state that the height of an ideal mixing stage ranges from several partile diameters to O.lm for liquid fluidized beds. For a tall olumnar reator the number of ideal mixing stages will beome large and progress toward plug flow. An examination of mixing in a sand-water

21 fluidized bed by Bruinzeel et al. (1962} revealed relatively low values for dispersion for the partile and fluid harateristis of this study. Kramers et at. (1962} also provided evidene that the effet of axial dispersion on kinetis would be negligible in liquid fluidized beds from a pratial point of view, and that a bed of about 1 m in height would ontain more than 100 mixing stages (CSTR's). From the graphial orrelation presented by Potter (1971} a dispersion oeffiient an be alulated showing an intermediate to small amount of dispersion in a liquid fluidized bed of typial height (~3m}. It is apparent that plug flow hydraulis are often an adequate desription of the mixing in 1iquid fluidized beds in wastewater treatment appliations. The traer studies by Beer et at. (1972} for a biologial fluidized bed denitrifiation reator indiate by inspetion that plug flow hydraulis predominate the residene time distribution.

EQUIPMENT AND PROCEDURES Pilot Plant Faility An extended aeration pakage plant loated at the Wastewater Tehnology Centre in Burlington was used to produe a ompletely nitrified effluent to feed the fluidized bed denitrifiation reator. Degritted raw sewage from the Burlington Skyway WPCP was fed to the pakage plant at a onstant flow rate, but with diurnal flutuations in nitrogen onentration. A ompletely mixed temperature ontrol tank ooled or heated the larified pakage plant effluent to the desired run temperature. When neessary, a KN03 solution, pumped into the ooler tank using a variable speed pump, furnished a residual N0 3 -N onentration greater than 1.0 mg/~ in the reator effluent. KN03 solution was also used to extend the range of feed N03-N onentrations for kineti measurements. A methanol solution, supplying arbon and energy for the denitrifiation reation, was pumped into the sution of the fluidized bed feed pump using a positive displaement variable speed hemial feed pump. The variable speed positive displaement reator feed pump pumped the nitrified effluent and hemials diretly into the olumn base. Figure 2 depits the pilot plant layout. The ylindrial, aryli olumn reator onsisted of setions 1.22m long and O.J4m inside diameter. Setions were onneted with a flange and gasket. Initially, three setions were used for a total reator height of 3.66m; a fourth setion was added later in the study (16 Deember 1975). The PVC flow distribution plate had forty-nine 22

23 Not to Sale Mehanial Sieve Nitrified Effluent From Pakage Plant Solids Removal Effluent KN0 3 Addition CH 3 0H -------~~ Addition Temperature Controller Media Plus Bio Film l 1.22 m *Sampling Point Feed Pump Steel Base * FIGURE 2. PILOT PLANT LAYOUT

24 2.38 mm holes symmetrially arranged around the plate's entre to a maximum distane of 60 mm from the entre. This left approximately a 20 mm unperforated border near the olumn walls to avoid exessive abrasion. Flow entered the 0.15 m long entrane setion below the distribution plate through four 9.53 mm holes drilled at 90 degree angles on a 19.05 mm pipe ap. On 8 January 1976 the entrane setion and base of the olumn were hanged to steel pipe onstrution beause of pressure buildup below the distribution plate and exessive abrasion of the aryli walls above the distribution plate. The steel base had a 0.15 m inside diameter, and was 0.46 m long above the distribution plate and 0.15 m long below the distribution plate. A mehanial mixer, plaed at the top of the olumn, was used in an attempt to shear biofilm and attahed gas bubbles from the media rising to the surfae to prevent partile arry over in the effluent. Initially, a three-blade propeller was used; this was replaed (16 Marh 1976) by a flat four-blade impeller with zero pith. The threeblade propeller sheared gas bubbles from the partiles, but did not shear exess biofilm or ontrol bed expansion due to growth. The flat four-blade impeller did shear bubbles and biofilm from rising partiles and was effetive in ontrolling bed expansion. Jeris and Owens (1975) reported suess using a similar devie in a 0.46 mdiameter olumn. Portions of the fluidized bed did not have to be removed to ontrol expansion after the flat four-blade impeller was installed; instead solids were ontinuously wasted in the effluent. Denitrified effluent was disharged by gravity and flowed

25 through a 0.19m diameter sieve with 1.19mm diameter openings. The sieve effetively aptured any biofilm oated media existing in the olumn disharge. This media plus biofilm was ombined with portions of the bed whih were removed to ontrol expansion and measured as waste solids.

26 Media Charateristis Charateristis of the ativated arbon and sand media used in this work are listed in Table 1. For the first three months of operation, granular ativated arbon was the medium used to support biofilm development. To remove arbon fines before biofilm development, the medium was bakwashed with tap water. Start-up was arried out beginning with an unexpanded bed height of one-third to one-half total olumn height. TABLE 1. ACTIVATED CARBON AND SAND CHARACTERISTICS Ativated Carbon * Sand~'d Effetive Size (mm) 0.55-0.65 0.75-0.85 Mean Size (mm) mass basis number basis 1.0 1.1 Uniformity Coeffiient Wetted Partile Density (g/m 3 ) Unexpanded Bed Porosity 1.9 1.4 1.3-1.4 2.65 0.40 0.41 * Calgon Corp., Pittsburgh, Pennsylvania, Filtrasorb 400 **William R. Barnes Co. Ltd., Waterdown, Ontario, #1 Sandblast Sand White quartzite sand medium replaed the ativated arbon medium (3 Deember 1975) after operating diffiulties were enountered with the arbon. Physial sand harateristis are given in Table 1. The sand was also bakwashed with tapwater to remove fines whih lassified at the surfae of the expanded bed. After startup new unwashed ativated arbon or sand were periodially harged to the bed as makeup.

27 After biofilm developed on the media, media plus biofilm ~ diameters were measured using a Zera-Meter mirosope~ with a builtin mirometer. Approximately 40 partiles were randomly seleted and measured to represent the size distribution. For the experimental runs, the number-weighted arithmeti mean partile size, d' was used to set the hydrauli flux to ahieve a speified p,so' expanded bed porosity. ~ ~Zeta Meter In., New York City.

28 Hydrauli Flux Hydrauli loadings were seleted using the orrelation from Wilhelm and Kwauk (1948). As partile size had been established, it was neessary only to estimate partile density. This was done as fo 11 ows: i) media and media plus biofilm were assumed to be spherial, ii) density of the media was taken from Table 1; ativated arbon density was taken as 1.35 g/m 3, iii) biofilm density was approximated as the density of water (1.0 g/m 3 ), iv) With the d~ p,so biofilm thikness was half the measured differene between the mediad and the media plus biofilm d", and p,so p,so v) overall partile density, pb' was obtained by alulating the media plus biofilm partile mass and dividing by the partile volume. measured and pb alulated, the dimensionless number K~p (Equation 16) an be alulated. For a speified expanded bed porosity, E, the partile Reynolds number, NRe' is obtained from Figure 1. The required hydrauli flux, v, is obtained from the Reynolds number and fluid and partile properties: (20) Q = v (olumn ross-setional area). (21) = v K

29 Sampling and Analysis On run days feed samples were olleted from the temperature ontrol tank using refrigerated 24-hour omposite samplers. With few exeptions, eah run lasted 24 hours and samples were olleted hourly. One grab sample was taken from the olumn base on run days for TOC analysis after methanol addition. Twenty-fourhour omposite effluent samples were taken diretly from the overflow of the olumn in a fashion similar to the influent sampling. Analyses inluded unfiltered TKN; filtered NH + 4 -N, N0 2 -N, N03-N, TKN, and TOC. filtered using 0.45 ~m Samples for analytial determinations were Gelman glass fibre filters. The Laboratory Servies Setion at the WTC performed these analyses, and a desription of the details is available (Laboratory Servies Setion, 1976). Sutton (1976) and Wilson (1975) have also summarized the analytial methods. Column feed and effluent suspended solids were also measured using 0.45 ~m Gelman glass fibre filters. Filter papers were ovendried before use at 100 to 105 C; filter papers and deposited solids were dried at the same temperature for at least 24 hours befor~ weighing. Dissolved oxygen (DO) was measured using a YSI model 54 oxygen meter. DO was measured in the temperature ontrol tank and at the top of the olumn above the fluidized bed- liquid interfae. An Orion Model 401 speifi ion meter was used to determine ph. Temperature was taken with a merury thermometer.

30 Fluidized Bed Expansion ~ Control Entrainment of the expanded bed in the effluent proved to be a signifiant problem. Inreasing biofilm thikness on individual media partiles and attahed gas bubbles resulted in a derease in bulk partile density to the density of the fluid medium allowing the bed to be transported. To ontrol this problem, the mehanial mixer devies desribed previously were installed in the olumn. The sharpened blade with zero pith was very effetive in ontrol) ing partile size and bed expansion; the three-blade propeller was less suessful. With the three-blade propeller, upper portions of the bed had to be removed daily, and stripped of attahed biomass external to the olumn. The portions were removed into a two-gallon pail and agitated for 15 to 30 minutes with a mehanial stirrer. This agita ' tion effetively stripped the gelatinous biofilm from the media and produed a fluid phase similar to onentrated mixed liquor (~1% solids) above the settled media in the pail. The mass and sol ids onentration of the wasted liquor were routinely measured. The media was reyled to the olumn. Though this rudimentary ontrol method was not ideal, it was effetive. It was more ritial to ontrol biofilm thikness on ativated arbon than on sand beause of the arbon's lesser den-1 sity.

RESULTS AND DISCUSSION Flow Model Two investigations to desribe flow patterns for the ativated arbon medium and for the sand medium were onduted using fluoride (F-) salt solution and a F- speifi ion eletrode. A hydrauli flux of about 0.26 m 3 /m 2 min for the ativated arbon and 1.5 m 3 /m 2 min for the sand medium was used. A method reported by Timpany (1966) was used to estimate the dispersion oeffiient, 0/~L. A tanks in series model was fitted to the data using the method desribed by Levenspiel (1962). For details, see Wilson (1975). Table 2 summarizes the experimental onditions and results (Appendies A.l and B), and Figures 3 and 4 provide a omparison of the traer data with the theoretial dispersion model in dimensionless oordinates. The dispersion model was a better approximation to the data than the tanks in series model. TABLE 2. FLOW MODEL EXPERIMENTAL CONDITIONS AND RESULTS Hydrauli Flux (m 3 /m 2 min) Reator void volume (R-) Reovery (%) D/11L T, theoretial retention time (min) 7.31 T, mean retention time (min) Ativated _ Carbon 13/11/75 0.26 28.5 73% 0.061 8.73 Sand 9/4/76 1.5 45. 113% 0.036 2.04 3. 13 31

2.0~--~--~~--~--~----,---~----,---~----~--~----~---r----~---r--~..::: 1.2 0 Theoretial Dispersion Model w (D/~L =0.061) w ""'' 1.0 0.8 Hydrauli Flux= 0.26 m 3 /m 2 min 0 _ 0 1 ' L,,A- ',,, r,, -, ~ -- - r= - r=-==- 1 0.0 0.2 0.4 0.6 0.8 1.0 1.2 1.4 1.6 1.8 2.0 2.2 2.4 2.6 2.8 3.0 t/t FIGURE 3. TRACER RESPONSE PATTERN USING ACTIVATED CARBON MEDIUM w N

2.0.----.----.----.----.---~----.----.----.----,r----r----r----r----T---~--~ 1.8 1.6 ~ Theoretial Dispersion Model 0 ~ (D/J.!L=0.036) w -' w o.o "'-. Hydrauli Flux= 1.5 mo/m 2 min."'., ~ --- ---- -- - el-e 0.0. 1.0 2.0 3.0 t/t FIGURE 4. TRACER RESPONSE PATTERN USING SAND MEDIUM w

34 A portion of the medium, about 1 metre, at the reator inlet failed to aumulate a biofilm and was signifiantly more turbulent than the seeded medium above. A distint interfae separated the two zones at all times. Based on qualitative observation of a visible dye (Rhodamine WT), a large amount of dye dispersion was experiened in the poorly fluidized bottom setion. This aounted for a signifiant but unmeasured fration of the overall dispersion predited by the traer results (Table 2). Turbulent eddies were less apparent in the smoothly fluidized seeded zone, and the dye front moved through the bed in plug-like flow. Visual inspetion of the Rhodamine WT dye intensity above the bed refleted the F- data in Figures 3 and 4 very well. The exhange of the traer between the bulk fluid and the biofilm aused deviations from the theoretial urve after the peak value. Harremoes and Riemer (1975) disussed dispersion in their paked bed denitrifiation olumns and suggested the like! ihood of pore diffusion. Beer et al. (1972) obtained traer responses similar to those shown in Figures 3 and 4 for their ativated arbon fluidized beds. A omparison of observed mean and theoretial retention times (Table 2) shows the influene of pore diffusion in the reator with the observed mean greater for eah ase. From the shapes of the traer responses, no shortiruiting or dead spae were evident in the reator; nor did hydrauli flux appear to markedly effet flow pattern, though about two times as muh dispersion existed in the ativated arbon bed (Table 2). Beause of nonideal entrane onditions, pore diffusion, and b'ased on assessment in the I iterature of flow onditions in 1iquid

35 fluidized beds, simplifiation to an ideal plug flow model was onsidered justifiable for kineti determinations. Jennings (1975) also justified this assumption in his modelling of a benh sale biologial fluidized bed proess.

Model for Denitrifiation Kinetis Data were obtained using sand and ativated arbon as biofilm support media (Appendix C).. Using sand media, two levels of h~drauli flux, ~1.0 m 3 /m 2 min and ~1.7 m 3 /m 2 min, were investigated to ahieve approximate bed porosities of about 0.5 and 0.6, respetively. Beause of start-up and operational problems assoiated with exess biofilm aumulation, only six experimental runs were ompleted during the use of ativated arbon media. With ativated arbon media, a hydrauli flux of about 0.25 m 3 /m 2 min was used. Wastewater temperature was ontrolled in the range 4 to 2] C. To assess the data, it was assumed that zero, half, and first-order models were potentially valid desriptions for denitrifiation in the fluidized bed reator; the Arrhenius relationship was used to desribe the temperature dependeny of the rate onstants. The model forms presented in Table 3 were obtained from the following expressions: dc/dt = ken (22) -~E(l/T - 1/T 0 )/R k = A exp { 12) TABLE 3. MODEL FOR DEN ITR IF ICAT ION / Model Reation Order, n Integrated Model Form 1 2 3 1 t 0 -k t = C exp 1 0 = - Ctk 1 t + 0 0 2 = - k t 0 0 (ktt/2) 2 2 {23) (24) (25) The parameters, A and ~E, for eah model form were estimated

37 from the data for eah flux setting separately by means of a nonlinear least squares tehnique whih is summarized in Appendix A.2. To determine whih of the three kineti models best represented the data, the Bayesian model disrimination method desribed by Reilly {1970) and Reilly and Blau (1974) was used (Appendix A.3). The data by Jeris e;t U.. (1977) and Sutton e;t U.. (1974) were used to estimate prior parameter distributions for A and ~E (Appendix D.2); equal prior probabilities were initially assigned to eah model. For the low hydrauli flux with sand media, Arrhenius parameters were estimated for eah model alternative for sixteen runs, and then used to determine the model likelihoods. Examples of the estimation and disrimination proedures are given in Appendies D.l and D.2. Table 4 summarizes the probabilities of eah model; the half-order model has the highest posterior probability (0.9989) and, between these models, is preferred as the most valid desription'of the kineti denitrifiation data. TABLE 4. DISCRIMINATION BETWEEN MODELS FOR SAND MEDIUMt Model j Reation Order Prior Probability L (M )* Prior j j Probability Posterior Probability 1 1 1/3 0. 6125,' 10-2 0 0.0010 2 t 1/3 0.61001:10-17 0.9989 3 0 1/3 o.2163''~1o- 21 0.0001 1 0000 2. 0360,' 1 o- 18 1. 0000 Good separation between the models was obtained, and there is greater than a 99% hane that the half-order model is orret. No signifi

38 ant improvement in fit over the half-order model was experiened by simultaneously estimating the reation order, n, by nonlinear least squares tehniques. The half-order model showed no evidene of lak of fit, no systemati deviations in the residuals were evident, and examination of the residuals indiated a onstant variane. An example of the alulations and methods for heking model adequay is given in Appendix 0.3. To represent the data graphial ly\using the Arrhenius relationship for the half-order model, the rate onstants were alulated for eah run, 1 k = 2(C! - 1 C~)/t (26) and plotted against temperature (Figure 5). The best fit 1ine was drawn using the parameters, A and ~E, previously derived from the data (Appendix D.l). The ativation energy of 11860 al/g-mol is about the same as the value 11090 al/g-mol reported by Sutton et al. {1974) for denitrifiation within a pilot sale paked olumn, but signifiantly less than their value of 15900 al/g-mol for suspended growth reators operated at SRT 1 s of six days. This indiates that fluidized bed denitrifiation may be less temperature dependent than denitrifiation in the ativated sludge proess. Data for the 1.7 m 3 /m 2 min flux setting with the sand medium were interpreted in a manner similar to that above. The three models {Table 3) were fit to estimate the Arrhenius parameters, A and ~E, and the Bayesian disrimination approah was used to test model preferene {Table 5). The data for this flux exhibited about

39 It... T"" I 2.0 r------,------,----r-----r----.-----. / T T+273 k = _ 68 exp -11860(1/T~ 1/288)/R 0 : I ~ /. / /~ :: ~ 8/ / E / / A / /./.. ~~ ~..,.,- / /' / / / / --;...-- ~ 95.% Limits ~ --------~.,- --- 0.0L------L------'----'----....L ---J 0 5 10 15 20 25 30 T( C) FIGURE 5. ARRHENIUS TEMPERATURE DEPENDENCY FOR SAND, MEDIUM (--1.0 m 3 /m 2 min}

40 TABLE 5. DISCRIMINATION BETWEEN MODELS FOR SAND MEDIUMt Model j Reation Order, n Prior Probability L (M )~~. Prior j j Probability Posterior Probab i 1 i ty I I 1/3 o.6093~qo- 17 0.6992 2 t 1/3 o. 25981:1 o- 1 7 0.2982 3 0 1/3 0.2238*10-19 0.0026 1. 0000 o.8714~':jo- 17 1.0000 three times more variability than the data for the 1.0 m 3 /m 2 min setting and, having assumed equal prior probabilities for eah model, the separation between the first-order and half-order models was inadequate. With equal prior probabilities, the probability of the first-order model being preferred is about twie as great as the probability for the half-order model; this amount of separation is not.. ' suffiient to differentiate between these two models learly.! If the equal prior probabilities, equated to a value of 1/3, are updated by the additional knowledge gained from the pos~erior probabilities and parameter distributions derived from the low flux runs (Table 4), then the half-order model would be preferred. Lak of fit was not deteted for the half-order model either for the pure error variane assoiated with this experiment or for the pure error variane assoiated with the experiment at the 1.0 m 3 /m 2 min hydrauli flux. The zero-order model was the least probable model alternative. For omparison to Figure 5, the half-order rate onstants were

41 alulated using equation (26) and plotted against temperature (Figure 6) to show the Arrhenius relationship for the 1.7 m 3 /m 2 min hydrauli flux data. This plot learly shows the larger variability in the data at this hydrauli flux. Although the temperature relationship is statistially adequate, it does not provide a well defined desription of the rate onstants. To remove the effet of temperature when omparing the two sets of runs at the different fluxes using sand medium, the rate onstants were ompared for a temperature of 15 C. At 15 C the rate onstants are equal to the Arrhenius preexponentials, A: at t -~ 2-1 k z 0. 82 mg Q. min at The variability in the data shown in Figures 5 and 6 suggests that the differene in the preexponentials is not signifiant. It is possible that the proess had not reahed steady-state during the high hydrauli flux runs; examination of the sol ids data (see Sol ids Prodution)showed -that the solids wastage may have been more variable during these runs. To demonstrate the degree of denitrifiation observed using the ativated arbon media and to ompare this to the sand media, the available denitrifiation rate data were evaluated. Arrhenius parameters were estimated using the nonlinear least squares subroutine desribed previously for the models listed' in.table 3, and the Bayesian model disrimination method was again used to ompare these models (Table 6). Clear separation between the first-order and half-order models was not obtained with this limited amount of data and, to be onsistent with the seletion of the half-order model for the ase of the sand media, the half-order model was seleted to desribe the data

42..-....,... I. E...JC\1. I ~.-JC\1. 2.0r------r----.,...---~---r------..--- / T =T+273 k = 0 _ 82 exp-1960(1/t'-1/288)/r E- 0>.,...-.--.::---- 5 10. --.... / /~ ~... / 95%Limits 0.0'------...l...o::----'----L----1-----L--~ 0 15 20 25 30 T( C) FIGURE 6. ARRHENIUS TEMPERATURE DEPENDENCY FOR SAND MEDIUM ("'1.7 mo/m 2 -min)