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Microfiltration and ultrafiltration as a post-treatment of biogas plant digestates for producing concentrated fertilizers

Biogas plant digestate liquid fractions can be concentrated by microfiltration and ultrafiltration. Two types of microfiltration membranes (polysulphone (PS) and surface-modified polyvinylidene fluoride (PVDF)) were used to process digestate liquid fractions, and to assess their applicability in the recovery of particulate phosphorus, compared to an ultrafiltration membrane (polyethersulphone (PES)). Results show that membrane material, operational conditions, and pore diameter influenced the permeate flux pattern during microfiltration. The PS membranes initially had a higher tendency to foul than PVDF membranes. However, during the filtration process, as fouling built up, the permeate flux behavior of the two membranes became very similar. During the concentration of digestate liquid fractions, the microfiltration PS membrane and the ultrafiltration PES membrane achieved the highest phosphorus rejection (80% w/w), suggesting that there was a correlation between the membrane material and both the fouling trend and phosphorus rejection. A two-step basic-acidic cleaning was unable to recover the initial water flux for the fouled microfiltration membranes. In conclusion, the PS microfiltration membranes might be a good strategy for recovering phosphorus from digestate liquid fractions. Further research leading to adequate cleaning procedures, for microfiltration PS and PVDF membranes treating digestate liquid fractions though, are needed.


27
Livestock production has increased dramatically in the last decades due to an increased 28 consumption in meat and an increase in the human population [1]. As large centralized livestock 29 production leads to large production of manure, centralized biogas production by anaerobic 30 digestion is of growing interest as a green energy solution in areas with industrialized livestock 31 production. Anaerobic digestion stabilizes and converts about 20-25% of the organic matter present 32 in animal manure into biogas; it is considered a reliable source of bioenergy and it minimizes 33 greenhouse gas (GHG) emissions and odors [2][3][4]. Anaerobic digestion mineralizes some of the 34 organic nitrogen and increases the NH 3 /NH 4 + content by about 15%. The same occurs with the 35 organic phosphorous, the concentration of which in the inorganic form also increases after 36 anaerobic digestion [4]. Digestates contain large amounts of nutrients in the form of nitrogen, 37 phosphorus, potassium, organic matter and micronutrients (e.g. calcium, chloride, sodium, copper, 38 zinc, manganese, etc.) [5]. It is common practice to use digestates as a fertilizer source and to spread 39 the resulting liquid fraction after solid-liquid separation on the fields [6]. However, this can lead to 40 environmental problems mostly related to NH 3 volatilization as well as surface and groundwater 41 eutrophication [1,7]. Moreover, digestates present an unbalanced distribution of nutrients with 42 respect to crop uptake and demand, which makes their application as reliable fertilizers difficult. 43 When digestates from centralized biogas plants are applied as fertilizer to fields far away from the 44 biogas plant, it is advantageous to separate digestate in a solid fraction rich in phosphorus and a 45 liquid fraction rich in nitrogen and potassium [8], but low in phosphorus. Separation in a solid and 46 liquid fraction both reduces transportation costs and makes it possible to obtain a more balanced 47 fertilizer composition. Traditionally, this separation is done using either a decanter centrifuge or a 48 However, the study of the fouling characteristics and filtration performance of the ultrafiltration 99 system is only included for comparison. It is not the objective of this paper. 100

Biogas plant digestates 102
Microfiltration experiments were carried out on the anaerobic digestate liquid fraction from Fangel 103 Biogas (Odense, Denmark). The feed to the anaerobic digester consisted of approximately 50% of pig 104 slurry, 15% cattle manure, 10% chicken manure and 25% food waste. The anaerobic digestion was 105 mesophilic (39-40 o C) and the hydraulic retention time was 30 days. The final digestate was 106 centrifuged via a decanter centrifuge (AD-1220, GEA Westfalia, Germany). The obtained solid 107 fraction is normally composted while the liquid fraction is distributed to local farmers and used as a 108 liquid fertilizer. Fangel biogas plant treats 260 tonnes·day -1 of animal slurry and 60 tonnes·day -1 of 109 industrial food waste. This allows an annual energy production of 16-18 million kWh·year -1 . The 110 digestate liquid fraction obtained after centrifugation was sieved using a 2-step manual sieve (Retsch 111 5657, Germany) of 350 and 125 µm mesh size. The resulting sieved liquid fraction was used as feed 112 to the microfiltration unit. The composition and characteristics of the digestate liquid fraction used 113 as feed to the microfiltration system is shown in Table 1. 114 The ultrafiltration experiments were carried out at Bioscan A/S' anaerobic digester ultrafiltration 115 pilot plant at mesophilic conditions (35°C) and at a hydraulic retention time of 12 days. The 116 concentrate was recirculated to the digester [13]. The digestate had a pH of 8 and a composition of 117 3.2% total solids; 2.7 (g·L -1 ) of total ammoniacal nitrogen (TAN); 0.53 g·L -1 of phosphorus and 0.46 118 g·L -1 of potassium. 119

Microfiltration unit 121
The microfiltration experiments were carried out on a LabStak® M10 laboratory plant (Alfa Laval, 122 Denmark) (Figure 1). The M10 is a cross flow filtration module with an active membrane area of 123 0.036 m 2 which can accommodate four membrane flat sheets (see Table 2). The laboratory plant had 124 three M10 membrane modules placed in parallel. This made it possible to test three different 125 membrane pore sizes at nearly identical operational conditions. The feed was pumped from one 126 shared 15L-feed tank through an external spur gear pump (GC6-KDT-KKU, Pulsafeeder Inc. USA) with 127 a maximum flow rate of 2.8 m 3 ·h -1 . Piping and connections in the plant were made of stainless steel 128 AISI 316. The inflow velocity to the microfiltration module was measured through a digital flowmeter 129 (IFS1000K, Krohne Altometer, Germany) and the transmembrane pressure by analogue pressure 130 gauges P1 and P2 (see Figure 1) (Tempress, Denmark). The temperature was controlled through a 131 multitube heat exchanger (Alfa Laval, Sweden). All the experiments were run at 30 o C, chosen as a 132 relevant temperature for mesophillic digestate liquid fractions obtained after digestate 133 centrifugation. 134

Ultrafiltration unit 135
The ultrafiltration experiments were performed by Bioscan A/S (Denmark) at their pilot plant ( Figure  136 2) unit. PES tubular membranes (STUF 2.5/1.7 AL (S) 719, Membratek, South Africa) having an active 137 area of 7.1 m 2 and a cut-off of 40 kDa. The system was run at 1.5-3.5 bar and a cross flow velocity of 138 2, 2.6, 2.9 and 4.3 m·s -1 . 139 The anaerobic digester was fed with screw-press pre-treated slurry from a pig farm. experiments. Before starting a filtration cycle, clean water tests were performed on the fresh 157 membranes, using deionized water at the same operational conditions as the related filtration 158 experiment. Each filtration experiment lasted 3-4 h. Following, another clean water test was 159 performed in order to compare the permeate flux before and after filtration of the digestate liquid 160 fraction. Then the membranes were chemically cleaned. The cleaning procedure was a sequence of a 161 basic cleaning using a 0.1 M NaOH solution (pH 12.1 ± 0.1), followed by an acidic cleaning using 0. were measured as a function of pH. This is crucial for understanding the acid-base properties of 227 membrane surface functional groups. Measuring the observed zeta potential of the feed as well as 228 for the obtained concentrates and permeates (Table 3) would help to elucidate the interaction 229 between the foulant and the microfiltration membrane surfaces. 230 Table 3 presents the observed zeta potential of the feed, concentrate and permeate samples 231 obtained during concentration experiments using PS5 and PVDF5. The typical pH of biogas plant 232 digestates is 8 to 8.5. All samples were measured at a pH of 8.2 ± 0.1, to simulate the typical feed 233 pH. All the tested samples had a zeta potential between -18.2 and -14.7 mV. These negative zeta 234 potential values at high pH are most likely related to the amino and carboxyl groups present in the 235 proteins of the analyzed samples and only to a smaller degree to the charge of the solids [33]. This 236 might explain why feed, concentrate and permeate show similar zeta potentials, as proteins are not 237 expected to be separated to a significant degree by microfiltration. 238 Figure 3 presents the zeta potential measurements for PVDF and PS membranes as a function of the 239 solution pH. For the studied pH range, the PS membranes showed a significantly less negative zeta 240 potential than the surface modified PVDF membranes. In the pH range of interest for the membrane 241 feed (pH 8-8.5), both PS and PVDF membrane materials showed stable values of observed zeta 242 potentials being -18.2mV and -16.0 for and PVDF2 and PVDF5, respectively; and -6.9mV and -7.1mV 243 for PS2 and PS5, respectively. The initial properties of the membranes suggest that foulants might 244 get more attracted to the PS membrane surface rather than to the PVDF membrane as the feed has 245 an observed zeta potential of about -18mV at a basic pH (Table 3). As the difference in charge is 246 larger between foulants and the PS than for the PVDF membrane, the PS membrane is thus less 247 repulsive initially. 248

Microfiltration of digestate liquid fraction during full recycling experiments 249
Water test on microfiltration membranes 250 Water tests were conducted on fresh membranes prior to each filtration experiment (Table 4) The relative permeate fluxes (J/J 0 ) to compare the effect of membrane material and pore diameter 279 during microfiltration using the two PS and PVDF membranes were plotted against each other in 280 scatter plots ( Figure 5 -Figure 7). These plots mostly give information about the effect of operational 281 conditions on the relative permeate flux pattern. The reason why the relative permeate flux for PS is 282 higher than the ones for PVDF is basically due to the higher clean water flux values obtained for 283 PVDF membranes. 284 The impact of the operational conditions, i.e. applied pressure and cross flow velocity, on the 285 relative permeate fluxes for the same pore size, for the two PS and PVDF membranes can be 286 observed in the individual graphs in Figure 5 a), b) and c). The relative permeate flux achieved with 287 the PS membranes was generally higher than that obtained when using the PVDF membranes for a 288 similar filtration time (i.e. the time track lines lay beneath the identity line). However, it can be 289 observed that the relative permeate flux decay with time was more pronounced for PS membranes 290 than for PVDF membranes for all the membrane pore sizes. The application of a higher pressure also 291 affected the relative permeate flux of the PS membranes more negatively than the PVDF 292 membranes, and enhanced the reduction of the relative permeate flux most probably due to an 293 increasing compression of the cake layer, especially for the smaller membrane pore sizes (i.e. 0.2 µm 294 and 0.5 µm). The combination of low pressures and high velocities appeared to be beneficial for all 295 membranes. It also seems that the larger the membrane pore size, the smaller the effect of the 296 material interaction, which causes an ultimate similar permeate relative fluxes for the two PS and 297 PVDF membranes (Figure 4). 298 For the PVDF membranes (Figure 6), the relative fluxes were higher when using 0.5 and 0.8 µm 299 membrane pore diameters (Figure 6 b), c)) than for the 0.2 µm pore diameter membranes (Figure 6  300 a)). The increase in cross flow velocity led to a higher improvement in relative flux for the PVDF8 301 membrane than for the PVDF5 membrane (Figure 6 c)). Further, this trend is more pronounced at 302 lower than at higher pressure. This situation has two implications as i) the increase in shear force, 303 related to higher velocities, enhanced the removal of the loosely deposited particles on PVDF 304 membranes leading to an increase in the critical flux [22] and ii) the increase in pressure more 305 negatively affected the relative permeate flux of the membranes with the larger pore diameters. 306 This suggests that an increase in pressure causes more membrane blocking in the larger membrane 307 pores of the PVDF membranes [35]. 308 Figure 7 shows the effect of PS membrane pore sizes on the relative permeate flux. It can be 309 observed that the effect of material interaction was more pronounced for smaller pore sizes in PS 310 membranes (Figure 7 a) and b)) especially for the experiments run at a lower pressure. An increase 311 in pressure negatively affected the relative permeate flux on PS membranes. This most probably is 312 due to an increasing compression in the cake layer, especially for smaller membrane pore sizes (0.2 313 µm and 0.5 µm). In contrast, an increase in the cross flow velocity affected positively the relative 314 permeate flux of the PS membranes. However, this improvement was limited as compared to PVDF 315 membranes, as foulants were strongly adsorbed on the PS membrane surfaces. The relative 316 permeate fluxes achieved with the PS2 membranes were higher compared to the PS5 and PS8. 317 However, the relative permeate flux values for both PS5 and PS8 were very alike (Figure 7 c). This 318 suggests that for smaller pore sizes, the fouling mechanism related to PS membranes might be 319 predominantly governed by the presence of an adsorbed layer that might enhance pore blocking 320 and cake filtration for larger pore sizes. 321

Summary of microfiltration experiments 322
The PS membranes achieved a lower permeate flux than the PVDF membranes during approximately 323 the first hour of the microfiltration experiment (shown in Figure 4), which suggests that material 324 interaction initially plays an important role during microfiltration of digestate liquid fractions. The 325 lower initial permeate flux for the PS membranes could be further explained by the fact that the PS 326 membrane surface is more hydrophobic than the surface modified PVDF membrane [18,19]. 327 Additionally, the PS membrane surfaces present a less negative zeta potential (Figure 3) than the 328 foulants (Table 3) and therefore were initially more prone to foul. For the PVDF membranes, 329 membrane and feed zeta potentials ( Figure 3) were comparable (Table 3) (Table 5). The decrease in relative permeate flux for the 350 PS membrane was 31% when the total solids increased from 2.5 to 3.9% (Table 5). For both 351 membranes, the decrease in relative flux can be considered identical within the experimental error. 352 A decrease in critical permeate flux has also been related to an increase in particle concentration 353 [22]. In general, the flux decline during the concentration process was found to be more acute for 354 PVDF membranes than for PS membranes as for the full recycling microfiltration experiments. This 355 was also observed during full recycling filtration (Figure 4) and as explained, it is the initial difference 356 in zeta potential between the feed and the membrane material that determines the permeate flux 357 decay pattern [12], as the initial higher resistance to fouling for the PVDF membrane is negated by a 358 slower cake layer build up which finally equalizes the morphology and composition of the fouling 359 layers on the two PS and PVDF membranes. 360

Ultrafiltration of digestate liquid fraction during concentration experiments 361
Ultrafiltration can be used as an alternative to microfiltration for concentration of particulate 362 phosphorous. As an alternative PES ultrafiltration membranes were used during the concentration 363 experiment carried out on the Bioscan A/S (Denmark) plant for determining their capacity to reject 364 phosphorus. Figure 9 shows a clear relationship between the increase in the volume reduction factor 365 (VRF) together with the increase in total solids in the concentrate and the decline in permeate flux. 366 As the total solids concentration increased, the flux decay became more acute being reduced by 367 more than 50% at a VRF 0.7 and 3.3 bar of applied pressure. During concentration, permeate flux 368 decline is thus a consequence of both fouling layer build up and concentration polarization [19]. 369 Moreover, increases in osmotic pressure could also affect the driving force for the permeate flux as 370 the retentate concentration increases. 371 Figure 10 shows the phosphorus rejection achieved with PVDF, PS and PES membranes. The 373 phosphorus rejection was found to be higher for the PS5 microfiltration membrane (82% w/w) than 374 for the PVDF5 membrane (60% w/w) for the same VRF. By extrapolation, the PES membrane and the 375 PVDF5 membrane seemed to achieve a similar phosphorus rejection. This might be another indicator 376 that the membrane material did have a significant influence on the phosphorus rejection. Most likely 377 a tighter cake layer builts up on the PS membrane than on the PVDF membrane. This could lead to a 378 higher retention of fine particles by the PS than the PVDF. A counter argument to this explanation 379 would be that the PES and the PVDF show similar rejection, but as the ultrafiltration system used in 380 this study was constructed specifically to reduce the particulate matter, in principle to about 100%, 381 the particle size distribution in the ultrafiltration feed setup was significantly different from that for 382 the microfiltration membranes at similar VRF. It can be concluded that a similar retention is achieved 383 for the ultrafiltration setup and the PVDF microfiltration membrane although making a direct 384 comparison of the two systems is difficult. However, as a similar or better phosphorous retention is 385 possible using microfiltration compared to ultrafiltration, microfiltration might be a cheaper 386 alternative option. 387

Chemical analysis of the obtained fractions during concentration experiments 388
While the permeate flux decline during microfiltration and ultrafiltration concentration of digestate 389 liquid fractions is severe, it is by no means detrimental. However, if the concentration of digestate 390 liquid fractions is to be of practical interest, a substantial concentration of phosphorus should also 391 be achieved. Table 5  form or attached to particles small enough to pass through the fouling cake layer and membrane 399 pores. As most of the nitrogen and potassium was present in dissolved form, the distribution of 400 nitrogen and potassium was not affected to the same degree by the increase in total solids in the 401 concentrate and thus the increase in VRF. 402

Cleaning of fouled microfiltration membranes 403
Membrane cleaning was applied to recover the membrane initial permeate flux after microfiltration 404 of digestate liquid fractions. A simple two-step basic and acidic cleaning procedure was tested. The 405 basic cleaning (0.1M NaOH, pH 12.1 ± 0.1) was intended to remove organic matter while the acidic 406 cleaning (0.1M citric acid, pH 2.2 ± 0.1) would remove inorganic foulants. The cleaning procedures 407 were performed after 3-4 hours of microfiltration of the digestate liquid fraction in full recycling 408 mode. Table 6 shows the clean water flux recovery achieved after chemical cleaning compared to 409 the clean water flux measured after the fouling cycle. 410 As can be observed in Table 6, the clean water flux recovery is insufficient for both fouled PVDF and 411 PS membranes. The water flux recovery after a two-step basic-acidic chemical cleaning was 42% for 412 PS membranes and 20% for PVDF membranes. However the clean water flux recovery after basic 413 cleaning was 19% for PS and 33% for PVDF. This indicates that the acidic cleaning might not be a 414 good strategy for PVDF membranes. Additionally, after flushing the membranes with water, PS 415 membranes achieved 75% of water flux recovery while PVDF membranes reached 84%, compared to 416 the respective final permeate flux value obtained at the end of the filtration cycle. Beier et al found 417 [19] that after ultrafiltration of a model bovine serum albumin (BSA) solution, the initial water flux 418 for PVDF membranes could be easily restored by flushing the membrane surface with water, while 419 PS membranes required chemical cleaning. The reason for a better flux recovery achieved for PVDF 420 membranes in this study could be related to the stronger attachment of the fouling layer to the PS 421 membrane. This made necessary the chemical cleaning of the PS membrane to restore the initial 422 water flux. Further investigations on PS and PVDF membranes fouled with digestate liquid fractions 423 are necessary as the two membranes would need to undergo different cleaning strategies, rather 424 than the ones performed in this paper, to achieve higher flux recoveries as obtained elsewhere [12]. 425

426
Two types of microfiltration membranes (PS and surface modified PVDF), used to process digestate 427 liquid fractions, were compared to an ultrafiltration membrane (PES) in terms of recovery of 428 particulate phosphorus from biogas plant digestates. It was shown that the membrane material 429 influences both the fouling mechanism and the rate of phosphorus recovery. When comparing the 430 two microfiltration materials used in this study, it was found that the PS membranes form a fouling 431 cake layer faster than PVDF membranes but as fouling starts to build up, the permeate flux of the 432 two membranes becomes similar. The PS membranes presented a higher initial fouling than the 433 PVDF membranes possibly due to a less negative zeta potential, which tended to initially attract 434 foulants more. On one hand, at lower pore sizes for PS membranes, the formation of an adsorbed 435 layer seems to play the main role during fouling while adsorption combined with pore blocking and a 436 cake layer formation seems to rule the fouling mechanism at larger PS membrane pore sizes. On the 437 other hand, PVDF membranes reached higher permeate fluxes for larger pore sizes. Moreover, 438 increases in the applied pressure affected more negatively PS membranes than PVDF membranes in 439 terms of the relative permeate flux, most probably due to cake layer compression, especially for 440 smaller membrane pore sizes. Although membrane pore blocking could be enhanced with an 441 increasing filtration pressure for larger pore sizes of PVDF membranes, increases in cross flow 442 velocity affected more positively the permeate flux on the PVDF membranes than the PS 443 membranes, as foulants were adsorbed stronger on the PS membrane surface. In general, larger 444 membrane pore sizes (i.e. 0.5 µm and 0.8 µm pore size) for the two PS and PVDF membranes 445 reduced the material effect on the fouling mechanism. Additionally, it was shown that when using PS 446 and PES membranes a phosphorus recovery of about 80% was achieved while 60% phosphorus 447 recovery was obtained when using PVDF membranes.