2.1. Materials
The Nannochloropsis sp. biomass used in this work was cultivated in photobioreactors (PBRs), mechanically disrupted and supplied by A4F—Algae for Future (Portugal). For the ultrafiltration studies, the Nannochloropsis sp. disrupted biomass (100 g·L−1, on a salt-free dry weight basis—SFDW) was centrifuged at 11,000× g (for 2 × 20 min, 5 °C) and the collected supernatant (30 gSFDW·L−1) was analyzed in terms of its salt-free dry weight, proteins, lipids, and chlorophyll-a content.
Previous studies of ultrafiltration using a hydrophilic ceramic membrane with a nominal molecular weight cut-off (MWCO) of 300 kDa revealed a significant volumetric flux decrease during the process, which may be related to high fouling, namely intrapore fouling (data not shown here).
Nannochloropsis sp. contains various proteins, which are associated with light-harvesting complexes, with molecular masses ranging between 21 and 32 kDa [
31]. Therefore, for optimization of soluble protein recovery, a hollow fiber module containing a tighter ultrafiltration membrane with an MWCO of 100 kDa was used (GE Healthcare, Chicago, IL, USA; model: UFP-100-C-5A). It was composed of 520 hydrophilic polysulfone (PS) fibers, with an internal diameter of 0.5 mm, length of 0.32 m, and total filtration area of 0.2 m
2. The membrane selected is hydrophilic, to enhance the permeation of soluble proteins, as the extract is aqueous, and to avoid protein denaturation that is prone to occurring at the surface of hydrophobic materials. The membrane-cleaning products Ultrasil 110 and Ultrasil 75 were purchased from Ecolab (Mississauga, Ontario, Canada) and ethanol (97%) was obtained from Panreac (Barcelona, Spain). For analytical procedures, acetone 90.0% and methanol 99.8% were purchased from JMGS (Odivelas, Portugal), chloroform was obtained from Honeywell/Riedel-de Haën (Seelze, Germany), and the Pierce™ BCA Protein Assay Kit was acquired from Thermo Fisher Scientific (Waltham, MA, USA).
2.2. Membrane Processing
In this work, the supernatant (30 g
SFDW·L
−1) from the centrifugation of
Nannochloropsis sp. disrupted biomass and its dilution, with a dilution of 1:3 corresponding to 10 g
SFDW·L
−1, were processed with the selected membrane system, as summarized in
Table 1. The supernatant (30 g
SFDW·L
−1 and 60 g
DW·L
−1) from the centrifugation of
Nannochloropsis sp. disrupted biomass was diluted with
Nannochloropsis sp. culture medium, which has a salt concentration of 30 g
DW·L
−1, to a dilution ratio of 1:3. After the dilution, the concentration of the diluted supernatant was 10 g
SFDW·L
−1 and 40 g
DW·L
−1.
Primarily, the membrane ultrafiltration experiments were conducted in a concentration operation mode, under controlled transmembrane pressure (TMP) conditions, as it is the most commonly reported in the literature (see the scheme of membrane setup in
Figure 1a). The membrane filtration unit constituted a feed tank (1), recirculation pump (2), pressure gauges (3), membrane module (4), and permeate tank (5). Additionally, the ultrafiltration experiments were also performed in a diafiltration operation mode followed by concentration, under controlled transmembrane pressure conditions (see the scheme of membrane set-up in
Figure 1b), where a water tank (6) and a peristaltic pump, for a water feed, were added. The value of transmembrane pressure was set as 0.2 bar (see
Table 1) to preserve mild operating conditions.
As shown in
Table 1, three experiments were also performed under controlled permeate flux conditions, where the permeate flux was fixed and the TMP was allowed to vary. The membrane setup for the controlled permeate flux experiments was similar to the preceding one (setup for the controlled TMP experiments), both for concentration and diafiltration operation modes (see the scheme of membrane setup in
Figure 1c,d, respectively). When operating under controlled permeate flux, its value was imposed by introducing a positive displacement pump (5) in the permeate circuit, which controls and defines the permeate flux, as presented in
Figure 1c,d. The value of permeate flux was set at 12 L·m
−2·h
−1 (see
Table 1). This value was selected after preliminary experiments, which showed that at this flux, no increase of transmembrane pressure occurs (sub-critical, sustainable flux conditions).
In each membrane experiment, the feed was pumped from the feed tank through the membrane module using a peristaltic pump and recirculated to the feed tank, with a cross-flow velocity of 0.25 m·s
−1. The feed reservoir was filled with 1.5 ± 0.5 L of protein aqueous extract at 20 °C and pH 6.7 ± 0.1. The permeate was collected in the permeate tank and the permeate flux was monitored by mass acquisition using an electronic balance (Kern 572, Kern, Balingen, Germany). For the ultrafiltration in a diafiltration operation mode, the diafiltration volume
D (-) was calculated through Equation (1):
where
Vwater added (L) is the volume of the water (diafiltration solvent) added and
Vfeed (L) is the volume of the feed. For the ultrafiltration process operated in a concentration operation mode, the concentration factor,
CF (-), was determined according to Equation (2):
where
mt0 (kg) represents the mass in the feed compartment in the beginning of the experiment and
mpermeate (kg) is the permeate mass collected during a given period.
The permeate volumetric flux, permeance, apparent rejection of soluble proteins, and percentage of soluble protein recovered in the permeate were calculated during the membrane ultrafiltration experiments.
The volumetric permeate flux
Jv (L·m
−2·h
−1) was calculated according to Equation (3):
where
mpermeate (kg) is the mass of permeate,
ρ (kg·m
−3) is the density of the permeate,
A (m
2) is the total membrane filtration area, and
t (h) is the time of permeation. The permeance
Lp (L·m
−2·h
−1·bar
−1) was calculated through Equation (4):
where
Jv (L·m
−2·h
−1) is the volumetric permeate flux and
TMP (bar) is the transmembrane pressure. In each test (controlled transmembrane pressure concentration/diafiltration and controlled permeate flux concentration/diafiltration), the global apparent rejection of a target compound
Ri (%) and the percentage of soluble protein recovered in the permeate were calculated during the filtration experiments through Equation (5) and Equation (6), respectively:
where in Equation (5),
Ci,perm (g·L
−1) is the concentration of the target compound in the collected permeate and
Ci,feed (g·L
−1) is the concentration of the same compound in the feed (retentate) compartment, and in Equation (6),
ms.protein,perm (
t) (g) is the mass of soluble protein in the permeate compartment at a defined time of the experiment (
t) and
ms.protein,feed (
t0) (g) is the mass of soluble protein in the feed compartment at the initial time of the experiment.
The soluble protein loss was defined as the soluble protein that is not recovered in the permeate for each experiment and was evaluated through a soluble protein mass balance, as shown in Equation (7):
where
ms.protein,feed (
t0) (g) is the mass of soluble protein in the feed compartment at the initial time of the experiment,
ms.protein,perm (
tF) (g) is the mass of soluble protein recovered in the permeate at the end of the experiment,
ms.protein,reten (
tF) (g) is the mass of soluble protein in the retentate at the end of the experiment, and
ms.protein,accum (
tF) (g) is the soluble protein adsorbed/accumulated on the membrane surface and/or pores (surface intra-pore fouling).
The fouling phenomena, defined as the undesirable adsorption/deposition of dissolved solutes or suspended particles, was evaluated through the resistance-in-series model [
32]:
where in Equation (8),
Jv (m·s
−1) is the volumetric permeate flux,
TMP (Pa) the transmembrane pressure, ղ (Pa·s) the dynamic viscosity of the permeate (we considered the water viscosity at 20 °C, 1.002 × 10
−3 Pa·s, since we are dealing with an aqueous extract), and
Rtotal (m
−1) represents the total resistance. We considered that the total resistance
Rtotal (m
−1) results from a series of resistances introduced by the intrinsic membrane resistance
Rm (m
−1) and the resistance caused by the fouling, which may comprise two distinct contributions:
Rrev (m
−1), standing for the resistance that disappears when the pressure across the membrane is released and pure water is passed through the system at the end of the experiment, removing unbounded solutes from the membrane surface—i.e., reversible fouling resistance;
Rirrev (m
−1), the resistance that is only removed when a cleaning cycle using chemical agents is applied, removing the compounds chemically bonded to the membrane—i.e., irreversible fouling resistance. Therefore, we have then for the total resistance:
The resistance that the membrane offers to the permeation of pure water,
Rm (m
−1), was calculated through the hydraulic permeance measured at the beginning of each experiment:
where
Lpw (m·Pa
−1·s
−1) is the hydraulic permeance and ղ
w (Pa·s) is the viscosity of water at 20 °C (1.002 × 10
−3 Pa·s).
To estimate the remaining resistances, as mentioned above (9), the total resistance, Rtotal (m−1) was calculated at the end of each Nannochloropsis sp. supernatant filtration experiment, using the last measured value of volumetric permeate flux, Jv (m·s−1), and Rm + Rirrev (m−1) values were obtained through the volumetric flux of water, measured after a flush with pure water (20 °C) to eliminate reversible fouling. By subtracting Rm + Rirrev (m−1) from Rtotal (m−1), it was then possible to calculate Rrev (m−1) and compute Rirrev (m−1), subtracting the known parameters, Rm (m−1) and Rrev (m−1), from the total resistance Rtotal (m−1).