- Email: [email protected]
Phosphorus removal from a real anaerobic supernatant by struvite crystallization
PII: S0043-1354(00)00498-X
Wat. Res. Vol. 35, No. 9, pp. 2167–2178, 2001 # 2001 Elsevier Science Ltd. All rights reserved Printed in Great Britain 0043-1354/01/$ - see front matter
PHOSPHORUS REMOVAL FROM A REAL ANAEROBIC SUPERNATANT BY STRUVITE CRYSTALLIZATION P. BATTISTONI1*, A. DE ANGELIS1, P. PAVAN2, M. PRISCIANDARO3 and F. CECCHI4 1
Engineering Faculty, Institute of Hydraulics, University of Ancona, Via Brecce Bianche, 60131 Ancona, Italy; 2 Department of Environmental Science, Science Faculty, Calle Larga S. Marta 2137, 30123 Venice, Italy; 3 Department of Chemistry, Engineering Faculty, Chemical Engineering and Materials, Monteluco di Roio, 67040 L’Aquila, Italy and 4 Scientific and Technological Department, Science Faculty, Strada Le Grazie, 37134 Verona, Italy (First received 2 December 1999; accepted in revised form 19 September 2000) Abstract}In this paper the phosphorus removal from a real anaerobic supernatant through the crystallization of struvite and or hydroxyapatite was investigated. A comparison between experimental results on phosphorus crystallization carried out in a fluidized-bed reactor (FBR) on a bench-scale and on a half-scale plant is presented, together with a double saturational model able to describe all experimental results, independent of the different geometry of the reactors, the distinct contact times and the unlike products obtained. Experimental results show that removal efficiencies are very satisfactory, and the maximum phosphorus removal is of 80%. # 2001 Elsevier Science Ltd. All rights reserved Key words}phosphorus, struvite, crystallization, fluidized-bed reactor, waste water treatment
NOMENCLATURE
B C Em F H HRT IP J k KSP n N Q t tc t1=2 v V X Xf
collision efficiency in (equation (21)) concentration, mol/m3 maximum crystallization yield, dimensionless ratio defined as Qair =Qi , dimensionless hydraulic head, m hydraulic retention time, s ionic product of supersaturated solution collision frequency rate constant for crystallization constant of solubility product number of cycles in the reactor, dimensionless particle concentration in equation (21) flowrate, m3/h time, s contact time, s half-time, s velocity, m/s volume, m3 phosphate conversion (equation (3)), dimensionless precipitation efficiency (equation (2)), dimensionless
Greek symbols a constant in equation (24) b supersaturation, dimensionless e porosity, dimensionless Z crystallization efficiency (equation (1)), dimensionless Subscripts 0 initial value ag aggregation *Author to whom all correspondence should be addressed. Tel.: +39-071-220-4530; fax: +39-071-2204236; e-mail: [email protected]
air c exp f i o s ric stripp t
relative to the crystallization relative to the final value influent outlet soluble relative to the relative to the total
air expanded bed
recycle flowrate stripping section
Abbreviations HAP Ca5(PO4)3OH–hydroxyapatite FBR Fluidized-bed reactor MAP NH4MgPO4}struvite
INTRODUCTION
The crystallization of hydroxyapatite (CA5(PO4)3OH }HAP) and/or struvite (NH4MgPO4}MAP) represents, among the different methods to remove phosphorus from secondary effluent wastewater, a technique free of sludge-handling problems with respect to the conventional process of phosphate precipitation. The crystallization of these calcium and magnesium phosphate salts can be achieved with various plant configurations (Kaneko and Nakajima, 1988; Joko, 1984) but fluidized-bed reactors (FBR) appear to be of particular interest (Van Dijk and Braakensiek, 1985; Egger et al., 1991) and their utilization is widespread in the field of nutrient removal (Fujimoto et al., 1991; Momberg and Oellerman, 1992). Fluidized-bed pilot plants have
2167
2168
P. Battistoni et al.
so far been conducted using a base dosage to obtain the pH necessary to reach the condition of supersaturation, but considering the high buffer capacity of the supernatant, the process becomes in this way very expensive; furthermore, it is necessary to add calcium or magnesium to satisfy the stoichiometric demand or to select MAP or HAP formation (Fujimoto et al., 1991). This research group has devised a phosphate removal technique by exclusively using the chemical physical properties of anaerobic supernatants without any addition of chemicals, reaching the operative pH only by aeration (Cecchi et al., 1994; Battistoni et al., 1995, 1996, 1997). In particular, as shown by Battistoni et al. (1997), a quantitative removal of phosphate is obtained by matching crystallization of MAP on sand quartz and CO2 stripping with air. Previously, Cecchi et al. (1994) had in fact demonstrated that MAP crystallization allows to operate in the metastable zone and requires a lower pH with respect to HAP precipitation, thus obtaining phosphate removal without the addition of alkali reagents. In a recent work Battistoni et al. (1998a) investigated the performance of an FBR plant fed by a real supernatant of an A2O process supplied from a centrifugation dewatering station, showing that the supernatant does not require any pretreatment to remove the low suspended solid content. Furthermore, the efficiency of the process was related to the pH and sand contact time in a double saturational model. The crystallization of phosphorus was also studied by operating with an FBR managed in batch mode (Battistoni et al., 2000); the experimental work allowed to individuate the feasibility and best operative conditions for the process without the addiction of reagents, showing that a good phosphate removal could be obtained through crystallization on quartz sand in about 100 min. The batch process assured excellent conversions, but if in the low phosphorus concentration range (30–60 mg PO4/l) satisfactory results were obtained as far as crystallization removal, at high phosphorus concentration range (100–150 mg PO4/l) the process needs further investigation to limit the precipitation. It was concluded that the management of FBR in a batch mode is not a reliable solution. All these investigations have so far been carried out in an FBR on a bench scale, by using an anaerobic supernatant coming from the dewatering section of a 100.000 population-equivalent (PE) civil wastewater treatment. The results obtained showed a percentage removal of more than 80% with MAP and HAP formation. The present paper extends the same process to an FBR on half-scale, located in the civil wastewater treatment plant of Ancona (Italy, 100.000 P.E., A2O Process) and continuously fed by the anaerobic supernatant coming from a centrifugation sludge station, without any treatment.
MATERIALS AND METHODS
Materials Real anaerobic supernatants from a centrifugation sludge section of a 100.000 PE civil biological nutrient removal plant (BNR) were used. The treatment plant uses an A2O configuration (anaerobic, anoxic, oxic process) to perform the carbon oxidation and nitrogen nitrification–dentirification with biological phosphorus removal. Mixed primary and secondary sludges are gravitationally thickened and anaerobically digested before being fed to the dewatering station. Analysis The chemical-physical characteristics were determined using the Standard Methods (APHA 1985). In particular phosphates were analyzed on influent and effluent average acidified samples as total phosphate (PO4t) and on 0.45 mm filtrate as soluble phosphate (PO4s). NH4, Mg and Ca ions concentrations were analyzed on acidified samples, while alkalinity was analyzed on natural sample. On the basis of average values the following quantities can be calculated. *
Crystallization efficiency (Z 100) Z% ¼
*
ðPO4tot out PO4sol out Þ 100 PO4tot in
ðPO4tot in PO4sol out Þ 100 ð3Þ PO4tot in
Crystallization percentage of PO4–MAP on total removed PO4 PO4 MAP% ¼
*
ð2Þ
Phosphate conversion (X%) X% ¼ Z% þ Xf % ¼
*
ð1Þ
Precipitation efficiency (Xf %) Xf % ¼
*
ðPO4tot in PO4tot out Þ 100 PO4tot in
ðMgin Mgout Þmol 100 ðPO4tot in PO4tot out Þmol
ð4Þ
Crystallization percentage of PO4–HAP on total removed PO4 PO4 HAP% ¼
ðPO4tot in PO4tot out Þmol ðMgin Mgout Þmol 100 ðPO4tot in PO4tot out Þmol
ð5Þ
Moreover, the trends with time of the pH, total and soluble phosphates, calcium, magnesium and alkalinity (carbonates and bicarbonates) entering and leaving the plant, were estimated. Pilot-plant The FBR plant, as sketched in Fig. 1, comprises two sections, a stripping tank and a fluidized-bed reactor. The reactor is a perpex column ðfin ¼ 0:09 m, height 2 m, vol. 12.8 l) filled with 9.5 kg of virgin silica sand (0.21–0.35 mm, fm ¼ 0:265 mm) to obtain a compress bed of 1 m height. At the bottom of the column a perpex cylinder ðfin ¼ 0:09 m, height 0.5 m), filled with gravel with decreasing size distribution, behaving as a filter, blocks the sand from coming back towards the pumps, and allows an homogenous distribution of the stream to the reactor. At the top of the column an expansion tank is provided in order to prevent the sand from coming out of the reactor.
Phosphorus removal from real anaerobic supernatant
2169
Fig. 1. Layout of pilot plant.
A stripper and a connected deaeration column compose the stripping section. The anaerobic supernatant, as supplied from the real plant, is previously stocked in a tank (3.0 m3 volume) to let a daily continuos feeding, then it was fed to the stripper with the recycle flowrate and an air flowrate necessary for CO2 stripping. The effluent exists from the deaeration column together with the recycle flowrate. Three levels from which the effluent can outlet are provided, thus allowing to work with different hydraulic heads (H ¼ 0:5, 0.8, 1.0 m) and different volumes of stripping section and deaeration column. Temperature and pH probes were positioned in the deaeration column, while a second temperature probe was placed in air pipeline. The recycle pump has a flowrate ranging from 50 to 750 l/h. In Table 1 the main characteristics of FBR system are reported. The pilot plant was located in the dewatering station of a wastewater treatment real plant, the three sections are built on three different platforms to avoid useless surface request (see Fig. 2), it was conducted in a continuous mode. Five sets of experimental runs were performed; each set is characterized by a constant feeding flowrate of anaerobic supernatant while the air flowrate can be changed to operate at different pH. The ratio between recycle and influent flowrate can range from 0.6 to 1.4, the value adopted is that able to operate with a constant volume of expanded bed in all experiments (up to the expansion tank). Each run was monitored either by sampling incoming and effluent from feeding tanks and plant effluent, or by logging temperature and pH on a personal computer.
sand was then examined by means of differential thermal analysis, chemical analysis and grain-size distribution analysis. Microstructural analysis Morphological and microstructural analyses of the deposits on sand grains were performed by scanning electron microscopy (SEM) techniques (Phillips XL20 instrument). Energy-dispersive X-ray (EDX) spectrometry was used to analyze the elements and deposits (Edax PV 9800 instrument). Differential thermal analysis The differential thermal analysis was performed on exhaust sand grains by slowly increasing the temperature in order to find the presence of carbonates that, at T¼ 8008C, decompose by releasing CO2. The loss in weight of the sample allows to establish the percentage of carbonates initially present. Hydrodynamic of the process In the FBR system four hydraulic retention times can be identified, referring to Fig. 1: HRTt ¼
ðV1 þ V2 þ V3 Þ Qi
ð6Þ
ðV1 þ V2 Þ Qi
ð7Þ
Sand analysis Sand removed from FBR plant is air dried for 72–96 h, then weighted and analyzed to measure the porosity e. The
HRTstripp ¼
2170
P. Battistoni et al. Table 1. Volumetric characteristics of FBR
Apparatus
Characteristics and material
Volume, geometry and flowrate
Feeding tanks FBR reactor
PVC tanks Inner diameter 0.09 m Height 2 m Perpex column PVC Inner diameter 0.08 m Height 1.6 m PVC column Inner diameter 0.31 m Height 2.05 m PVC column PVC Power 0.25 kW Recycle pump Type: Progressive cavity Power max. 0.74 kW Feeding pumps Type: Peristaltic, Power 0.06 kW Progressive cavity power 2.0 kW Tair, T stripper, pH stripper
Variable volume up to 3.01 m3 Compressed-bed height: 1.5 m free volume 3.2 l 1.0 m 6.4 l 0.5 m 9.5 l Volume 18 l Hydraulic head 0.5, volume 2.7 l Hydraulic head 0.8, volume 4.2 l Hydraulic head 1.0, volume 5.2 l Hydraulic head 0.5, volume 40 l Hydraulic head 0.8, volume 62.6 l Hydraulic head 1.0, volume 77.7 l Volume 63 l
Expansion tank Deaeration column
Stripper E sampling tank B-blower Pl-
P2 Signal to personal computer
Flowrate 0.5–2.5 m3/h Peristaltic pump flowrate 0–100 l/h Peristaltic pump flowrate 0–150 l/h Progressive cavity pump flowrate 0–1500 l/h
RESULTS AND DISCUSSION
Fig. 2. Pilot organization.
V3 Qric
ð8Þ
Vexp e Qric
ð9Þ
HRTFBR ¼
HRTexp ¼
where Qi is the feed flowrate, Qric the recycle flowrate, V1 the stripping section volume, V2 the volume of the quiet zone, V3 the column free volume (including tube, expansion tank and filter volumes), Vexp the volume of the fluidized bed and e the bed porosity. In the upward movement of the liquid inside the FBR, the number of passages (n) is defined as the ratio between the mean hydraulic retention time in the plant and the time spent in the expanded bed, while the contact time (tc ) on sand grains depends on the number of cycles and the time needed for one single passage: HRTT HRTFBR
ð10Þ
tc ¼ nHRTexp
ð11Þ
n¼
Anaerobic supernatants are supplied from a centrifugation sludge station, without any previous treatment, and their averaged data are reported in Table 2. These data, compared with those obtained in the previous experimentation on bench-scale pilot plant (Battistoni et al., 1998a; see Table 3), allow to observe the lower content of total phosphates in the inlet stream (68 mg/l against 139 mg/l of the previous experimentation), the increased content of calcium, magnesium, ammonia and alkalinity and the ratio Ca/Mg almost one-half. It is worth noting that both the anaerobic supernatant used in this work and the anaerobic supernatant previously used (Battistoni et al., 1998a) are real liquors coming from an A2O process. The different composition is related to a change in technical management of the plant that happened a year later. Anyway, on the basis of the average PO4 content in supernatants, it is possible to note that the stoichiometric demand of calcium for HAP formation, or of calcium and magnesium for MAP formation, is completely satisfied. In the previous experimentation, due to the higher PO4 content, the stoichiometric demand was satisfied for HAP, whereas it was satisfied only to a limited extent for MAP (Mg/PO4=0.7 mol/mol, see Table 3). Five sets of experiments were carried out (A–Q Table 4a) characterized by different feed flowrates (17.6, 50–53, 67, 110 and 190 l/h), up to 10 times higher with respect to experimentation on laboratory scale. At the beginning of each experiment the feed flowrate, recycle flowrate and air flowrate were set to a fixed value. For each set of experiments, the air flowrate was modified in order to investigate the variations in the crystallization and precipitation efficiency while changing pH and the parameter F, given by the ratio between air and influent flowrates ðF ¼ Qair =Qi Þ. Each experiment lasted from 16 to 44 h, even though short durations were preferred,
Phosphorus removal from real anaerobic supernatant
2171
Table 2. Characteristics of liquors Parameters
Unit
Average
Min.
Max.
S.D.
n
pH PO4 NH4 Ca Mg HCO3 CO3 Ca/Mg Ca/PO4 Mg/PO4 NH4/PO4
mg/1 mg/1 mg/1 mg/1 mgCaCO3/1 mgCaCO3/1 mol/mol mol/mol mol/mol mol/mol
7.62 68 1055 200 54 7438 0 2.2 7.1 3.3 85.9
7.41 43 810 108 40 6400 0 1.6 4.2 1.9 60.41
7.87 89 1170 324 73 8050 0 3.0 13.3 5.4 133.7
0.11 14 92 50 11 393 0 0.5 2.4 1.1 21.6
22 22 22 22 22 22 22 22 22 22 22
Table 3. Comparison among liquors from laboratory and full-scale plants Plant scale
PO4t ot (mg/l)
NH4 (mg/l)
HCO3 (mgCaCO3/l)
Ca (mg/l)
Mg (mg/l)
Ca/Mg (mol/mol)
Ca/PO4 (mol/mol)
Mg/PO4 (mol/mol)
Laboratory Half
139 68
914 1055
3550 7400
153 200
24 54
3.7 2.2
2.6 7.2
0.7 3.3
Table 4. (a) Experimental results Set
Influent
Characteristics
Operative
Conditions
Ca Mg NH4 Qi time QAIR No. Run PO4 tot PO4 sol HCO3 (mg/l) (mg/l) (meq/l) (mg/l) (mg/l) (mg/l) (1/h) (h) (m3/h) 1 1 1 1 2 2 3 3 2 3 3 4 4 3 5
A B C D E F G H I L M N O P Q
58.2 41.3 43.7 45.3 40.4 42.6 56.7 74.1 82.1 83.0 75.4 68.5 73.5 73.5 57.5
57.4 31.0 39.4 41.3 34.5 38.3 53.2 30.0 54.2 53.4 58.1 52.5 59.0 48.4 32.2
69 72 76 77 73 73 77 68 70 70 72 76 76 75 69
132 113 139 156 181 183 194 214 216 205 206 205 203 199 195
41 42 63 60 61 61 57 61 45 43 44 47 47 48 42
1090 1130 1054 1135 1100 980 1160 990 972 1062 990 1030 998 1080 954
18 18 18 18 50 50 67 67 53 67 67 110 110 67 190
43.5 25.5 34.0 27.0 41.2 27.7 21.7 25.0 18.0 22.0 18.0 20.0 16.0 34.8 19.0
18.4 6.0 2.6 5.7 1.8 4.5 12.3 4.9 2.9 1.5 3.4 7.8 1.9 1.3 8.2
Effluent
F
tc (h)
pH
1047 339 145 324 36 94 183 74 54 22 51 71 18 19 43
1.43 1.41 1.32 1.31 0.43 0.39 0.26 0.23 0.47 0.37 0.35 0.20 0.19 0.30 0.09
8.78 8.55 8.55 8.47 8.07 8.45 8.68 8.45 8.31 8.06 8.30 8.44 8.07 8.04 8.22
Performances
PO4tot out PO4s out Z (mg/1) (mg/1) (%) 11.6 11.6 12.1 10.1 17.0 13.3 12.7 25.9 26.4 35.5 27.3 23.8 34.5 30.7 26.5
11.1 7.9 11.6 9.8 16.7 12.9 12.2 23.0 25.6 34.4 27.0 21.9 32.0 28.5 23.5
80 74 74 78 55 72 78 65 68 57 63 65 53 58 55
X (%)
Xf (%)
84 81 73 81 72 91 82 76 71 60 64 73 60 65 66
4 7 1 3 2 3 4 11 3 3 1 8 7 17 1
(b) Experimental results}HAP and MAP formation No. A B C D E F G H I L M N O P Q a b
DPO4 (mg/l)
DCa (mg/l)
DMg (mg/l)
Dalkal. (mg/l)
PO4HAPa (%)
PO4MAPa (%)
HAPb (%)
MAPb (%)
CaCO3b (%)
46.6 29.7 31.6 35.2 23.7 29.3 44 48.2 55.7 47.5 48.1 44.7 39 42.8 31
110.2 94.6 45.9 94.2 95.5 117.4 143.2 112.5 87.9 51.5 96.7 112.7 31.1 55 88.9
4.7 2.33 1.2 7.9 6.7 9.2 11.5 13.3 9.6 7 10.9 11.8 7.7 8.3 7.2
7.41 16.5 14.5 15.5 7.3 12.9 13.9 7.5 5 3.3 6.1 7.4 4.2 6.3 5.6
59 68 85 8 0 0 0 0 30 40 8 0 20 24 5
41 32 15 92 100 100 100 100 70 60 92 100 80 79 95
4 3 11 0 0 0 0 0 2 5 0 0 3 2 0
8 5 6 13 11 12 12 17 17 23 17 15 33 22 12
88 92 83 87 89 88 88 83 80 72 83 85 64 76 87
Percentage relative to phosphorus distribution. Molar ratios on total.
since dewatering station was open only 8 h a day. The five main sets of experiments, classified on the basis of feed flowrates, allow to study 16 operation conditions, different for pH and contact time. Among
various experiments, variability in supernatants total phosphorus content is observed, probably due to the different way of operating of the post-thickener in real plant. Moreover, during each run, a decrease in
2172
P. Battistoni et al.
phosphorus soluble fraction ðPO4s =PO4tot Þ is found. This circumstance is mainly dependent on supernatant residence times in the feed tank (see Fig. 1). In fact, the aging of anaerobic supernatants results in a spontaneous precipitation of HAP and/or MAP, and therefore in a decrease of PO4s/PO4 tot as the supernatant comes to equilibrium with the air. This fact was observed both for anaerobic supernatants coming from nitrogen biological removal wastewater treatment plants added with orthophosphates; Battistoni et al., 1997) and for supernatants coming from nutrients biological removal wastewater treatment plants (Battistoni et al., 1998b); in this last case, supernatants are similar to those used in the present study. An analysis of runs shows that in all of them, crystallization efficiencies vary from 53 to 80%, the precipitation efficiency is normally very low: in 9 of the 15 runs precipitation efficiency is less than the cut-off value chosen to integrate the plant with a post-filtration operation, while in three of the 15 runs precipitation efficiency ranges from 7 to 8% and in three of the 15 runs it is higher than 10%. Each run has a typical behavior, made by a first phase (lasting up to 3 h, depending on F value and on hydraulic level in the deaeration column), in which pH exponentially grows from supernatant to experimental value, and a stationary phase, in which the average pH at 300 remains on constant. Variations in influent characteristics depend on stock or on feed renewal when the run is working. In Fig. 3(a), a typical behavior of operating parameters for D run is reported; in particular, pH at the stationary phase, after 2.5 h is equal to 8.46 (S.D. 0.04), and the initial concentration of PO4 sol is constant, while slight variations in PO4 tot are present at the beginning, at half and at the end of the run, all below 5 mg/l. At the end of the run a variation is present in PO4 tot out and PO4 sol out, and moreover pH is varied by 0.05% with respect to the previous value. The change in pH value reported in Fig. 3(b) is ascribed to the higher value of alkalinity, and therefore to the buffer effect. In conclusion, the alkalinity loss between influent and effluent is observable from the periodic check of their chemical-physical characteristics (Fig. 3(b)). Salt precipitation Several salts may form from a P-enriched or not real anaerobic supernatant; when a certain amount of magnesium is present, only MAP can be obtained in the presence of ammonium and phosphate ions (Seckler et al., 1991), since the formation of Mg(OH)2 was found only operating in a bed of MgO (Angel, 1999), or at high values of pH (ShulzeRettmer, 1991); in particular, when pH10.5 the retarding effect of Mg ions on phosphate formation tends to disappear (Jenkins et al., 1971). Moreover, the coprecipitation of Mg3(PO4)2 with amorphous Ca3(PO4)2 was observed in clear waters added with Ca, PO4 and Mg ions (Seckler et al., 1991, 1996b). On
Fig. 3. Typical behavior of (a) phosphates and pH with time (run D), and (b) alkalinity (run D).
the other hand, different salts can precipitate when calcium is present, namely dicalcium phosphate (CaHPO4), tricalcium phosphate (Ca3(PO4)2), octocalcium phosphate (Ca4H(PO4)3) and apatites (Zoltek, 1974). The formation of these phosphate salts seems to conform to the Ostwald–Gay Lussac Step Rule, which states that the less thermodynamically stable, and hence the more soluble, form of a polymorphous solid is generally the most readily precipitable; therefore the likely order of formation would be dicalcium phosphate, tricalcium phosphate, octocalcium phosphate and finally, the apatites (Zoltek, 1974). Amorphous calcium phosphates are always present in clear water in supersaturated conditions (Seckler et al., 1996a, b), or sometimes in the main effluent stream of a civil wastewater treatment plant (Van Dijk and Braakensiek, 1985; Angel, 1999), while crystalline HAP was found working in the metastable zone in effluents from lagoons, secondary sedimentation tanks or anaerobic supernatants from sludge treatment devices (Momberg et al., 1992; Joko, 1984). Finally, HAP
Phosphorus removal from real anaerobic supernatant
crystallization was observed in synthetic liquor also, when a low phosphorous concentration was present (Kaneko et al., 1998; Joko, 1984). As regards the socalled calcite (CaCO3), carbonate ions have a tendency to crystallize as calcite when the stoichiometric ratio Mg/Ca is less than 0.6 (as in the current work, see Table 3); otherwise as dolomite CaMg(CO3)2 when Mg/Ca is higher than 0.6 (Diaz et al., 1994). In conclusion, in the present work experimental results will be analyzed considering the formation of merely HAP, MAP and calcite. Other kinds of salts will not be taken into consideration, giving to the analysis of the exhaust grains the task of confirming what was supposed. Moreover, in a previous work carried out on the same supernatant in a bench-scale FBR (Battistoni et al., 1998a), the analysis of sand after the crystallization process demonstrated the presence of HAP and MAP only. As for the molar losses, unlike results previously obtained (Battistoni et al., 1998a), together with MAP formation the presence of a great deal of calcite (CaCO3) was found (see Table 4b). The analysis of the table reveals the MAP formation is predominant in almost all experiments. This can be related to the Ca/Mg ratio (2.2}Table 3) lower than that previously found (3.7}Table 3) and a Mg/PO4 ratio equal to 3.3. However, these ratios do not explain the calcite crystallization because the molar ratio Ca/PO4 is enough to guarantee a complete HAP formation. A possible explanation of the phenomena can be found in the high alkalinity and Mg content and their inhibitory effect on HAP formation (Battistoni et al., 1998b; Joko, 1984; Ferguson et al., 1970, 1973; Jenkins et al., 1971). This circumstance could be explained by assuming that Ca precipitates as salt different from HAP, determining the prevailing MAP formation despite its massive presence in the supernatant, this hypothesis being confirmed by the analysis of the exhaust grain sand reported in subsequent section. Operative pH The operative pH can be approached by fixing an air flowrate, and therefore a value for F, which is related to pH by means of a saturational relationship. For experiments carried out on laboratory scale the data can be modeled (multiple regression analysis}Anova method) according to the following equation: pH ¼ 1:125
F þ 7:70 171 þ F
2173
between pH and F given by the following expression: pH ¼ 1:100
F þ 7:63 33 þ F
with R2 ¼ 0:89 and S.E.=0.08. Once the value of air flowrate, and therefore of F, has been fixed, the value of pH in condition of saturation depends on the hydraulic head, and on whether FBR is active or not. In particular, in Fig. 4 first pH increase, when only the stripping section is active, and a second increase, when the whole plant is working, are shown. Three tests are performed at a constant feeding flowrate of 621/h and an F ¼ 40 l air/l feeding, while the hydraulic head changes from 0.5 to 0.8 and 1.0 m. The main conclusions that can be drawn are: *
*
a higher hydraulic head allows to obtain a higher pH, even if a head of 0.8 m is enough to reach the operative conditions; when FBR also works, a pH higher than that obtained when operating with the stripping section only, can be obtained.
This behavior can be interpreted considering that a higher hydraulic head results in a major use of the diffused air flowrate, a lower buffer capacity of waste when calcite crystallizes in the FBR and an increased CO2 stripping by recirulation pump; however, up to 3 h are required to obtain the operative pH (steadystate conditions). Consequently, in order to make a comparison among different reactors, the only parameter to be taken into account is the pH. Crystallization efficiency Efficiency is strongly related to the operative pH, both for experiments carried out on laboratory scale (Battistoni et al., 1998a) and for experiments on a half scale; in particular, in this last case the highest efficiency, equal to 80%, occurred at the highest pH, equal to 8.78 (Table 4a, run A). This strong dependence is shown in Fig. 5, and on this basis an empirical correlation, linear or saturational, between
ð12Þ
with R2 ¼ 0:41 and S.E.=0.13. The low value obtained for R2 is due to the different kinds of feed used, namely liquor without any additives, liquor with calcium ions added up, liquor with magnesium ions added. On the contrary, the experimental results relative to the half-scale plant individuate a relationship
ð13Þ
Fig. 4. Dependence of pH on hydraulic head.
2174
P. Battistoni et al. Table 5. Statistical parameters obtained with multiple regression analysis ANOVA Equation
R2
S.E.
(14) (15) (16) (17)
0.81 0.85 0.79 0.81
0.07 0.10 3.03 4.05
Fig. 5. Crystallization efficiency as a function of pHexperimental results.
pH and Z was found for each plant configuration, (linear and multiple regression, Anova method) expressed by the following equations: pH ¼ 6:62 þ 0:022Z%
bench scale
pH ¼ 6:81 þ 0:023 Z% half scale
ð14Þ ð15Þ
Z% ¼ 100
ðpH 7:65Þ bench scale ð16Þ ðpH 7:65Þ þ 0:201
Z% ¼ 100
ðpH 7:65Þ ðpH 7:65Þ þ 0:340
half scale ð17Þ
with the statistical parameters obtained with multiple regression analysis ANOVA reported in Table 5. Unfortunately, this first linear model does not explain much more than 85% of the result variance. Furthermore, this saturational model at a deeper insight appeared to be not sufficiently adequate to describe the dependence of crystallization efficiency on process parameters, since a strong dependence of Z not only on pH, but also on another variable, was found. In Fig. 6 the behavior of efficiency with contact time (tc ) is reported for both plant configurations, showing that the contact time on sand grains represents a major parameter in the characterization of process performances. A double-saturational model, taking into account both pH and tc , was devised, with a multiple regression (Anova method) which exhibits an excellent agreement with experimental results. For the experiments carried out in the half plant, the model is expressed by the following equation: Z% ¼ 100
ðpH 7:65Þ tc ðpH 7:65Þ þ 0:290 tc þ 0:017
ð18Þ
with R2 ¼ 99:84 and S.E.=2.76. The double-saturational model was originally presented for laboratory-scale experiments (Battistoni et al., 1998a), and even in that case it allowed a
Fig. 6. Crystallization efficiency as a function of contact time-experimental results.
successful interpretation of experimental data. The versatility of the model, able to describe experimental results carried out in very unlike reactors, relies on the fact that it is well sustained by hydraulic and engineering parameters, as defined by equations (6)–(9). Therefore experimental results, obtained in different geometry reactors, with distinct contact times and giving unlike products can be accurately described by using the same mathematical model, as demonstrated by the following equation: Z% ¼ 100
ðpH 7:63Þ tc ðpH 7:63Þ þ 0:208 tc þ 0:039
ð19Þ
with R2 ¼ 99:73 and S.E.=3.78. Equation (17) has been in fact obtained by correlating experimental data relative to both bench-scale and half-scale pilot plant configurations. The half-efficiency constants for pH and tc introduced in equation (18) give useful information: an operative pH of 7.94 (calculated as 0.29+7.65, see equation (18)) reduces the efficiency by half, and the same result is obtained when operating at a contact time equal to 0.017 h (about 1 min). The contact time is defined by equation (11) as a function of the number of cycles in the reactor, but for the same tc different n are feasible since e changes with phosphate growth on sand and consequently HRTexp changes even if Qric and Vexp remain constant. In fact, e varies from a maximum of 0.655 to a minimum of 0.405 as shown in Table 6, passing from virgin to exhaust
Phosphorus removal from real anaerobic supernatant
2175
Table 6. Porosity and contact times in FBR experiments No. A B C D E F G H I L M N O P Q a
e
tc (h)
0.655 0.641 0.594 0.583 0.546 0.489 0.431 0.365a 0.653 0.634 0.611 0.568 0.530 0.505 0.405
1.4326 1.4076 1.3198 1.3070 0.4324 0.3937 0.2633 0.2275 0.4745 0.3662 0.3548 0.2033 0.1917 0.3017 0.0878
Exhaust sand substituted with virgin sand.
Fig. 7. Crystallization efficiency as a function of contact time, double saturational model (equation (18)).
sand. Definitely, the contact time can be expressed with the following expression: HRTt tc ¼ n HRTexp: ¼ HRTexp HRTFBR ðV1 þ V2 þ V3 Þ Qric eVexp aeVtot ¼ ¼ V3 Qric Qi Qi
ð20Þ
where a is a dimensionless constant, used to express the column free volume (V3 ) as a function of the expanded bed volume (Vexp ). Following the definition given by equation (20), the contact time appears to be the best operative parameter to be considered for comparing results from different pilot plants. The introduction of the double-saturational model (equation (18)) in a pH range which can be easily gained (8.1–8.7) allows to fix the most suitable contact time to reach in order to get a prefixed efficiency. As shown in Fig. 7, it can be seen that for a pH equal to 8.5, in order to get an efficiency Z of 70%, a tc equal to 0.3 h is needed, while efficiencies of 80% are very difficult to get with a pH less than 8.7. Several models are available in the literature to describe the phosphorus precipitation process; they can be distinguished in models based on primary nucleation mechanism (e.g. nucleation is caused by pure supersaturation), and models based on secondary nucleation mechanisms (e.g. nucleation and growth take place on pre-existent seeds, in metastable supersaturation conditions). The work of Seckler (Seckler et al., 1996a–c) belongs to the first group, proposing a theoretical model for fine-particle aggregation with sand grains. This model takes into account both aggregations of fines with sand and particle breakage when a high dissipation of the kinetic energy takes place. The following equation derives from a particle number balance, and expresses the decrease in the fine concentration by aggregation with grains in a fluidized bed: dNi ¼ BJii b dt
ð21Þ
where N is the initial particle concentration, Jii is the collision frequency, b is the supersaturation, t is the reaction time and B is the collision efficiency, which is influenced by the energy dissipation rate. By integrating equation (21) from t ¼ 0 to t ¼ tout , the phosphate removal efficiency Zag (by aggregation only) can be calculated, by assuming the particle size of the fines to be constant in time: Zag ¼
Ni;in Ni;out Ni;in
ð22Þ
It was found, both theoretically and experimentally, that the aggregation could be increased by spreading the supersaturation more evenly throughout the reactor, while choosing fluidization conditions where the energy dissipation rate in the bed is minimized can reduce the breakage. This process has been studied in the precipitation of phosphorus from tap water or from a phosphate-enriched stream of a sewage treatment plant. Anywhere, the model describes experimental results only to a limited extent, since the aggregation of fines can be strongly influenced by molecular growth. Regarding the models based on secondary nucleation mechanism, De Rooij et al. (1984) studied the formation of different phases of calcium phosphates on a seed material, in well-defined experimental conditions, such as a fixed value of temperature (T ¼ 378C), ionic strength (I=0.10 mol/l), Ca/P ratio (Ca/P=1.333) and for different pH values (55pH58). In order to avoid any interference by traces of impurities in reagent-grade chemicals, ultrapure reagents were used. The model obtained is that shown by the following equation: p dn 1=v ¼ ks IP1=v KSP ð23Þ dt where n is the number of precipitated moles, k the rate constant for crystallization, s a factor proportional to the number of sites available for growth, p
2176
P. Battistoni et al.
the effective order of reaction, v the number of ions in the formula unit, IP the ionic product of supersaturated solution and KSP the thermodynamic solubility product. The Gibbs free energy related to the transfer from supersaturated solution to an assumed saturated solution at the surface of the developing solid phase, gives the relationship between IP and KSO (De Rooij et al., 1984). Kaneko and Nakajima (1988), in order to describe the crystal growth of HAP, starting from a metastability condition of synthetic water solutions in which the dependence of pH is not present, made an approximation of equation (23) expressed as dC ¼ ksC2 ð24Þ dt where C is the molar concentration and t is the retention time. Since it was demonstrated that for the crystallization of MAP and HAP from anaerobic supernatant in fixed-bed reactors FBR works in conditions of metastability (Battistoni et al., 1998b), the approximation made by Kaneko and Nakajima (1988) is suitable to be used. This kinetic equation can be integrated from 0 to t, leading to the empirical saturational model in tc (equation (18)) of crystallization efficiency Z (Battistoni et al., 2000a): Z ¼ Em
tc t1=2 þ tc
ð25Þ
Moreover, the maximum crystallization efficiency (Em ) and the half-time (t1=2 ) can be introduced: C0 Ct C0
ð26Þ
1 ðC0 Ct Þk
ð27Þ
Em ¼ t1=2 ¼
where Ct is the final concentration, C0 the initial concentration and tc the contact time. The empirical double saturational model thus follows from the theoretical model of crystallization in conditions of limiting supersaturation for what concerns the saturation in tc , while the behavior as a function of pH has always been reported in the literature but never described in a theoretical model. Experimental runs (see Table 4) show that an almost wide range of precipitation efficiency was individuated, for which no explanation is given. In conclusion, with respect to a previous work (Battistoni et al., 1998a), in this paper it has been possible to reduce the contact time with a higher pH by increasing the air-to-influent ratio. A further experimental work is required to improve the validity of the model proposed, even at a lower contact time.
analysis of the grain; comparison among the total amount of phosphorus removed calculated through a mass balance carried out on experimental results and exhaust matrix; grain size distribution on virgin and exhaust silica sands; thermal differential analysis on grains and scanning electron microscopy analysis. Sand samples were taken from FBR in the middle (M) and at the end of the experimentation (E). The chemical analysis highlighted a first difference with respect to bench-scale experiments (Battistoni et al., 1998): in that case phosphorus was removed entirely as HAP (62%) and MAP (38%), while in this case the percentage of MAP ranges from 34% on weight referred to the grain growth (M) to 50% (E) and HAP varies from 0 (M) to 5% (E). Truly, on grain surface a great deal of calcite (CaCO3) has formed (66% (M)–45% (E)), subtracting Ca ions for HAP cyrstallization. MAP formation is thus explained despite the high value of the ratio Ca/Mg. A comparison between the amount of phosphorus present in matrix and that removed, computed as the total of all the phosphorus crystallized in the different runs, allowed to confirm the worth of process, since all phosphorus removed from the liquid phase was found in the reactor solid matrix. This circumstance strengthens the validity of equations introduced to define the crystallization and precipitation efficiencies (see equations (1) and (2)) since a significant precipitation in the stripping and deaeration section is not expected. In Table 7 the increment of grain average dimensions obtained through a grain size distribution is reported: the nominal diameter passes from 0.27 mm of virgin sand to 0.35 mm of exhaust sand, with a growth of 26%, on which the presence of calcite had a major role. The differential thermal analysis reveals the classic weight loss at 8008C temperature connected with the presence of carbonate in exhaust sand. In particular, the conversion of CO2 loss into CaCO3 gives an amount equivalent to that calculated from chemical analysis of exhaust sand where the Ca ions in excess to HAP formation are computed as CaCO3 formation. This evidence is repeated either in M or in E sand samples. The SEM analysis interpretation requires all the information obtained by previous techniques. The photograph of an exhaust grain (M condition} Fig. 7) shows the morphological aspect of exhaust sand with different zones classified as sheet (marker A) and sphere (marker B) types. X-ray dispersive
Table 7. Silica sand nominal diameter
Analysis of the exhaust sand grain
Sand type
d (mm)
Some features of the process were clarified through an analysis of the exhaust sand grain, which was carried out according to the following steps: chemical
Virgin sand Exhaust sand (M) Exhaust sand (E)
0.2672 0.3156 0.3478
Phosphorus removal from real anaerobic supernatant
has been presented. The conclusions arising from this comparison can be summarized as follows:
Table 8. Elements on deposit (EDX measurement) Zone
P (%)
Mg (%)
Ca (%)
35.0 61.4 13.6
18.5 32.2 7.2
46.5 6.4 79.1
*
Average Sheet Sphere
2177
*
*
*
*
Fig. 8. Morphological analysis of sand grain.
spectrometry was performed in both zones and on middle one with dimensions of 200 200 mm. The results (Table 8) do not highlight trace of Si. This means that EDX spectrometry is not able to investigate the whole deposits, but only a fraction and most probably the last. This agrees with the average thickness of deposit equivalent to 40.3 mm (see Table 8). Furthermore, the elements in different zones reveal that the sheet deposit is mainly formed by MAP while the sphere zones composition can be translated in a MAP (15%), HAP (2%) and CaCO3 (83%) molar distribution, this last being very similar to those calculated in P and Q experiments on FBR plant (see Fig. 8 and Table 4b). Coming to conclusions, in spite of all investigations carried out on grains, a clear evidence whether the deposit on sand is Ca3(PO4)2 or HAP was not achieved. Analytical mistakes in chemical analysis may be sufficient to cover the slight difference between the formations of one species with respect to another. This is the reason why X-ray spectroscopy should be desirable, in order to ascertain which kind of salt is formed. Notwithstanding the process modeling is still well founded, since it has been carried out on phosphorus removal, independent of the crystallized salt.
CONCLUSIONS
In this paper a comparison between experimental results on phosphorus crystallization carried out in an FBR reactor on a bench-scale and on half-scale
the double saturational model is able to describe all experimental results, independent of the different geometry of the reactors, the distinct contact times and the unlike products obtained; a further experimental work is required to further improve the validity of the model proposed, even at a lower contact time; with respect to bench-scale experiments in half-scale experiments phosphorus is removed almost entirely as MAP and a great deal of calcite forms on grains; a further investigation is needed to foresee the formation of MAP or HAP on silica sand grains; also in the case of plant on a half scale the efficiencies are very satisfactory, and the maximum phosphorus removal is of 80%.
Acknowledgements}The authors would like to thank the Ministry of the University and of the Scientific and Technology Research of Italy for financial support.
REFERENCES
Angel R. (1999) Removal of phosphate from sewage as amorphous calcium phosphate. Environ. Technol. 20, 709–720. APHA, AWWA and WPCF (1985) Standard Methods for the Examination of Water and Wastewater, 16th ed. American Public Health Association, Wastington, DC. Battistoni P., Fava G., Pavan P., Musacco A. and Cecchi F. (1995) Phosphate crystallization in a fluidized bed reactor. Preliminary results. Second International Symposium on Wastewater Reclamation and Reuse, Iraklio, Crete, Greece, October 17–20. Battistoni P., Fava G., Pavan P. and Mata-Halvarez J. (1996) Phosphate removal by crystallization. Anaerobic liquors treated in a fluidized bed reactor. Proceedings of Water Quality International 96, Singapore, June 23–28. Battistoni P., Fava G., Pavan P., Musacco A. and Cecchi F. (1997) Phosphate removal in anaerobic liquors by struvite crystallization without addition of chemicals. Preliminary results. Water Res. 31, 2925–2929. Battistoni P., Pavan P., Cecchi F. and Mata-Halvarez J. (1998a) Phosphate removal in real anaerobic supernatants: modeling and performance of a fluidized bed reactor. Water Sci. Technol. 31, 275–283. Battistoni P., Pavan P., Cecchi F. and Mata-Alvarez J. (1998b) Effect of composition of anaerobic supernatants from an anaerobic, anoxic and oxic (A2O) process on struvite and hydroxyapatite formation. Ann. Chim. 88, 761–772. Battistoni P., Pavan P., Prisciandaro M. and Cecchi F. (2000) Struvite crystallization: a feasible and reliable way to fix phosphorus in anaerobic supernatants. Water Res. 34, 3033–3041. Cecchi F., Battistoni P., Pavan P., Fava G. and MataHalvarez J. (1994) Anaerobic digestion of OFMSW and BNR processes: a possible integration. Preliminary results. In Proceedings of the 17th Biennal Conference, Water Sci. Technol. 30, 65–71.
2178
P. Battistoni et al.
De Roij J. F., Heughebaert J. C. and Nancollas G. H. (1984) A pH study of calcium phosphate seeded precipitation. J. Colloid Interface Sci. 100, 350–358. Diaz O. A., Reddy K. R. and Moore P. A. (1994) Solubility of inorganic phosphorus in stream water as influenced by pH and calcium concentration. Water Res. 28, 1755– 1763. Egger E., Dirkzwager A. H. and Van der Honing H. (1991) Full-scale experiences with phosphate crystallization in a crystalactor. Water Sci. Technol. 23, 819–824. Ferguson J. F., Jenkins D. and Stumm W. (1970) Calcium phosphate precipitation in wastewater. Chem. Prog. Symp. Ser. 67, 279–287. Ferguson J. F., Jenkins D. and Fasrman J. (1973) Calcium phosphate precipitation at slightly alkaline pH values. J. Water Pollut. Control Fed. 45, 620–631. Fujimoto N., Mizouchi T. and Togami Y. (1991) Phosphorus fixation in the sludge treatment system of a biological phosphate removal-accomplishments and needs. Water Res. 25, 1471–1478. Jenkins D., Ferguson J. F. and Menar A. B. (1971) Chemical process for phosphate removal. Review paper. Water Res. 5, 369–389. Joko I. (1984) Phosphorus removal from wastewater by the crystallization method. Water Sci. Technol. 17, 121–132. Kaneko S. and Nakajima K. (1988) Phosphorus removal by crystallization using a granular activated magnesia clinker. J. Water Pollut. Control Fed. 60, 1239–1244.
Momberg G. A. and Oellerman R. A. (1992) The removal of phosphate by hydroxyapatite and struvite crystallization in South Africa. Water Sci. Technol. 26, 967–976. Schulze-Rettmer R. (1991) The simultaneous chemical precipitation of ammonium and phosphate in the form of magnesium–ammonium–phosphate. Water Sci. Technol. 23, 659–667. Seckler M. M., Bruinsma O. S. L. and Van Rosmalen G. M. (1991) Crystallization of calcium and magnesium phosphates in a fluidized bed. Proceedings of the Third European Conference on Crystal Growth, Budapest, Hungary, pp. 932–937. Seckler M. M., Van Leeuwen M. L. J., Bruinsma O. S. L. and Van Rosmalen G. M. (1996a) Phosphate removal in a fluidized bed. II. Process optimization. Water Res. 30, 1589–1596. Seckler M. M., Bruinsma O. S. L. and Van Rosmalen G. M. (1996b) Calcium phosphate precipitation in a fluidized bed in relation to process conditions: a black box approach. Water Res. 30, 1677–1685. Seckler M. M., Bruinsma O. S. L. and Van Rosmalen G. M. (1996c) Phosphate removal in a fluidized bed}I. Identification of physical processes. Water Res. 30, 1585–1588. Van Dijk J. C. and Braakensiek H. (1985) Phosphate removal by crystallization in a fluidized bed. Water Sci. Technol. 17, 133–142. Zoltek J. (1974) Phosphorus removal by orthophosphate nucleation. J. Water Pollut Control Fed. 46, 2498–2520.
Wat. Res. Vol. 35, No. 9, pp. 2167–2178, 2001 # 2001 Elsevier Science Ltd. All rights reserved Printed in Great Britain 0043-1354/01/$ - see front matter
PHOSPHORUS REMOVAL FROM A REAL ANAEROBIC SUPERNATANT BY STRUVITE CRYSTALLIZATION P. BATTISTONI1*, A. DE ANGELIS1, P. PAVAN2, M. PRISCIANDARO3 and F. CECCHI4 1
Engineering Faculty, Institute of Hydraulics, University of Ancona, Via Brecce Bianche, 60131 Ancona, Italy; 2 Department of Environmental Science, Science Faculty, Calle Larga S. Marta 2137, 30123 Venice, Italy; 3 Department of Chemistry, Engineering Faculty, Chemical Engineering and Materials, Monteluco di Roio, 67040 L’Aquila, Italy and 4 Scientific and Technological Department, Science Faculty, Strada Le Grazie, 37134 Verona, Italy (First received 2 December 1999; accepted in revised form 19 September 2000) Abstract}In this paper the phosphorus removal from a real anaerobic supernatant through the crystallization of struvite and or hydroxyapatite was investigated. A comparison between experimental results on phosphorus crystallization carried out in a fluidized-bed reactor (FBR) on a bench-scale and on a half-scale plant is presented, together with a double saturational model able to describe all experimental results, independent of the different geometry of the reactors, the distinct contact times and the unlike products obtained. Experimental results show that removal efficiencies are very satisfactory, and the maximum phosphorus removal is of 80%. # 2001 Elsevier Science Ltd. All rights reserved Key words}phosphorus, struvite, crystallization, fluidized-bed reactor, waste water treatment
NOMENCLATURE
B C Em F H HRT IP J k KSP n N Q t tc t1=2 v V X Xf
collision efficiency in (equation (21)) concentration, mol/m3 maximum crystallization yield, dimensionless ratio defined as Qair =Qi , dimensionless hydraulic head, m hydraulic retention time, s ionic product of supersaturated solution collision frequency rate constant for crystallization constant of solubility product number of cycles in the reactor, dimensionless particle concentration in equation (21) flowrate, m3/h time, s contact time, s half-time, s velocity, m/s volume, m3 phosphate conversion (equation (3)), dimensionless precipitation efficiency (equation (2)), dimensionless
Greek symbols a constant in equation (24) b supersaturation, dimensionless e porosity, dimensionless Z crystallization efficiency (equation (1)), dimensionless Subscripts 0 initial value ag aggregation *Author to whom all correspondence should be addressed. Tel.: +39-071-220-4530; fax: +39-071-2204236; e-mail: [email protected]
air c exp f i o s ric stripp t
relative to the crystallization relative to the final value influent outlet soluble relative to the relative to the total
air expanded bed
recycle flowrate stripping section
Abbreviations HAP Ca5(PO4)3OH–hydroxyapatite FBR Fluidized-bed reactor MAP NH4MgPO4}struvite
INTRODUCTION
The crystallization of hydroxyapatite (CA5(PO4)3OH }HAP) and/or struvite (NH4MgPO4}MAP) represents, among the different methods to remove phosphorus from secondary effluent wastewater, a technique free of sludge-handling problems with respect to the conventional process of phosphate precipitation. The crystallization of these calcium and magnesium phosphate salts can be achieved with various plant configurations (Kaneko and Nakajima, 1988; Joko, 1984) but fluidized-bed reactors (FBR) appear to be of particular interest (Van Dijk and Braakensiek, 1985; Egger et al., 1991) and their utilization is widespread in the field of nutrient removal (Fujimoto et al., 1991; Momberg and Oellerman, 1992). Fluidized-bed pilot plants have
2167
2168
P. Battistoni et al.
so far been conducted using a base dosage to obtain the pH necessary to reach the condition of supersaturation, but considering the high buffer capacity of the supernatant, the process becomes in this way very expensive; furthermore, it is necessary to add calcium or magnesium to satisfy the stoichiometric demand or to select MAP or HAP formation (Fujimoto et al., 1991). This research group has devised a phosphate removal technique by exclusively using the chemical physical properties of anaerobic supernatants without any addition of chemicals, reaching the operative pH only by aeration (Cecchi et al., 1994; Battistoni et al., 1995, 1996, 1997). In particular, as shown by Battistoni et al. (1997), a quantitative removal of phosphate is obtained by matching crystallization of MAP on sand quartz and CO2 stripping with air. Previously, Cecchi et al. (1994) had in fact demonstrated that MAP crystallization allows to operate in the metastable zone and requires a lower pH with respect to HAP precipitation, thus obtaining phosphate removal without the addition of alkali reagents. In a recent work Battistoni et al. (1998a) investigated the performance of an FBR plant fed by a real supernatant of an A2O process supplied from a centrifugation dewatering station, showing that the supernatant does not require any pretreatment to remove the low suspended solid content. Furthermore, the efficiency of the process was related to the pH and sand contact time in a double saturational model. The crystallization of phosphorus was also studied by operating with an FBR managed in batch mode (Battistoni et al., 2000); the experimental work allowed to individuate the feasibility and best operative conditions for the process without the addiction of reagents, showing that a good phosphate removal could be obtained through crystallization on quartz sand in about 100 min. The batch process assured excellent conversions, but if in the low phosphorus concentration range (30–60 mg PO4/l) satisfactory results were obtained as far as crystallization removal, at high phosphorus concentration range (100–150 mg PO4/l) the process needs further investigation to limit the precipitation. It was concluded that the management of FBR in a batch mode is not a reliable solution. All these investigations have so far been carried out in an FBR on a bench scale, by using an anaerobic supernatant coming from the dewatering section of a 100.000 population-equivalent (PE) civil wastewater treatment. The results obtained showed a percentage removal of more than 80% with MAP and HAP formation. The present paper extends the same process to an FBR on half-scale, located in the civil wastewater treatment plant of Ancona (Italy, 100.000 P.E., A2O Process) and continuously fed by the anaerobic supernatant coming from a centrifugation sludge station, without any treatment.
MATERIALS AND METHODS
Materials Real anaerobic supernatants from a centrifugation sludge section of a 100.000 PE civil biological nutrient removal plant (BNR) were used. The treatment plant uses an A2O configuration (anaerobic, anoxic, oxic process) to perform the carbon oxidation and nitrogen nitrification–dentirification with biological phosphorus removal. Mixed primary and secondary sludges are gravitationally thickened and anaerobically digested before being fed to the dewatering station. Analysis The chemical-physical characteristics were determined using the Standard Methods (APHA 1985). In particular phosphates were analyzed on influent and effluent average acidified samples as total phosphate (PO4t) and on 0.45 mm filtrate as soluble phosphate (PO4s). NH4, Mg and Ca ions concentrations were analyzed on acidified samples, while alkalinity was analyzed on natural sample. On the basis of average values the following quantities can be calculated. *
Crystallization efficiency (Z 100) Z% ¼
*
ðPO4tot out PO4sol out Þ 100 PO4tot in
ðPO4tot in PO4sol out Þ 100 ð3Þ PO4tot in
Crystallization percentage of PO4–MAP on total removed PO4 PO4 MAP% ¼
*
ð2Þ
Phosphate conversion (X%) X% ¼ Z% þ Xf % ¼
*
ð1Þ
Precipitation efficiency (Xf %) Xf % ¼
*
ðPO4tot in PO4tot out Þ 100 PO4tot in
ðMgin Mgout Þmol 100 ðPO4tot in PO4tot out Þmol
ð4Þ
Crystallization percentage of PO4–HAP on total removed PO4 PO4 HAP% ¼
ðPO4tot in PO4tot out Þmol ðMgin Mgout Þmol 100 ðPO4tot in PO4tot out Þmol
ð5Þ
Moreover, the trends with time of the pH, total and soluble phosphates, calcium, magnesium and alkalinity (carbonates and bicarbonates) entering and leaving the plant, were estimated. Pilot-plant The FBR plant, as sketched in Fig. 1, comprises two sections, a stripping tank and a fluidized-bed reactor. The reactor is a perpex column ðfin ¼ 0:09 m, height 2 m, vol. 12.8 l) filled with 9.5 kg of virgin silica sand (0.21–0.35 mm, fm ¼ 0:265 mm) to obtain a compress bed of 1 m height. At the bottom of the column a perpex cylinder ðfin ¼ 0:09 m, height 0.5 m), filled with gravel with decreasing size distribution, behaving as a filter, blocks the sand from coming back towards the pumps, and allows an homogenous distribution of the stream to the reactor. At the top of the column an expansion tank is provided in order to prevent the sand from coming out of the reactor.
Phosphorus removal from real anaerobic supernatant
2169
Fig. 1. Layout of pilot plant.
A stripper and a connected deaeration column compose the stripping section. The anaerobic supernatant, as supplied from the real plant, is previously stocked in a tank (3.0 m3 volume) to let a daily continuos feeding, then it was fed to the stripper with the recycle flowrate and an air flowrate necessary for CO2 stripping. The effluent exists from the deaeration column together with the recycle flowrate. Three levels from which the effluent can outlet are provided, thus allowing to work with different hydraulic heads (H ¼ 0:5, 0.8, 1.0 m) and different volumes of stripping section and deaeration column. Temperature and pH probes were positioned in the deaeration column, while a second temperature probe was placed in air pipeline. The recycle pump has a flowrate ranging from 50 to 750 l/h. In Table 1 the main characteristics of FBR system are reported. The pilot plant was located in the dewatering station of a wastewater treatment real plant, the three sections are built on three different platforms to avoid useless surface request (see Fig. 2), it was conducted in a continuous mode. Five sets of experimental runs were performed; each set is characterized by a constant feeding flowrate of anaerobic supernatant while the air flowrate can be changed to operate at different pH. The ratio between recycle and influent flowrate can range from 0.6 to 1.4, the value adopted is that able to operate with a constant volume of expanded bed in all experiments (up to the expansion tank). Each run was monitored either by sampling incoming and effluent from feeding tanks and plant effluent, or by logging temperature and pH on a personal computer.
sand was then examined by means of differential thermal analysis, chemical analysis and grain-size distribution analysis. Microstructural analysis Morphological and microstructural analyses of the deposits on sand grains were performed by scanning electron microscopy (SEM) techniques (Phillips XL20 instrument). Energy-dispersive X-ray (EDX) spectrometry was used to analyze the elements and deposits (Edax PV 9800 instrument). Differential thermal analysis The differential thermal analysis was performed on exhaust sand grains by slowly increasing the temperature in order to find the presence of carbonates that, at T¼ 8008C, decompose by releasing CO2. The loss in weight of the sample allows to establish the percentage of carbonates initially present. Hydrodynamic of the process In the FBR system four hydraulic retention times can be identified, referring to Fig. 1: HRTt ¼
ðV1 þ V2 þ V3 Þ Qi
ð6Þ
ðV1 þ V2 Þ Qi
ð7Þ
Sand analysis Sand removed from FBR plant is air dried for 72–96 h, then weighted and analyzed to measure the porosity e. The
HRTstripp ¼
2170
P. Battistoni et al. Table 1. Volumetric characteristics of FBR
Apparatus
Characteristics and material
Volume, geometry and flowrate
Feeding tanks FBR reactor
PVC tanks Inner diameter 0.09 m Height 2 m Perpex column PVC Inner diameter 0.08 m Height 1.6 m PVC column Inner diameter 0.31 m Height 2.05 m PVC column PVC Power 0.25 kW Recycle pump Type: Progressive cavity Power max. 0.74 kW Feeding pumps Type: Peristaltic, Power 0.06 kW Progressive cavity power 2.0 kW Tair, T stripper, pH stripper
Variable volume up to 3.01 m3 Compressed-bed height: 1.5 m free volume 3.2 l 1.0 m 6.4 l 0.5 m 9.5 l Volume 18 l Hydraulic head 0.5, volume 2.7 l Hydraulic head 0.8, volume 4.2 l Hydraulic head 1.0, volume 5.2 l Hydraulic head 0.5, volume 40 l Hydraulic head 0.8, volume 62.6 l Hydraulic head 1.0, volume 77.7 l Volume 63 l
Expansion tank Deaeration column
Stripper E sampling tank B-blower Pl-
P2 Signal to personal computer
Flowrate 0.5–2.5 m3/h Peristaltic pump flowrate 0–100 l/h Peristaltic pump flowrate 0–150 l/h Progressive cavity pump flowrate 0–1500 l/h
RESULTS AND DISCUSSION
Fig. 2. Pilot organization.
V3 Qric
ð8Þ
Vexp e Qric
ð9Þ
HRTFBR ¼
HRTexp ¼
where Qi is the feed flowrate, Qric the recycle flowrate, V1 the stripping section volume, V2 the volume of the quiet zone, V3 the column free volume (including tube, expansion tank and filter volumes), Vexp the volume of the fluidized bed and e the bed porosity. In the upward movement of the liquid inside the FBR, the number of passages (n) is defined as the ratio between the mean hydraulic retention time in the plant and the time spent in the expanded bed, while the contact time (tc ) on sand grains depends on the number of cycles and the time needed for one single passage: HRTT HRTFBR
ð10Þ
tc ¼ nHRTexp
ð11Þ
n¼
Anaerobic supernatants are supplied from a centrifugation sludge station, without any previous treatment, and their averaged data are reported in Table 2. These data, compared with those obtained in the previous experimentation on bench-scale pilot plant (Battistoni et al., 1998a; see Table 3), allow to observe the lower content of total phosphates in the inlet stream (68 mg/l against 139 mg/l of the previous experimentation), the increased content of calcium, magnesium, ammonia and alkalinity and the ratio Ca/Mg almost one-half. It is worth noting that both the anaerobic supernatant used in this work and the anaerobic supernatant previously used (Battistoni et al., 1998a) are real liquors coming from an A2O process. The different composition is related to a change in technical management of the plant that happened a year later. Anyway, on the basis of the average PO4 content in supernatants, it is possible to note that the stoichiometric demand of calcium for HAP formation, or of calcium and magnesium for MAP formation, is completely satisfied. In the previous experimentation, due to the higher PO4 content, the stoichiometric demand was satisfied for HAP, whereas it was satisfied only to a limited extent for MAP (Mg/PO4=0.7 mol/mol, see Table 3). Five sets of experiments were carried out (A–Q Table 4a) characterized by different feed flowrates (17.6, 50–53, 67, 110 and 190 l/h), up to 10 times higher with respect to experimentation on laboratory scale. At the beginning of each experiment the feed flowrate, recycle flowrate and air flowrate were set to a fixed value. For each set of experiments, the air flowrate was modified in order to investigate the variations in the crystallization and precipitation efficiency while changing pH and the parameter F, given by the ratio between air and influent flowrates ðF ¼ Qair =Qi Þ. Each experiment lasted from 16 to 44 h, even though short durations were preferred,
Phosphorus removal from real anaerobic supernatant
2171
Table 2. Characteristics of liquors Parameters
Unit
Average
Min.
Max.
S.D.
n
pH PO4 NH4 Ca Mg HCO3 CO3 Ca/Mg Ca/PO4 Mg/PO4 NH4/PO4
mg/1 mg/1 mg/1 mg/1 mgCaCO3/1 mgCaCO3/1 mol/mol mol/mol mol/mol mol/mol
7.62 68 1055 200 54 7438 0 2.2 7.1 3.3 85.9
7.41 43 810 108 40 6400 0 1.6 4.2 1.9 60.41
7.87 89 1170 324 73 8050 0 3.0 13.3 5.4 133.7
0.11 14 92 50 11 393 0 0.5 2.4 1.1 21.6
22 22 22 22 22 22 22 22 22 22 22
Table 3. Comparison among liquors from laboratory and full-scale plants Plant scale
PO4t ot (mg/l)
NH4 (mg/l)
HCO3 (mgCaCO3/l)
Ca (mg/l)
Mg (mg/l)
Ca/Mg (mol/mol)
Ca/PO4 (mol/mol)
Mg/PO4 (mol/mol)
Laboratory Half
139 68
914 1055
3550 7400
153 200
24 54
3.7 2.2
2.6 7.2
0.7 3.3
Table 4. (a) Experimental results Set
Influent
Characteristics
Operative
Conditions
Ca Mg NH4 Qi time QAIR No. Run PO4 tot PO4 sol HCO3 (mg/l) (mg/l) (meq/l) (mg/l) (mg/l) (mg/l) (1/h) (h) (m3/h) 1 1 1 1 2 2 3 3 2 3 3 4 4 3 5
A B C D E F G H I L M N O P Q
58.2 41.3 43.7 45.3 40.4 42.6 56.7 74.1 82.1 83.0 75.4 68.5 73.5 73.5 57.5
57.4 31.0 39.4 41.3 34.5 38.3 53.2 30.0 54.2 53.4 58.1 52.5 59.0 48.4 32.2
69 72 76 77 73 73 77 68 70 70 72 76 76 75 69
132 113 139 156 181 183 194 214 216 205 206 205 203 199 195
41 42 63 60 61 61 57 61 45 43 44 47 47 48 42
1090 1130 1054 1135 1100 980 1160 990 972 1062 990 1030 998 1080 954
18 18 18 18 50 50 67 67 53 67 67 110 110 67 190
43.5 25.5 34.0 27.0 41.2 27.7 21.7 25.0 18.0 22.0 18.0 20.0 16.0 34.8 19.0
18.4 6.0 2.6 5.7 1.8 4.5 12.3 4.9 2.9 1.5 3.4 7.8 1.9 1.3 8.2
Effluent
F
tc (h)
pH
1047 339 145 324 36 94 183 74 54 22 51 71 18 19 43
1.43 1.41 1.32 1.31 0.43 0.39 0.26 0.23 0.47 0.37 0.35 0.20 0.19 0.30 0.09
8.78 8.55 8.55 8.47 8.07 8.45 8.68 8.45 8.31 8.06 8.30 8.44 8.07 8.04 8.22
Performances
PO4tot out PO4s out Z (mg/1) (mg/1) (%) 11.6 11.6 12.1 10.1 17.0 13.3 12.7 25.9 26.4 35.5 27.3 23.8 34.5 30.7 26.5
11.1 7.9 11.6 9.8 16.7 12.9 12.2 23.0 25.6 34.4 27.0 21.9 32.0 28.5 23.5
80 74 74 78 55 72 78 65 68 57 63 65 53 58 55
X (%)
Xf (%)
84 81 73 81 72 91 82 76 71 60 64 73 60 65 66
4 7 1 3 2 3 4 11 3 3 1 8 7 17 1
(b) Experimental results}HAP and MAP formation No. A B C D E F G H I L M N O P Q a b
DPO4 (mg/l)
DCa (mg/l)
DMg (mg/l)
Dalkal. (mg/l)
PO4HAPa (%)
PO4MAPa (%)
HAPb (%)
MAPb (%)
CaCO3b (%)
46.6 29.7 31.6 35.2 23.7 29.3 44 48.2 55.7 47.5 48.1 44.7 39 42.8 31
110.2 94.6 45.9 94.2 95.5 117.4 143.2 112.5 87.9 51.5 96.7 112.7 31.1 55 88.9
4.7 2.33 1.2 7.9 6.7 9.2 11.5 13.3 9.6 7 10.9 11.8 7.7 8.3 7.2
7.41 16.5 14.5 15.5 7.3 12.9 13.9 7.5 5 3.3 6.1 7.4 4.2 6.3 5.6
59 68 85 8 0 0 0 0 30 40 8 0 20 24 5
41 32 15 92 100 100 100 100 70 60 92 100 80 79 95
4 3 11 0 0 0 0 0 2 5 0 0 3 2 0
8 5 6 13 11 12 12 17 17 23 17 15 33 22 12
88 92 83 87 89 88 88 83 80 72 83 85 64 76 87
Percentage relative to phosphorus distribution. Molar ratios on total.
since dewatering station was open only 8 h a day. The five main sets of experiments, classified on the basis of feed flowrates, allow to study 16 operation conditions, different for pH and contact time. Among
various experiments, variability in supernatants total phosphorus content is observed, probably due to the different way of operating of the post-thickener in real plant. Moreover, during each run, a decrease in
2172
P. Battistoni et al.
phosphorus soluble fraction ðPO4s =PO4tot Þ is found. This circumstance is mainly dependent on supernatant residence times in the feed tank (see Fig. 1). In fact, the aging of anaerobic supernatants results in a spontaneous precipitation of HAP and/or MAP, and therefore in a decrease of PO4s/PO4 tot as the supernatant comes to equilibrium with the air. This fact was observed both for anaerobic supernatants coming from nitrogen biological removal wastewater treatment plants added with orthophosphates; Battistoni et al., 1997) and for supernatants coming from nutrients biological removal wastewater treatment plants (Battistoni et al., 1998b); in this last case, supernatants are similar to those used in the present study. An analysis of runs shows that in all of them, crystallization efficiencies vary from 53 to 80%, the precipitation efficiency is normally very low: in 9 of the 15 runs precipitation efficiency is less than the cut-off value chosen to integrate the plant with a post-filtration operation, while in three of the 15 runs precipitation efficiency ranges from 7 to 8% and in three of the 15 runs it is higher than 10%. Each run has a typical behavior, made by a first phase (lasting up to 3 h, depending on F value and on hydraulic level in the deaeration column), in which pH exponentially grows from supernatant to experimental value, and a stationary phase, in which the average pH at 300 remains on constant. Variations in influent characteristics depend on stock or on feed renewal when the run is working. In Fig. 3(a), a typical behavior of operating parameters for D run is reported; in particular, pH at the stationary phase, after 2.5 h is equal to 8.46 (S.D. 0.04), and the initial concentration of PO4 sol is constant, while slight variations in PO4 tot are present at the beginning, at half and at the end of the run, all below 5 mg/l. At the end of the run a variation is present in PO4 tot out and PO4 sol out, and moreover pH is varied by 0.05% with respect to the previous value. The change in pH value reported in Fig. 3(b) is ascribed to the higher value of alkalinity, and therefore to the buffer effect. In conclusion, the alkalinity loss between influent and effluent is observable from the periodic check of their chemical-physical characteristics (Fig. 3(b)). Salt precipitation Several salts may form from a P-enriched or not real anaerobic supernatant; when a certain amount of magnesium is present, only MAP can be obtained in the presence of ammonium and phosphate ions (Seckler et al., 1991), since the formation of Mg(OH)2 was found only operating in a bed of MgO (Angel, 1999), or at high values of pH (ShulzeRettmer, 1991); in particular, when pH10.5 the retarding effect of Mg ions on phosphate formation tends to disappear (Jenkins et al., 1971). Moreover, the coprecipitation of Mg3(PO4)2 with amorphous Ca3(PO4)2 was observed in clear waters added with Ca, PO4 and Mg ions (Seckler et al., 1991, 1996b). On
Fig. 3. Typical behavior of (a) phosphates and pH with time (run D), and (b) alkalinity (run D).
the other hand, different salts can precipitate when calcium is present, namely dicalcium phosphate (CaHPO4), tricalcium phosphate (Ca3(PO4)2), octocalcium phosphate (Ca4H(PO4)3) and apatites (Zoltek, 1974). The formation of these phosphate salts seems to conform to the Ostwald–Gay Lussac Step Rule, which states that the less thermodynamically stable, and hence the more soluble, form of a polymorphous solid is generally the most readily precipitable; therefore the likely order of formation would be dicalcium phosphate, tricalcium phosphate, octocalcium phosphate and finally, the apatites (Zoltek, 1974). Amorphous calcium phosphates are always present in clear water in supersaturated conditions (Seckler et al., 1996a, b), or sometimes in the main effluent stream of a civil wastewater treatment plant (Van Dijk and Braakensiek, 1985; Angel, 1999), while crystalline HAP was found working in the metastable zone in effluents from lagoons, secondary sedimentation tanks or anaerobic supernatants from sludge treatment devices (Momberg et al., 1992; Joko, 1984). Finally, HAP
Phosphorus removal from real anaerobic supernatant
crystallization was observed in synthetic liquor also, when a low phosphorous concentration was present (Kaneko et al., 1998; Joko, 1984). As regards the socalled calcite (CaCO3), carbonate ions have a tendency to crystallize as calcite when the stoichiometric ratio Mg/Ca is less than 0.6 (as in the current work, see Table 3); otherwise as dolomite CaMg(CO3)2 when Mg/Ca is higher than 0.6 (Diaz et al., 1994). In conclusion, in the present work experimental results will be analyzed considering the formation of merely HAP, MAP and calcite. Other kinds of salts will not be taken into consideration, giving to the analysis of the exhaust grains the task of confirming what was supposed. Moreover, in a previous work carried out on the same supernatant in a bench-scale FBR (Battistoni et al., 1998a), the analysis of sand after the crystallization process demonstrated the presence of HAP and MAP only. As for the molar losses, unlike results previously obtained (Battistoni et al., 1998a), together with MAP formation the presence of a great deal of calcite (CaCO3) was found (see Table 4b). The analysis of the table reveals the MAP formation is predominant in almost all experiments. This can be related to the Ca/Mg ratio (2.2}Table 3) lower than that previously found (3.7}Table 3) and a Mg/PO4 ratio equal to 3.3. However, these ratios do not explain the calcite crystallization because the molar ratio Ca/PO4 is enough to guarantee a complete HAP formation. A possible explanation of the phenomena can be found in the high alkalinity and Mg content and their inhibitory effect on HAP formation (Battistoni et al., 1998b; Joko, 1984; Ferguson et al., 1970, 1973; Jenkins et al., 1971). This circumstance could be explained by assuming that Ca precipitates as salt different from HAP, determining the prevailing MAP formation despite its massive presence in the supernatant, this hypothesis being confirmed by the analysis of the exhaust grain sand reported in subsequent section. Operative pH The operative pH can be approached by fixing an air flowrate, and therefore a value for F, which is related to pH by means of a saturational relationship. For experiments carried out on laboratory scale the data can be modeled (multiple regression analysis}Anova method) according to the following equation: pH ¼ 1:125
F þ 7:70 171 þ F
2173
between pH and F given by the following expression: pH ¼ 1:100
F þ 7:63 33 þ F
with R2 ¼ 0:89 and S.E.=0.08. Once the value of air flowrate, and therefore of F, has been fixed, the value of pH in condition of saturation depends on the hydraulic head, and on whether FBR is active or not. In particular, in Fig. 4 first pH increase, when only the stripping section is active, and a second increase, when the whole plant is working, are shown. Three tests are performed at a constant feeding flowrate of 621/h and an F ¼ 40 l air/l feeding, while the hydraulic head changes from 0.5 to 0.8 and 1.0 m. The main conclusions that can be drawn are: *
*
a higher hydraulic head allows to obtain a higher pH, even if a head of 0.8 m is enough to reach the operative conditions; when FBR also works, a pH higher than that obtained when operating with the stripping section only, can be obtained.
This behavior can be interpreted considering that a higher hydraulic head results in a major use of the diffused air flowrate, a lower buffer capacity of waste when calcite crystallizes in the FBR and an increased CO2 stripping by recirulation pump; however, up to 3 h are required to obtain the operative pH (steadystate conditions). Consequently, in order to make a comparison among different reactors, the only parameter to be taken into account is the pH. Crystallization efficiency Efficiency is strongly related to the operative pH, both for experiments carried out on laboratory scale (Battistoni et al., 1998a) and for experiments on a half scale; in particular, in this last case the highest efficiency, equal to 80%, occurred at the highest pH, equal to 8.78 (Table 4a, run A). This strong dependence is shown in Fig. 5, and on this basis an empirical correlation, linear or saturational, between
ð12Þ
with R2 ¼ 0:41 and S.E.=0.13. The low value obtained for R2 is due to the different kinds of feed used, namely liquor without any additives, liquor with calcium ions added up, liquor with magnesium ions added. On the contrary, the experimental results relative to the half-scale plant individuate a relationship
ð13Þ
Fig. 4. Dependence of pH on hydraulic head.
2174
P. Battistoni et al. Table 5. Statistical parameters obtained with multiple regression analysis ANOVA Equation
R2
S.E.
(14) (15) (16) (17)
0.81 0.85 0.79 0.81
0.07 0.10 3.03 4.05
Fig. 5. Crystallization efficiency as a function of pHexperimental results.
pH and Z was found for each plant configuration, (linear and multiple regression, Anova method) expressed by the following equations: pH ¼ 6:62 þ 0:022Z%
bench scale
pH ¼ 6:81 þ 0:023 Z% half scale
ð14Þ ð15Þ
Z% ¼ 100
ðpH 7:65Þ bench scale ð16Þ ðpH 7:65Þ þ 0:201
Z% ¼ 100
ðpH 7:65Þ ðpH 7:65Þ þ 0:340
half scale ð17Þ
with the statistical parameters obtained with multiple regression analysis ANOVA reported in Table 5. Unfortunately, this first linear model does not explain much more than 85% of the result variance. Furthermore, this saturational model at a deeper insight appeared to be not sufficiently adequate to describe the dependence of crystallization efficiency on process parameters, since a strong dependence of Z not only on pH, but also on another variable, was found. In Fig. 6 the behavior of efficiency with contact time (tc ) is reported for both plant configurations, showing that the contact time on sand grains represents a major parameter in the characterization of process performances. A double-saturational model, taking into account both pH and tc , was devised, with a multiple regression (Anova method) which exhibits an excellent agreement with experimental results. For the experiments carried out in the half plant, the model is expressed by the following equation: Z% ¼ 100
ðpH 7:65Þ tc ðpH 7:65Þ þ 0:290 tc þ 0:017
ð18Þ
with R2 ¼ 99:84 and S.E.=2.76. The double-saturational model was originally presented for laboratory-scale experiments (Battistoni et al., 1998a), and even in that case it allowed a
Fig. 6. Crystallization efficiency as a function of contact time-experimental results.
successful interpretation of experimental data. The versatility of the model, able to describe experimental results carried out in very unlike reactors, relies on the fact that it is well sustained by hydraulic and engineering parameters, as defined by equations (6)–(9). Therefore experimental results, obtained in different geometry reactors, with distinct contact times and giving unlike products can be accurately described by using the same mathematical model, as demonstrated by the following equation: Z% ¼ 100
ðpH 7:63Þ tc ðpH 7:63Þ þ 0:208 tc þ 0:039
ð19Þ
with R2 ¼ 99:73 and S.E.=3.78. Equation (17) has been in fact obtained by correlating experimental data relative to both bench-scale and half-scale pilot plant configurations. The half-efficiency constants for pH and tc introduced in equation (18) give useful information: an operative pH of 7.94 (calculated as 0.29+7.65, see equation (18)) reduces the efficiency by half, and the same result is obtained when operating at a contact time equal to 0.017 h (about 1 min). The contact time is defined by equation (11) as a function of the number of cycles in the reactor, but for the same tc different n are feasible since e changes with phosphate growth on sand and consequently HRTexp changes even if Qric and Vexp remain constant. In fact, e varies from a maximum of 0.655 to a minimum of 0.405 as shown in Table 6, passing from virgin to exhaust
Phosphorus removal from real anaerobic supernatant
2175
Table 6. Porosity and contact times in FBR experiments No. A B C D E F G H I L M N O P Q a
e
tc (h)
0.655 0.641 0.594 0.583 0.546 0.489 0.431 0.365a 0.653 0.634 0.611 0.568 0.530 0.505 0.405
1.4326 1.4076 1.3198 1.3070 0.4324 0.3937 0.2633 0.2275 0.4745 0.3662 0.3548 0.2033 0.1917 0.3017 0.0878
Exhaust sand substituted with virgin sand.
Fig. 7. Crystallization efficiency as a function of contact time, double saturational model (equation (18)).
sand. Definitely, the contact time can be expressed with the following expression: HRTt tc ¼ n HRTexp: ¼ HRTexp HRTFBR ðV1 þ V2 þ V3 Þ Qric eVexp aeVtot ¼ ¼ V3 Qric Qi Qi
ð20Þ
where a is a dimensionless constant, used to express the column free volume (V3 ) as a function of the expanded bed volume (Vexp ). Following the definition given by equation (20), the contact time appears to be the best operative parameter to be considered for comparing results from different pilot plants. The introduction of the double-saturational model (equation (18)) in a pH range which can be easily gained (8.1–8.7) allows to fix the most suitable contact time to reach in order to get a prefixed efficiency. As shown in Fig. 7, it can be seen that for a pH equal to 8.5, in order to get an efficiency Z of 70%, a tc equal to 0.3 h is needed, while efficiencies of 80% are very difficult to get with a pH less than 8.7. Several models are available in the literature to describe the phosphorus precipitation process; they can be distinguished in models based on primary nucleation mechanism (e.g. nucleation is caused by pure supersaturation), and models based on secondary nucleation mechanisms (e.g. nucleation and growth take place on pre-existent seeds, in metastable supersaturation conditions). The work of Seckler (Seckler et al., 1996a–c) belongs to the first group, proposing a theoretical model for fine-particle aggregation with sand grains. This model takes into account both aggregations of fines with sand and particle breakage when a high dissipation of the kinetic energy takes place. The following equation derives from a particle number balance, and expresses the decrease in the fine concentration by aggregation with grains in a fluidized bed: dNi ¼ BJii b dt
ð21Þ
where N is the initial particle concentration, Jii is the collision frequency, b is the supersaturation, t is the reaction time and B is the collision efficiency, which is influenced by the energy dissipation rate. By integrating equation (21) from t ¼ 0 to t ¼ tout , the phosphate removal efficiency Zag (by aggregation only) can be calculated, by assuming the particle size of the fines to be constant in time: Zag ¼
Ni;in Ni;out Ni;in
ð22Þ
It was found, both theoretically and experimentally, that the aggregation could be increased by spreading the supersaturation more evenly throughout the reactor, while choosing fluidization conditions where the energy dissipation rate in the bed is minimized can reduce the breakage. This process has been studied in the precipitation of phosphorus from tap water or from a phosphate-enriched stream of a sewage treatment plant. Anywhere, the model describes experimental results only to a limited extent, since the aggregation of fines can be strongly influenced by molecular growth. Regarding the models based on secondary nucleation mechanism, De Rooij et al. (1984) studied the formation of different phases of calcium phosphates on a seed material, in well-defined experimental conditions, such as a fixed value of temperature (T ¼ 378C), ionic strength (I=0.10 mol/l), Ca/P ratio (Ca/P=1.333) and for different pH values (55pH58). In order to avoid any interference by traces of impurities in reagent-grade chemicals, ultrapure reagents were used. The model obtained is that shown by the following equation: p dn 1=v ¼ ks IP1=v KSP ð23Þ dt where n is the number of precipitated moles, k the rate constant for crystallization, s a factor proportional to the number of sites available for growth, p
2176
P. Battistoni et al.
the effective order of reaction, v the number of ions in the formula unit, IP the ionic product of supersaturated solution and KSP the thermodynamic solubility product. The Gibbs free energy related to the transfer from supersaturated solution to an assumed saturated solution at the surface of the developing solid phase, gives the relationship between IP and KSO (De Rooij et al., 1984). Kaneko and Nakajima (1988), in order to describe the crystal growth of HAP, starting from a metastability condition of synthetic water solutions in which the dependence of pH is not present, made an approximation of equation (23) expressed as dC ¼ ksC2 ð24Þ dt where C is the molar concentration and t is the retention time. Since it was demonstrated that for the crystallization of MAP and HAP from anaerobic supernatant in fixed-bed reactors FBR works in conditions of metastability (Battistoni et al., 1998b), the approximation made by Kaneko and Nakajima (1988) is suitable to be used. This kinetic equation can be integrated from 0 to t, leading to the empirical saturational model in tc (equation (18)) of crystallization efficiency Z (Battistoni et al., 2000a): Z ¼ Em
tc t1=2 þ tc
ð25Þ
Moreover, the maximum crystallization efficiency (Em ) and the half-time (t1=2 ) can be introduced: C0 Ct C0
ð26Þ
1 ðC0 Ct Þk
ð27Þ
Em ¼ t1=2 ¼
where Ct is the final concentration, C0 the initial concentration and tc the contact time. The empirical double saturational model thus follows from the theoretical model of crystallization in conditions of limiting supersaturation for what concerns the saturation in tc , while the behavior as a function of pH has always been reported in the literature but never described in a theoretical model. Experimental runs (see Table 4) show that an almost wide range of precipitation efficiency was individuated, for which no explanation is given. In conclusion, with respect to a previous work (Battistoni et al., 1998a), in this paper it has been possible to reduce the contact time with a higher pH by increasing the air-to-influent ratio. A further experimental work is required to improve the validity of the model proposed, even at a lower contact time.
analysis of the grain; comparison among the total amount of phosphorus removed calculated through a mass balance carried out on experimental results and exhaust matrix; grain size distribution on virgin and exhaust silica sands; thermal differential analysis on grains and scanning electron microscopy analysis. Sand samples were taken from FBR in the middle (M) and at the end of the experimentation (E). The chemical analysis highlighted a first difference with respect to bench-scale experiments (Battistoni et al., 1998): in that case phosphorus was removed entirely as HAP (62%) and MAP (38%), while in this case the percentage of MAP ranges from 34% on weight referred to the grain growth (M) to 50% (E) and HAP varies from 0 (M) to 5% (E). Truly, on grain surface a great deal of calcite (CaCO3) has formed (66% (M)–45% (E)), subtracting Ca ions for HAP cyrstallization. MAP formation is thus explained despite the high value of the ratio Ca/Mg. A comparison between the amount of phosphorus present in matrix and that removed, computed as the total of all the phosphorus crystallized in the different runs, allowed to confirm the worth of process, since all phosphorus removed from the liquid phase was found in the reactor solid matrix. This circumstance strengthens the validity of equations introduced to define the crystallization and precipitation efficiencies (see equations (1) and (2)) since a significant precipitation in the stripping and deaeration section is not expected. In Table 7 the increment of grain average dimensions obtained through a grain size distribution is reported: the nominal diameter passes from 0.27 mm of virgin sand to 0.35 mm of exhaust sand, with a growth of 26%, on which the presence of calcite had a major role. The differential thermal analysis reveals the classic weight loss at 8008C temperature connected with the presence of carbonate in exhaust sand. In particular, the conversion of CO2 loss into CaCO3 gives an amount equivalent to that calculated from chemical analysis of exhaust sand where the Ca ions in excess to HAP formation are computed as CaCO3 formation. This evidence is repeated either in M or in E sand samples. The SEM analysis interpretation requires all the information obtained by previous techniques. The photograph of an exhaust grain (M condition} Fig. 7) shows the morphological aspect of exhaust sand with different zones classified as sheet (marker A) and sphere (marker B) types. X-ray dispersive
Table 7. Silica sand nominal diameter
Analysis of the exhaust sand grain
Sand type
d (mm)
Some features of the process were clarified through an analysis of the exhaust sand grain, which was carried out according to the following steps: chemical
Virgin sand Exhaust sand (M) Exhaust sand (E)
0.2672 0.3156 0.3478
Phosphorus removal from real anaerobic supernatant
has been presented. The conclusions arising from this comparison can be summarized as follows:
Table 8. Elements on deposit (EDX measurement) Zone
P (%)
Mg (%)
Ca (%)
35.0 61.4 13.6
18.5 32.2 7.2
46.5 6.4 79.1
*
Average Sheet Sphere
2177
*
*
*
*
Fig. 8. Morphological analysis of sand grain.
spectrometry was performed in both zones and on middle one with dimensions of 200 200 mm. The results (Table 8) do not highlight trace of Si. This means that EDX spectrometry is not able to investigate the whole deposits, but only a fraction and most probably the last. This agrees with the average thickness of deposit equivalent to 40.3 mm (see Table 8). Furthermore, the elements in different zones reveal that the sheet deposit is mainly formed by MAP while the sphere zones composition can be translated in a MAP (15%), HAP (2%) and CaCO3 (83%) molar distribution, this last being very similar to those calculated in P and Q experiments on FBR plant (see Fig. 8 and Table 4b). Coming to conclusions, in spite of all investigations carried out on grains, a clear evidence whether the deposit on sand is Ca3(PO4)2 or HAP was not achieved. Analytical mistakes in chemical analysis may be sufficient to cover the slight difference between the formations of one species with respect to another. This is the reason why X-ray spectroscopy should be desirable, in order to ascertain which kind of salt is formed. Notwithstanding the process modeling is still well founded, since it has been carried out on phosphorus removal, independent of the crystallized salt.
CONCLUSIONS
In this paper a comparison between experimental results on phosphorus crystallization carried out in an FBR reactor on a bench-scale and on half-scale
the double saturational model is able to describe all experimental results, independent of the different geometry of the reactors, the distinct contact times and the unlike products obtained; a further experimental work is required to further improve the validity of the model proposed, even at a lower contact time; with respect to bench-scale experiments in half-scale experiments phosphorus is removed almost entirely as MAP and a great deal of calcite forms on grains; a further investigation is needed to foresee the formation of MAP or HAP on silica sand grains; also in the case of plant on a half scale the efficiencies are very satisfactory, and the maximum phosphorus removal is of 80%.
Acknowledgements}The authors would like to thank the Ministry of the University and of the Scientific and Technology Research of Italy for financial support.
REFERENCES
Angel R. (1999) Removal of phosphate from sewage as amorphous calcium phosphate. Environ. Technol. 20, 709–720. APHA, AWWA and WPCF (1985) Standard Methods for the Examination of Water and Wastewater, 16th ed. American Public Health Association, Wastington, DC. Battistoni P., Fava G., Pavan P., Musacco A. and Cecchi F. (1995) Phosphate crystallization in a fluidized bed reactor. Preliminary results. Second International Symposium on Wastewater Reclamation and Reuse, Iraklio, Crete, Greece, October 17–20. Battistoni P., Fava G., Pavan P. and Mata-Halvarez J. (1996) Phosphate removal by crystallization. Anaerobic liquors treated in a fluidized bed reactor. Proceedings of Water Quality International 96, Singapore, June 23–28. Battistoni P., Fava G., Pavan P., Musacco A. and Cecchi F. (1997) Phosphate removal in anaerobic liquors by struvite crystallization without addition of chemicals. Preliminary results. Water Res. 31, 2925–2929. Battistoni P., Pavan P., Cecchi F. and Mata-Halvarez J. (1998a) Phosphate removal in real anaerobic supernatants: modeling and performance of a fluidized bed reactor. Water Sci. Technol. 31, 275–283. Battistoni P., Pavan P., Cecchi F. and Mata-Alvarez J. (1998b) Effect of composition of anaerobic supernatants from an anaerobic, anoxic and oxic (A2O) process on struvite and hydroxyapatite formation. Ann. Chim. 88, 761–772. Battistoni P., Pavan P., Prisciandaro M. and Cecchi F. (2000) Struvite crystallization: a feasible and reliable way to fix phosphorus in anaerobic supernatants. Water Res. 34, 3033–3041. Cecchi F., Battistoni P., Pavan P., Fava G. and MataHalvarez J. (1994) Anaerobic digestion of OFMSW and BNR processes: a possible integration. Preliminary results. In Proceedings of the 17th Biennal Conference, Water Sci. Technol. 30, 65–71.
2178
P. Battistoni et al.
De Roij J. F., Heughebaert J. C. and Nancollas G. H. (1984) A pH study of calcium phosphate seeded precipitation. J. Colloid Interface Sci. 100, 350–358. Diaz O. A., Reddy K. R. and Moore P. A. (1994) Solubility of inorganic phosphorus in stream water as influenced by pH and calcium concentration. Water Res. 28, 1755– 1763. Egger E., Dirkzwager A. H. and Van der Honing H. (1991) Full-scale experiences with phosphate crystallization in a crystalactor. Water Sci. Technol. 23, 819–824. Ferguson J. F., Jenkins D. and Stumm W. (1970) Calcium phosphate precipitation in wastewater. Chem. Prog. Symp. Ser. 67, 279–287. Ferguson J. F., Jenkins D. and Fasrman J. (1973) Calcium phosphate precipitation at slightly alkaline pH values. J. Water Pollut. Control Fed. 45, 620–631. Fujimoto N., Mizouchi T. and Togami Y. (1991) Phosphorus fixation in the sludge treatment system of a biological phosphate removal-accomplishments and needs. Water Res. 25, 1471–1478. Jenkins D., Ferguson J. F. and Menar A. B. (1971) Chemical process for phosphate removal. Review paper. Water Res. 5, 369–389. Joko I. (1984) Phosphorus removal from wastewater by the crystallization method. Water Sci. Technol. 17, 121–132. Kaneko S. and Nakajima K. (1988) Phosphorus removal by crystallization using a granular activated magnesia clinker. J. Water Pollut. Control Fed. 60, 1239–1244.
Momberg G. A. and Oellerman R. A. (1992) The removal of phosphate by hydroxyapatite and struvite crystallization in South Africa. Water Sci. Technol. 26, 967–976. Schulze-Rettmer R. (1991) The simultaneous chemical precipitation of ammonium and phosphate in the form of magnesium–ammonium–phosphate. Water Sci. Technol. 23, 659–667. Seckler M. M., Bruinsma O. S. L. and Van Rosmalen G. M. (1991) Crystallization of calcium and magnesium phosphates in a fluidized bed. Proceedings of the Third European Conference on Crystal Growth, Budapest, Hungary, pp. 932–937. Seckler M. M., Van Leeuwen M. L. J., Bruinsma O. S. L. and Van Rosmalen G. M. (1996a) Phosphate removal in a fluidized bed. II. Process optimization. Water Res. 30, 1589–1596. Seckler M. M., Bruinsma O. S. L. and Van Rosmalen G. M. (1996b) Calcium phosphate precipitation in a fluidized bed in relation to process conditions: a black box approach. Water Res. 30, 1677–1685. Seckler M. M., Bruinsma O. S. L. and Van Rosmalen G. M. (1996c) Phosphate removal in a fluidized bed}I. Identification of physical processes. Water Res. 30, 1585–1588. Van Dijk J. C. and Braakensiek H. (1985) Phosphate removal by crystallization in a fluidized bed. Water Sci. Technol. 17, 133–142. Zoltek J. (1974) Phosphorus removal by orthophosphate nucleation. J. Water Pollut Control Fed. 46, 2498–2520.