**Removal of Di**ff**erent Dye Solutions: A Comparison Study Using a Polyamide NF Membrane**

**Asunción María Hidalgo 1,\* , Gerardo León 2 , María Gómez <sup>1</sup> , María Dolores Murcia <sup>1</sup> , Elisa Gómez <sup>1</sup> and José Antonio Macario <sup>1</sup>**


Received: 9 November 2020; Accepted: 9 December 2020; Published: 10 December 2020

**Abstract:** The removal of organic dyes in aquatic media is, nowadays, a very pressing environmental problem. These dyes usually come from industries, such as textiles, food, and pharmaceuticals, among others, and their harm is produced by preventing the penetration of solar radiation in the aquatic medium, which leads to a great reduction in the process of photosynthesis, therefore damaging the aquatic ecosystems. The feasibility of implementing a process of nanofiltration in the purification treatment of an aqueous stream with small size dyes has been studied. Six dyes were chosen: Acid Brown-83, Allura Red, Basic Fuchsin, Crystal Violet, Methyl Orange and Sunset Yellow, with similar molecular volume (from 250 to 380 Å). The nanofiltration membrane NF99 was selected. Five of these molecules with different sizes, shapes and charges were employed in order to study the behavior of the membrane for two system characteristic parameters: permeate flux and rejection coefficient. Furthermore, a microscopy study and a behavior analysis of the membrane were carried out after using the largest molecule. Finally, the Spiegler–Kedem–Katchalsky model was applied to simulate the behavior of the membrane on the elimination of this group of dyes.

**Keywords:** characterization; dyes; molecular structure; nanofiltration; physico-chemical properties

### **1. Introduction**

Organic dyes (such as Acid Brown-83, Allura Red, Basic Fuchsin, Crystal Violet, Methyl Orange and Sunset Yellow) can be found in effluents of different industries (food, medical, painting) but the most pollutant industry is the textile.

The discharge of these pollutants into the aquatic environment has a strong environmental impact due to the amount of toxic compounds they have and also due to the fact that they cause a decrease in the self-purification capacity of the water they are discharged into. This phenomenon prevents plants from performing photosynthesis and microorganisms from developing their biological activity [1].

Therefore, there are numerous methods of disposal of dye aqueous solutions, which can be grouped into physical, chemical and biological methods, but none of them stand out among the others [2–5]. Following this, new techniques are being investigated, including membrane technologies, because they offer low costs and give good yields [1,6].

As a result, membrane technology is attracting great interest. This technology is based on the separation of compounds by size and charge, as the membrane acts as a filter that retains the molecules which are larger than the pore and allows the water to pass through. In the last decade, more than 65% of research works have been based on the fabrication strategies of nanoporous membranes and their

applications in the field of water purification [7–12]. According to Wang et al. the solute and water permeability play important roles in the membrane performance. The membrane is able to separate pollutants from water mainly through size exclusion and solute diffusion [7].

The application of pressure driven membrane processes for the removal of low molecular weight organic compounds from aqueous solutions has been described in several recent publications for example, phenol and chlorophenol compounds [13]. A comparative study, using different organic compounds (atrazine, aniline and phenol, and their derivatives 4-chlorophenol, 4-nitrophenol and 4-nitroaniline) in aqueous solution and their elimination through NF-97 polyamide membrane, was carried out. The different physicochemical parameters of the organic compounds, the permeate flux and the rejection coefficient values were found to be correlated. The best correlation for the rejection coefficient was obtained using the molar refractivity and the water solubility of the compound simultaneously. For permeate flux, the best correlation value was obtained using the surface tension and molecular weight [14].

It is clear that removal efficiency depends on the membrane type and solute, and the interaction between them. Temperature, pH, pressure and concentration also influence rejection [15]. Whether nanofiltration should be used in the treatment of wastewater containing dyes depends on the rejection capacity of the membranes and the permeate flux.

In addition, distilled water tests were performed in order to characterize the membrane, and selectivity tests facing salt solutions before and after dyes tests were carried out in order to know the membrane permeability, studying performance and its changes during the process. In that way, membrane fouling can be analyzed, as well as the phenomenon in which membrane pores get wider because substances passing, known as swelling, can be observed [16–18].

The discussion on membrane-based treatment processes is incomplete without an elaborate perception of the mechanism governing the transport of solute across the membrane and compressive modeling of a membrane-based technique [1].

The main goal of this research work is to study the behavior of the NF99 membrane on the elimination of several dyes, which are molecules of different structure, charge and shape, the following ones being chosen: Allura Red, Basic Fuchsin, Crystal Violet, Methyl Orange and Sunset Yellow. These molecules were selected since in the bibliography there are no studies for some of them, such as Basic Fuchsin and Allura Red. Solutions of each dye were used to characterize the system and to obtain the values of the permeate flux and rejection coefficient. Furthermore, a preliminary study on the characterization of the membrane treated with salt solutions was carried out before and after the dye treatment. Such a study was complemented by scanning electric microscopy (SEM) morphologic study of the membrane using the Acid Brown-83 dye. This molecule was selected because it is a real case of a leather tanning industry located in Murcia (Spain). Finally, the Spiegler–Kedem–Katchalsky model was applied to simulate the behavior of the membrane on the elimination of this group of dyes.

### **2. Materials and Methods**

### *2.1. Materials*

### 2.1.1. Membrane

A nanofiltration membrane was employed in this research. Its main technical characteristics are shown in Table 1.


**Table 1.** Characteristics of the membrane used in the experimental test module.

### 2.1.2. Reagents

The following reagents were used in the assays:


In Table 2, Log Kwo, pK<sup>a</sup> and water solubility data, obtained using PubChem, are shown.

**Table 2.** Chemical properties of some of the dyes employed in the study. Data obtained using PubChem. https://pubchem.ncbi.nlm.nih.gov/.


(\*) https://www.carlroth.com/medias/.

### *2.2. Equipment*

The research was carried out in a membrane module from INDEVEN CF (Spain), which has been designed at laboratory scale to obtain further information on the behavior of plane membranes with small surface area. In addition to the membrane module, other equipment was used to obtain valuable parameters for further comparison and discussion among the different dyes.

### 2.2.1. Membrane Module

The membrane module consists of three main stages of installation: feed tank, fluid impulsion pump and membrane settlement. Furthermore, there is a manometer and a rotameter that measure rejection pressure.

The feed tank is cylindric and it maintains the internal fluid at a constant temperature. Its capacity is of 12 L. The fluid passes from the feed tank to the driving pump through a flexible rubber pipe. The pump is a triple plunger pump from Flowmax (Spain). It consists of three AISI 316 Steel valves and of corrosion resistant double collectors. Flow rate is controlled by a manual needle valve.

The membrane inflows are divided into two: permeate flow and concentrate flow. The last one re-enters the feed tank. Moreover, the vent plug discharge and the caudal control are carried out by a flow that leaves the impulsion pump and arrives to the feed tank.

The membrane, whose surface is 30 cm<sup>2</sup> , is placed near the feeding spacer, with the active layer looking towards the mainboard. The following step involves placing the permeate spacer and finally the closing plate. Two o-rings seal the set.

Continuous functioning is guaranteed because the concentrate flow discharges in the feed tank. Operating pressure is regulated by a valve and a manometer, and the flow is measured by a rotameter from TechFluid, which detects flows ranging from 50 to 400 L/h.

### 2.2.2. Spectrophotometer

A spectrophotometer from Shimadzu (UV–160) (Japan) was employed to measure the absorbance of the dye samples. The measurements were carried out at specific wavelengths, which were 443 nm for Acid Brown-83, 485 nm for Sunset Yellow, 596 nm for Crystal Violet, 500 nm for Allura Red, 460 nm for Methyl Orange and 550 nm for Basic Fuchsin.

### 2.2.3. Variable Pressure Scanning Electron Microscope

To develop the membrane fouling study, a scanning electron microscope from Hitachi (S-3500N model) (Japan) was employed; its main characteristics are the following:


### *2.3. Experimental Series*

In order to obtain further knowledge about the membrane behavior, a series of experiments with different dyes were carried out. In these series of experiments, all the operating conditions remained unchanged excluding that which was subject of study.

### 2.3.1. Distilled Water Assays

The tank was filled with distilled water and afterwards a series of experiments were carried out. The experiments were of 15 min of length, at 5, 10 and 15 bar operating pressures and with a 150 L/h flow. The main goal of these assays was to get to know the permeability of the membrane.

### 2.3.2. Salts Assays

The objective of the experimental assays using salt solutions was to obtain the membrane rejection coefficient and, as a result, its selectivity. In the same way that distilled water assays, the experiments were carried out at 5, 10 and 15 bar operating pressure and with a 150 L/h flow; however, the duration time was of 20 min. Aqueous solutions of 1 g/L were used to carry out the experimental assays. The salts employed in the experiments were magnesium chloride and sodium chloride.

### 2.3.3. Dyes Assays

In order to elucidate the membrane elimination power in detail, a 50 mg/L dye dissolution was employed to fill the feed tank and 30 min assays were carried out, in which the samples were taken every 5 min with different operating pressures. Duplicate assays were carried out.


### **3. Results and Discussion**

### *3.1. Membrane Characterization*

The initial membrane characterization was carried out by determining its permeability coefficient, its performance regarding flows and its selectivity against two different salt solutions: sodium chloride and magnesium chloride.

### Permeability Coefficient Determination

In order to determine the permeability coefficient, the following equation was used:

$$I\_{\upsilon} = L\_p \cdot (\Delta P - \Delta II) \tag{1}$$

The osmotic pressure gradient can be ignored only if the solvent is employed alone. As a result, the previous equation can be described as:

$$J\_{\upsilon} = L\_{p} \cdot \Delta P \tag{2}$$

The permeability coefficient value is obtained by representing the final values of the solvent mass flow against applied pressure.

In Table 3, the permeability coefficient for the solvent (*Lp*) values obtained in this research for different pressure ranges and the permeability coefficient values found in the literature are shown. As can be observed, the values obtained in this research are of the same order as those found in the literature [19–21].


**Table 3.** Properties of the membrane used in the experimental assays.

### *3.2. Determination of Selectivity and Performance of the Membrane against Salt Solutions*

The characterization of NF membranes is often carried out using divalent salt solutions. In this research, two salt solutions were used: sodium chloride and magnesium chloride.

To determinate the selectivity of the membrane, the rejection coefficient was established:

$$\%R = \frac{\left(\mathbb{C}\_0 - \mathbb{C}\_p\right)}{\mathbb{C}\_0} \cdot 100\tag{3}$$

The experimental values obtained for permeate flux and rejection coefficient were treated by applying the solution-diffusion model [22]. As a result, the permeability coefficient for the solute (Ps) for each salt solution was obtained.

Table 2 shows the Ps values for each salt solution assayed, which are very close to those obtained by previous authors [13].

### *3.3. Influence of the Chemical Structure of Di*ff*erent Dyes*

Usually, parameters such as molecular weight, log Kw and pKa, were used to explain the membrane selectivity and the rejection coefficient. However, in recent years, to attempt to explain the behavior of the nanofiltration systems, based on the two characteristic parameters, the permeate flux and the rejection coefficient, the influence of chemical structure parameters could represent an important factor to consider.

Table 4 shows the chemical structure parameters of the dye molecules. The parameters, such as area, radio, length and volume were obtained by the program, MarvinSketch version 15.12.7, using ChemAxon. Furthermore, Figure 1 shows the charge, shape and geometry of different molecules using a tridimensional draw. These parameters were proven to influence permeate flux and rejection coefficient.

**Table 4.** Structure parameters of the dye molecules. Data obtained using ChemAxon. https: //chemaxon.com/products/marvin.


**Figure 1.** Molecular properties of the dyes.

In the literature, some authors have described that the most influential parameter is molecular volume, but other parameters related to chemical properties can also be used to predict the behavior of these systems. Figure 2 shows the influence of molecular volume in rejection coefficients and permeate flux using a pressure of 10 bar (a), and 15 bar (b) for the different dyes.

**1.01**

**5.0**

♦ ■ **Figure 2.** Rejection coefficient () and permeate flux () variation with molecular volume for colorants: (MO) Methyl Orange, (BF) Basic Fuchsin, (SY) Sunset Yellow, (AR) Allure Red, (CV) Crystal Violet. Experimental conditions: pH = 7, [Dyes] = 50 mg/L and pressure values (**a**) 10 bar and (**b**) 15 bar.

As can be seen in Figure 2a, when the molecular volume is increased, the selectivity of the membrane increases too. When comparing the values obtained for the different pressures applied, it is seen that high-volume molecules present a small decrease in the rejection coefficient when the pressure applied is high (15 bar). Besides that, the permeate flux has no predictive behavior because both the small dye molecules (MO and BF) that have a similar size to the molecular weight cut off MWCO of the membrane and the large dye molecules (CV) have low permeate fluxes.

According to Cheng et al. (2016), membrane water permeability and solute rejection can be attributed to sensitive pore size and membrane charge. This separation discerned three mechanisms, size exclusion (sieving), electrostatic repulsion (Donnan exclusion) and adsorption. The rejection of neutral molecules and large dye molecules (CV) was mostly size exclusion. The rejection of the low-charged solutes was dominated by the electrostatic interactions, including repulsion (cations) and attraction (anions) (BF) [15].

Furthermore, different parameters of the structure of the molecules were correlated with the permeate fluxes and rejection coefficients obtained (Figures S1–S4, Supplementary Material), and it was tested that the parameter that presents a greater incidence is length perpendicular to the maximum area. Figure 3 shows the influence of perpendicular length to the maximum area on permeate flux and rejection coefficient for (a) 10 bar and (b) 15 bar pressures. In this case, the highest value of length perpendicular to the maximum area corresponds to Sunset Yellow dye, and the values obtained for permeate flux and rejection coefficient show a lineal correlation with this parameter, in this particular range studied.

♦ ■ **Figure 3.** Rejection coefficient () and permeate flux () variation with length perpendicular to the maximum area for colorants: (MO) Methyl Orange, (BF) Basic Fuchsin, (SY) Sunset Yellow, (AR) Allure Red, (CV) Crystal Violet. Experimental conditions: pH = 7, [Dyes] = 50 mg/L and pressure values (**a**) 10 bar and (**b**) 15 bar.

### *3.4. Influence of pH: Comparison of Electrostatic Interaction and Membrane Performance*

The permeate fluxes and rejection coefficients obtained from the dye molecule assays were studied with different pH values using the NF99 membrane under identical conditions and the obtained values are shown in Table 5.

**Table 5.** Rejection coefficient and permeate flux variation with pH of feed for the different dyes. P = 15 bar; [Dye] = 50 mg/L.


According to different authors [11,15,16], electrostatic interactions between the membrane and charged molecules is an important parameter which determines flux decline. Usually, the pH values of effluents from the dyeing industry are between neutral and basic pH [4]. In this pH range, the polyamide NF99 membrane possesses negative charge, and therefore the negatively charged dye molecules (AB83, AR, MO and SY) are not electrostatically attracted towards the membrane, and hence they do not significantly reduce the permeate flux. However, AB83 and MO dyes molecules showed a decrease in permeate flux, being more significant in the case of MO. This behavior could be explained due to other types of interactions, such as hydrophobic ones (between the aromatic rings of both the dyes and the polyamide membrane selective layer) or hydrogen bonds that can play an important role in membrane blocking, especially under the conditions in which the acidic or basic groups in dyes are partially dissociated.

Positively charged dye molecules (CV) of relatively low molecular weight exhibit a strong fouling effect in the neutral as well as the alkaline pH of the feed solution. This behavior is according to the results obtained by Chindambaram et al. [16].

### *3.5. Fouling Phenomenon after Treatment of Dyes Solutions*

A simple means of evaluating the fouling phenomenon effect on the membrane is to repeat the distilled water assays after the dye assays are carried out. In this research, the fouling factor of the membrane, %*FF*, was calculated in order to quantify the fouling phenomenon by comparing the initial, *Lp*0, and final, *Lpf*, values of the permeability coefficient. The equation is the following:

$$\%FF = \frac{\left(L\_{p0} - L\_{pf}\right)}{L\_{p0}} \cdot 100\tag{4}$$

Table 6 shows the results of the fouling factor of the membrane after the use of the different dyes studied.


**Table 6.** Values of fouling factor of the membrane for the different dye assays.

Considering the molecular weight and molecular volume values obtained for each dye, it was found that the smaller dye molecules (MO and BF), whose sizes were close to the molecular weight cut-off (MWCO) of the membrane (200 Da), presented a higher fouling factor. This fact showed that these dyes were absorbed in the membrane and, consequently, the fluxes were reduced. Some authors also described adsorption phenomena for SY [15]. When comparing two dyes of similar molecular size (MO and BF, or CV and AB83), the dye molecules with negative charges and of a linear size gave a lower fouling factor that those of positive charges and with flat disc shape. These results were already described in other studies [23,24].

### *3.6. Morphologic Study of the Membrane*

Even though there are many available techniques for observing the membrane surface (including the active layer and the sublayer that sustains it), the most employed technique for nanofiltration membrane characterization is Scanning Electron Microscopy (SEM).

In this research, the samples of native membrane and used membrane (after carrying out the assay of the dye of higher molecular weight, AB83) were analyzed.

Figure 4a,b shows an SEM picture (300×) of the membrane Alfa Laval NF before starting the assays and after them. Figure 5 shows the energy-dispersive X-ray spectrum of the membrane (a) before the initial assay and (b) after the pass of Acid Brown-83 solutions through the membrane.

**Figure 4.** SEM taken picture (300x) of the Alfa Laval NF membrane (**a**) before starting the assays and (**b**) after them.

**Figure 5.** EDX analysis of the membrane (**a**) before the initial assay and (**b**) after the pass of Acid Brown-83 solutions through the membrane.

When comparing Figures 4 and 5 for the study of the evolution of the Alfa Laval NF membrane after its use, it can be observed that the SEM picture shows that there is a certain degree of fouling. Membrane fouling is mainly observed on the active layer.

Furthermore, according to the energy-dispersive X-ray spectrum, new elements, such as chlorine, iron and nitrogen appear to be on the membrane surface after the assays. The presence of these elements can be explained because of the pass of Sodium Chloride and the dye solution through the membrane, and because of metallic rests from the installation.

### *3.7. Application of the Spiegler–Kedem–Katchalsky Model*

In the bibliography, some adequate models to explain the behavior of the separation process for a thin-layer membrane have been described [25–27]—for example, the solution-diffusion model. Therefore, other models are based on the use of coefficients that relate the permeate flux and the fouling factor of the membrane, but in recent years, the most-used models are based on phenomenological transport.

Those models correlate driving force and flow linearly:

$$J\_{\mathbf{i}} = L\_{\mathbf{i},\mathbf{j}} \cdot \mathbf{X}\_{\mathbf{j}} \tag{5}$$

where *J<sup>i</sup>* is the flow density of the component, *X<sup>j</sup>* is the driving force and *Li,j* is the proportionality coefficient.

The driving forces that dominate the transference of matter in membrane processes are the gradient of pressure and the gradient of concentration.

−

−

 <sup>൰</sup>

 <sup>൰</sup>

<sup>௩</sup> = ∙ ൬

<sup>௩</sup> = ∙ ൬

The Spiegler–Kedem–Katchalsky model [28,29] expresses the initial equations of the previous model in a differential way; not linearly. As a result, it considers that the densities of flux vary through the thickness of the membrane.

$$J\_{\upsilon} = L\_p \cdot \left(\frac{dP}{d\mathbf{x}} - \sigma \frac{d\Pi}{d\mathbf{x}}\right) \tag{6}$$

$$J\_s = P\_s \cdot \frac{d\mathbb{C}\_s}{d\mathbb{x}} + (1 - \sigma) \cdot \mathbb{C}\_s \cdot l\_v \tag{7}$$

When expressing both equations incrementally:

$$J\_v = L\_p \cdot (\Delta P - \sigma \cdot \Delta I I) \tag{8}$$

$$J\_s = P\_s \cdot \left(\mathbb{C}\_m - \mathbb{C}\_p\right) + (1 - \sigma) \cdot l\_v \cdot \mathbb{C}\_s \tag{9}$$

The Spiegler–Kedem–Katchalsky model was initially developed for reverse osmosis processes; however, it has been proven that it is also applicable in some nanofiltration processes [30,31].

This model assumes that transport coefficients are independent of solute concentration. Nevertheless, these coefficients depend on solute concentration for ionic solutions in nanofiltration membranes. As a result, some authors made some changes in the model to consider this fact [32].

There are two parameters to be determined for the Spiegler–Kedem–Katchalsky model:


The pass of a solute flux through the membrane is caused by two different fluxes: a convective flux, which is caused by the application of a gradient of pressure through the membrane, and a diffusive flux, which is caused by the gradient of concentration in both sides of the membrane. The reflection coefficient is also an indicator of what type of flux prevails: the closer the σ values are to 1, the lower participation has the convective flux [33].

For ideal reverse osmosis membranes, σ values are close to 1 as they present a dense structure and no pores that would enable convective transport.

The observed rejection was calculated using the following expression:

$$\% \mathcal{R}\_{obs} = \frac{(1 - F)}{1 - \sigma \cdot F} \cdot 100 \tag{10}$$

where *F* is a parameter that depends on the reflection coefficient, solvent flux, and solute permeability coefficient [34]:

$$F = e^{\left(1 - \frac{1-a}{P\_s} \cdot I\_v\right)} \tag{11}$$

The transport phenomenon through the membrane is, in fact, a combination of convection, solution, and diffusion. In this case, the transport process can be described as an irreversible thermodynamic phenomenon. The following relations among the parameters of the process: reflection coefficient and solute permeability (σ and *Ps*), solvent flux (*Jv*) and observed rejection (*Robs*) were proposed by Spiegler, Kedem and Katchalsky:

$$L n[X] = 1 - \frac{1 - \sigma}{P\_s} \cdot l\_v \tag{12}$$

$$X = \left(\frac{1}{(1-\sigma)} - \frac{1}{1-R\_{obs}}\right) \frac{(1-\sigma)}{\sigma} \tag{13}$$

The parameters of the model were obtained by employing both Equations (12) and (13) along with (9). When combining Equations (12) and (13), Equation (14) is obtained:

$$\ln L n \left[ \left( \frac{1}{(1-\sigma)} - \frac{1}{1-R\_{\rm obs}} \right) \cdot \frac{(1-\sigma)}{\sigma} \right] = 1 - \frac{1-\sigma}{P\_s} \cdot l\_v \tag{14}$$

The average *Robs* value was calculated from the experimental data of rejection coefficients; thus it is now a known value. From this value, a parameter z ( <sup>1</sup> 1−*Robs* ) was calculated. *σ*

Equations (9) and (14) were employed to determine the rejection coefficient (σ) and the permeability coefficient (*Ps*). It was determined that solute concentration in the feed was the same as the solute concentration in the membrane (*C<sup>m</sup>* ≈ *C*0), as few polarization processes occur. The analyzed solute feeding and permeate concentrations are converted to mol/m<sup>3</sup> by dividing by the molecular weight of the different dyes. ≈ *σ*

When replacing *J<sup>s</sup>* , *Jv*, *C*<sup>0</sup> and *C<sup>p</sup>* in Equation (9), and after isolating *P<sup>s</sup>* , the following value, dependent on σ, is obtained: <sup>௦</sup> = ௦ − <sup>௩</sup> · <sup>௦</sup> · ሺ1−)

$$P\_s = \frac{f\_s - f\_v \cdot \mathbb{C}\_s \cdot (1 - \sigma)}{\mathbb{C}\_0 - \mathbb{C}\_p} \tag{15}$$

This would lead to a *P<sup>s</sup>* = *a* − *b*·(1 − σ) type of equation, so Equation (14) would become the following: ቈ൬ <sup>1</sup> − ൰ ∙ ሺ1−) −1+ 1−

$$L\text{tr}\left[\left(\frac{1}{(1-\sigma)} - z\right)\frac{(1-\sigma)}{\sigma}\right] - 1 + \frac{1-\sigma}{a - b \cdot (1-\sigma)} \cdot \frac{f}{f\_{\upsilon}} = 0\tag{16}$$

*σ*

where *a* = *Js C*0−*C<sup>p</sup>* and *b* = *Jv*·*C<sup>s</sup> C*0−*C<sup>p</sup>* . = బି

ೞ

In order to solve this equation of one unknown parameter (σ), it is necessary to use a numeric method, since there is no analytical solution. The program Solver from Excel was employed for that purpose. As a result, the parameters σ and *Ps* were obtained for each different case. The results are shown in Table 7. σ

**Table 7.** Solute permeability coefficient and reflection parameter for the different dyes obtained using SKK model.


To verify the model, the values of F and Robs were calculated. The following figures (Figure 6a–f) show the good correlation in most cases between the experimental values of the rejection coefficient and those calculated by the model. Table 5 shows the standard deviation values being the highest lower than 4%.

**Figure 6.** *Cont*.

♦ ■ ♦ ■ ● ◊ **Figure 6.** Correlation between the experimental values of the rejection coefficient and those calculated by the model. (**a**) AB83, (**b**) AR, (**c**) BF, (**d**) CV, (**e**) • MO, (**f**) ♦ SY.

### *3.8. Comparative Study of the Results*

A comparison of the results obtained on permeate flux and rejection coefficient using NF99 for the different dyes molecules was carried out. Table S1 (Supplementary Material) shows the results obtained by other authors using other membranes (native and modified) for the removal of dyes.

### **4. Conclusions**

− − − − The performance of a polyamide nanofiltration membrane on the removal of six different dyes, Acid Brown-83, Allura Red, Basic Fuchsin, Crystal Violet, Methyl Orange and Sunset Yellow, has been studied. Firstly, the membrane characterization was carried out, obtaining a water permeability coefficient value of 1.665 × 10−<sup>8</sup> s m−<sup>1</sup> . The membrane selectivity was also determined, and the solute permeability coefficients were 6.705 × 10−<sup>6</sup> and 1.632 × 10−<sup>7</sup> for NaCl and MgCl2, respectively. It has been proven that the chemical structure of the dyes has an important influence on the permeate fluxes and rejection coefficients obtained, these being the molecular volume and the length perpendicular to the maximum area the most relevant parameters. The pH influence was also studied, these being the membrane negatively charged at neutral and basic pH and therefore being repelled by the dye molecules of negative charge (AB83, AR, MO and SY). However, AB83 and MO dye molecules showed a decrease in permeate flux, which can be explained due to other types of interactions (hydrophobic interactions and the presence of hydrogen bonds that cause membrane blocking). Membrane fouling was determined by calculating a fouling factor, showing that the smaller dye molecules (Methyl Orange and Basic Fuchsin) presented the highest fouling. Additionally, when comparing dyes of similar molecular sizes, those with negative charges and linear size gave lower values of fouling factor. The morphologic study of the membrane by Scanning Electron Microscopy (SEM) and infrared spectrum confirmed the observed degree of fouling. Finally, the Spiegler–Kedem–Katchalsky model that simulates the membrane behavior was successfully applied, with a high degree of agreement between the experimental and calculated rejection coefficients.

**Supplementary Materials:** The following are available online at http://www.mdpi.com/2077-0375/10/12/408/s1, Figure S1: Rejection coefficient () and permeate flux () variation with dreiding energy (A&B) and with MMFF94

Energy (C&D) for colorants: (MO) Methyl Orange, (BF) Basic Fuchsin, (SY) Sunset Yellow, (AR) Allure Red, (CV) Crystal Violet. Experimental conditions: pH = 7, [Dyes] = 50 mg/L and pressure 10 bar (A&C) and 15 bar (B&D). Figure S2: Rejection coefficient () and permeate flux () variation with minimal projection area (A&B) and with maximal projection area (C&D) for colorants: (MO) Methyl Orange, (BF) Basic Fuchsin, (SY) Sunset Yellow, (AR) Allure Red, (CV) Crystal Violet. Experimental conditions: pH = 7, [Dyes] = 50 mg/L and pressure 10 bar (A&C) and 15 bar (B&D). Figure S3: Rejection coefficient () and permeate flux () variation with minimal projection radius (A&B) and with maximal projection radius (C&D) for colorants: (MO) Methyl Orange, (BF) Basic Fuchsin, (SY) Sunset Yellow, (AR) Allure Red, (CV) Crystal Violet. Experimental conditions: pH = 7, [Dyes] = 50 mg/L and pressure 10 bar (A&C) and 15 bar (B&D). Figure S4: Rejection coefficient () and permeate flux () variation with length perpendicular to the minimal area (A&B) and with molecular weight (C&D) for colorants: (MO) Methyl Orange, (BF) Basic Fuchsin, (SY) Sunset Yellow, (AR) Allure Red, (CV) Crystal Violet. Experimental conditions: pH = 7, [Dyes] = 50 mg/L and pressure 10 bar (A&C) and 15 bar (B&D), Table S1: Comparison of dye removal between previous studies and this study in terms of water flux and rejection.

**Author Contributions:** Conceptualization, A.M.H. and G.L.; methodology, A.M.H. and G.L.; investigation, J.A.M.; resources, J.A.M.; writing—original draft preparation, A.M.H., M.G. and M.D.M.; writing—review and editing, A.M.H., G.L., J.A.M., M.G., M.D.M. and E.G.; supervision, A.M.H., G.L. and E.G. All authors have read and agreed to the published version of the manuscript.

**Funding:** This research received no external funding.

**Acknowledgments:** The authors would like to thank SAIT, from Polytechnic University of Cartagena for the SEM made to the membrane.

**Conflicts of Interest:** The authors declare no conflict of interest.

### **Nomenclature**


### **References**


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## *Article* **Dependence of Water-Permeable Chitosan Membranes on Chitosan Molecular Weight and Alkali Treatment**

### **Ryo-ichi Nakayama \*, Koki Katsumata, Yuta Niwa and Norikazu Namiki**

Department of Environmental Chemistry & Chemical Engineering, School of Advanced Engineering, Kogakuin University, 2665-1 Nakano-machi, Hachioji, Tokyo 192-0015, Japan; s316020@ns.kogakuin.ac.jp (K.K.); st13562@g.kogakuin.jp (Y.N.); nnamiki@cc.kogakuin.ac.jp (N.N.)

**\*** Correspondence: bionakayama.ryo@cc.kogakuin.ac.jp; Tel.: +81-42-628-4876; Fax: +81-42-628-4531

Received: 1 September 2020; Accepted: 6 November 2020; Published: 18 November 2020 -

**Abstract:** Chitosan membranes were prepared by the casting method combined with alkali treatment. The molecular weight of chitosan and the alkali treatment influenced the water content and water permeability of the chitosan membranes. The water content increased as the NaOH concentration was increased from 1 to 5 mol/L. The water permeation flux of chitosan membranes with three different molecular weights increased linearly with the operating pressure and was highest for the membrane formed from chitosan with the lowest molecular weight. Membranes with a lower water content had a higher water flux. The membranes blocked 100% of compounds with molecular weights above methyl orange (MW = 327 Da). At 60 ≤ MW ≤ 600, the blocking rate strongly depended on the substance. The results confirmed that the membranes are suitable for compound separation, such as in purification and wastewater treatment.

**Keywords:** chitosan; membrane; water content; water permeability; alkali treatment

### **1. Introduction**

Chitin and chitosan are biopolymers contained in the exoskeletons of crustaceans that have recently attracted attention as reproducible biogenic components [1,2]. They are significant for effective resource utilization, because they can be obtained from shells that are discarded during the processing of crabs and shrimps for food products [3,4].

Chitin is formed from N-acetyl-D-glucosamine that is linked linearly with β-1,4 units, whereas chitosan is formed from D-glucosamine (i.e., the deacetylation product of chitin). Both structures are similar to cellulose [5,6].

Conventional industrial applications of chitosan include as a flocculent [7], adsorbent [8,9], and fiber [10], because it is commercially and continuously available at low cost. Chitosan is also anticipated to be a biocompatible material in functional gels for drug delivery systems [11–13] and as a membrane material for industrial separation tools [14–16]. Membranes offer several advantages over other separation techniques because of their low energy consumption, bulk production at continuous operation, and production of bio-products that are not thermally denatured. Industrial applications of membrane separation are wide-ranging and include fruit-juice condensation [17], artificial dialysis [18], desalination of seawater [19], and wastewater treatment [20–22].

A membrane is characterized by its mechanical strength (i.e., stress–strain relationship) and mass-transfer characteristics. The mechanical strength determines the handling fatigue life of the membrane in a module. The mass-transfer characteristics determine the molecular diffusion rate through the membrane, which is the main rate-limiting step of the separation process. In general, chitosan is dissolved in aqueous acetic acid [23]. To form a water-insoluble chitosan membrane, the acetic acid must be neutralized by basic components such as sodium hydroxide (NaOH). During preparation, the type and concentration of the basic aqueous solution is known to influence the coagulation rate and structure of the chitosan gel. Moreover, the deacetylation degree of chitosan affects the distilled water permeation characteristics of the membrane [24]. The film-forming properties of chitosan are affected by the molecular weight of the chitosan and the alkali treatment at the time of membrane formation. The latter is essential to stabilizing the film formation against dissolution in water.

In this study, the water content, mechanical strength, water permeability, and mass-transfer characteristics of chitosan membranes were regulated by controlling the molecular weight of chitosan and the alkali treatment of the casting solution.

### **2. Materials and Methods**

### *2.1. Materials*

Powders of chitosan with three different molecular weights were purchased from Sigma-Aldrich (St Louis, MO, USA). Table 1 lists the mean molecular weights of the chitosan powders, which were determined from the measured viscosity. The guaranteed viscosity range of chitosan in the database was based on special-grade sodium hydroxide. Acetic acid and other chemicals were purchased from Fujifilm Wako Pure Chemical Industries, Ltd. (Osaka, Japan).


**Table 1.** Measured molecular weights of chitosan powders.

\* Mean molecular weight of chitosan was determined by viscosity measurement. \*\* The range of viscosity was quoted from Sigma-Aldrich.

### *2.2. Preparation of Chitosan Membrane*

Figure 1 shows the procedure for preparing the chitosan membranes. Chitosan was dissolved in 1.7 mol/L of acetic acid solution (20 g/L). The chitosan solution (20 g) was dispensed into a petri dish (diameter of 7.5 cm) and then dried for 12 h at 333 K in a thermostatic chamber. Subsequently, the chitosan was gelled by immersion in NaOH at a sufficient concentration (volume of = 25 mL, NaOH concentration = 1–5 mol/L). The chitosan in the petri dish was continuously immersed in the NaOH solution for 15–360 min. The resulting membrane was washed with distilled water. After the alkali treatment, the swollen membrane easily separated from the glass plate; it was thoroughly washed with distilled water to remove any excess NaOH. The neutralized state was checked by immersing pH paper in the wash water.

**Figure 1.** Experimental procedure for preparing the chitosan membranes.

### *2.3. Scanning Electron Microscopy*

The membranes were snap-frozen in liquid nitrogen and then dried in a vacuum freeze dryer (RLE-103, Kyowa Vacuum Engineering. Co., Ltd., Tokyo, Japan) at 298 K for 24 h. Then, the dried membranes were sputter-coated with a thin Pt membrane, using a sputter-coater (E-1010 Ion Sputter, Hitachi, Ltd., Tokyo, Japan). Finally, cross-sectional images of the membranes were obtained using a scanning electron microscope (SEM) (Miniscope TM-1000, Hitachi, Ltd., Tokyo, Japan).

### *2.4. Water Content*

To determine the internal structure of a swollen membrane, the volumetric water content (*HV*) was determined from the water content of the membrane. For this purpose, each membrane was cut into 4 × 4 cm squares. Because a membrane sequesters water in its void spaces, the volumetric water content reasonably approximates the void fraction of a membrane in the swollen state. Each membrane square was blotted with filter paper to remove the excess surface water and was then dried in a thermo-controlled oven (333 K, 24 h). The water loss was measured gravimetrically with an electronic balance (ER-180A; A&D Co. Ltd., Tokyo). The volumetric water content each membrane square was obtained by calculating its gravimetric change after swelling:

$$H\_V = \frac{V\_w}{V\_m} \tag{1}$$

 where *V<sup>w</sup>* is the volume of water in the membrane, and *V<sup>m</sup>* is the volume of the membrane.

characterize the mechanical strength. The maximum stress σ was calculated as follows:

*λ*

### *2.5. Mechanical Strength*

− The mechanical strengths of the membranes were measured with a rheometer (CR-DX500, Sun Scientific Co., Ltd., Tokyo, Japan). The swollen membranes were cut into 1 × 4 cm samples and stretched at a rate of 1.0 mm s −1 . The maximum stress and maximum strain were measured to characterize the mechanical strength. The maximum stress σ was calculated as follows:

$$
\sigma = \frac{B\_{\text{max}}}{A\_c} \tag{2}
$$

where *Bmax* is the maximum pre-breaking load, and *A<sup>c</sup>* is the cross-sectional area of the initial membrane. The maximum strain λ was calculated as follows:

$$
\lambda = \frac{L\_0 - L\_i}{L\_i} \times 100\tag{3}
$$

where *L<sup>i</sup>* and *L*<sup>0</sup> are the membrane lengths in the initial and breaking states, respectively.

### *2.6. Water Permeability*

The water permeability of the membrane was measured with an ultrafiltration apparatus (UHP-62K, Advantec Tokyo Kaisha, Ltd., Tokyo, Japan) [25]. Figure 2 presents a schematic of the module and the setup for testing the water permeation. The initial volume of the aqueous phase was 200 mL, and the effective membrane surface area was 2.13 × 10 <sup>−</sup><sup>3</sup> m<sup>2</sup> . The operating pressure ∆P (50–200 kPa) was adjusted by introducing N<sup>2</sup> gas at room temperature (298 K). A magnetic stirring bar was installed near the membrane surface and stirred at a constant speed of 3 s −1 in the aqueous phase. The mass of the permeated water was measured on an electric balance and was converted to the volumetric amount of permeated water according to the permeated water density. The volumetric water flux *J<sup>v</sup>* was then calculated as follows: <sup>−</sup> Δ – −

$$J\_V = \frac{V\_p}{A\_m \cdot \ell \cdot t} \tag{4}$$

where *V<sup>p</sup>* is the volumetric amount of permeated water, *A<sup>m</sup>* is the membrane surface area, ℓ is the thickness of the swollen membrane, and *t* is the operating time. These tests were replicated three times. The results of the water permeability test of the membranes were presented with the associated standard deviation (±SD). *ℓ*

**Figure 2.** Schematic of the water permeation apparatus: (**a**) N<sup>2</sup> gas inlet, (**b**) regulator, (**c**) transducer, (**d**) magnetic stirring bar, (**e**) magnetic stirrer, and (**f**) electronic balance.

### *2.7. Measurement of the Mass Transfer Flux*

Ⅼ − A chitosan membrane prepared by the method described in Section 2.2 was installed in the ultrafiltration device, and 190 mL of the sample solution was poured in the permeation cell. After the device was filled with the sample solution, it was sealed, and a vial was attached to the permeation outlet. The experiment was started by pressurizing the device to 100 kPa with N<sup>2</sup> gas. The stirring speed was 190 min −1 . Each experimental sample solution (urea (MW = 60 Da), D-glucose (MW = 180 Da), methyl orange (MW = 327 Da), and bordeaux S (MW = 604 Da)) was dissolved in water as a solvent. After the substance permeation experiment, the absorbance of the sample solution before and after permeation was measured with an extra-visible visible spectrophotometer (V-630IRM, JASCO). After the absorbance was measured, the concentrations *C<sup>f</sup>* and *C<sup>p</sup>* before and after permeation, respectively, were determined from the calibration curve of each sample solution. The apparent rejection rate *R* was then calculated as follows:

$$R = \frac{\left(\mathbb{C}\_f - \mathbb{C}\_p\right)}{\mathbb{C}\_f} \times 100\tag{5}$$

### **3. Results and Discussion**

### *3.1. Scanning Electron Microscopy*

Figure 3 shows scanning electron microscopy (SEM) images of the surfaces and cross-sections of the chitosan membranes prepared in solutions with various NaOH concentrations (CNaOH = 1.0 and 5.0 mol/L) and crosslinking times (*t*<sup>N</sup> = 15 and 180 min). The surfaces of chitosan membranes prepared in 1.0 mol/L NaOH were smooth, and more membrane formed with a longer crosslinking time. Meanwhile, the cross-section showed that the structure became denser with a longer crosslinking time. The chitosan membranes prepared in 5.0 mol/L NaOH developed a rough surface with a longer crosslinking time, and the membrane surface peeled off and deteriorated. Furthermore, the cross-section of the membrane showed voids in the internal structure with a shorter crosslinking time. Previous SEM images demonstrated a measurable change in the biopolymer networks induced by the alkali treatment [26,27].

**Figure 3.** SEM images of the surfaces and cross-sections of chitosan membranes prepared in 1.0 mol/L NaOH (*t*<sup>N</sup> = 15 and 180 min, upper panels) and 5.0 mol/L NaOH (*t*<sup>N</sup> = 15 and 180 min, lower panels).

Figure 4 shows SEM images of the membranes prepared from chitosans with different molecular weights (low, medium, and high) in 5 mol/L NaOH. All of the membranes had uniform and dense internal structures in the thickness direction, but the membrane prepared from chitosan with a high molecular weight developed voids through its cross-section.

**Figure 4.** Cross-sectional SEM images of membranes prepared in 5 mol/L NaOH (*t*<sup>N</sup> = 180 min) from chitosan of different molecular weights: (**a**) low, (**b**) medium, and (**c**) high.

### *3.2. Volumetric Water Content*

Figure 5 shows the effect of the crosslinking time on the water contents of chitosan membranes prepared in 1.0 and 5.0 mol/L NaOH. For the chitosan membranes prepared in 1.0 mol/L NaOH, the water content decreased with increasing crosslinking time. This is probably because the network structure within the membrane densified as the crosslinking progressed. Conversely, the water content of the chitosan membranes prepared in 5.0 mol/L NaOH showed no significant change regardless of the crosslinking time. This was attributed to the rapid progression of the crosslinking in the concentrated NaOH aqueous solution, so the membrane was fully formed within a short time.

**Figure 5.** Effect of crosslinking time on the water content of chitosan membranes prepared in 1.0 and 5.0 mol/L NaOH.

Figure 6 shows the effects of the chitosan molecular weight and NaOH concentration on the volume porosity. For all chitosan membranes, the volumetric water content increased with the NaOH concentration. This trend can be explained by the hydrogen bonds that crosslink the amino and hydroxyl groups of chitosan [28]. Chitosan polymer networks are principally crosslinked by hydrogen bonds between hydrogel groups and amino groups. Increasing the concentration of the basic aqueous solution is equivalent to increasing its ionic strength; thus, when the NaOH concentration was high, the ionic strength was also high and the hydrogen bonds were weakened. This may have increased the clearance between polymer chains owing to the weakened hydrogen bonds form the higher ionic

strength of NaOH [29]. This suggests that the concentration of the basic aqueous solution contributes greatly to the surface and cross-sectional structures of chitosan membranes.

**Figure 6.** Effect of the chitosan molecular weight on the porosity and concentration of the aqueous NaOH solution during membrane formation.

### *3.3. Mechanical Strength*

Figures 7 and 8 show the effect of the NaOH concentration on the maximum stress and maximum strain, respectively, at the time of membrane rupture. Increasing the NaOH concentration decreased the maximum breaking stress and the maximum strain. These trends might be explained by the decreased number of hydrogen bonds and weakening bonds between chitosan molecules as the NaOH concentration increased. The membrane prepared from chitosan with a high molecular weight exhibited greater mechanical strength than the other two membranes. This may be explained by the stronger crosslinking of its polymer chains, which contained many crosslinking points [30].

**Figure 7.** Effect of the NaOH concentration on the maximum stress of the three chitosan membranes at the time of rupture.

**Figure 8.** Effect of the NaOH concentration on the maximum strain of the three chitosan membranes at the time of rupture.

### *3.4. Water Permeability*

Figure 9 shows the effect of the crosslinking time on the water flux through the chitosan membranes prepared in 1.0 and 5.0 mol/L NaOH. The water permeation flux decreased with increasing crosslinking time, regardless of the NaOH concentration. At longer crosslinking times, the interior of the membrane grew denser and suppressed the water flux. Lengthening the crosslinking time probably increased the tortuosity of the permeation pathway through the membrane.

**Figure 9.** Effect of crosslinking time on the water permeation flux through chitosan membranes prepared in 1.0 and 5.0 mol/L NaOH.

Figure 10 shows the effect of the operating pressure on the water permeation flux through the membranes prepared from chitosan with different molecular weights. The water permeation flux increased linearly with the operating pressure for all membranes and was highest for the membrane

formed from chitosan with a low molecular weight. This appears to be because the molecular chain length of the chitosan influences the water permeation pathway through the membrane.

**Figure 10.** Effect of the operating pressure on the water flux through the chitosan membranes (*t*<sup>N</sup> = 180 min).

Figure 11 plots the water flux as a function of porosity for the membranes prepared from different-molecular-weight chitosan in different NaOH concentrations. Increasing the molecular weight of chitosan increased the volumetric water content and decreased the water permeation flux of the membrane. In general, the water permeation path increased with porosity. These results suggest that many moisture regions were immobilized by the molecular chains in the cell structure of the chitosan membrane. These regions could not function as permeation pathways for water. However, in the membrane formed from chitosan with a low molecular weight, the volumetric water content decreased and the permeation flux of pure water increased. The superior water permeation performance of this sample can be explained by the molecular chain length of the chitosan.

**Figure 11.** Correlation between the pure water flux and porosity of the chitosan membranes in different NaOH concentrations.

### *3.5. Mass Permeation Performance of the Chitosan Membranes*

Figure 12 plots the glucose inhibition rate a function of the crosslinking time for chitosan membranes prepared in 1.0 and 5.0 mol/L NaOH. Increasing the crosslinking time increased the glucose inhibition rate of the chitosan membranes in 1.0 mol/L NaOH. This trend can be explained by the reduced number of pores in the membrane as the crosslinking time elapsed; this blocked or narrowed the permeation channels to below the molecular size of glucose (8.7 Å). In contrast, the inhibition rate of membranes prepared in 5.0 mol/L NaOH did not change significantly after 120 min. Since the crosslinking in the membrane was rapidly completed during the formation process with excessive NaOH, the diameters of the mass permeation channels may have been robust against extended crosslinking times.

membranes prepared in different NaOH concentrations (ΔP **Figure 12.** Effect of crosslinking time on the apparent rejection rate of the glucose ratio for chitosan membranes prepared in different NaOH concentrations (∆P = 100 kPa).

When molecules with high and low molecular weights are separated with a polymer, the membrane must block the target polymer and pass the smaller molecules. This property can be evaluated with the membrane fractionation performance. The fractional molecular weight of a membrane is defined as the molecular weight at which the apparent rejection is 90% or more. Since the separation by molecular weight is non-uniform, the fractional molecular weight covers a range of molecular weights. Thus, the molecular weight cutoff is an important performance index of biopolymer membranes and is a helpful guide for selecting a suitable membrane for a given purpose.

(327 Da) was 100%. At 60 ≤ MW ≤ 600, the blocking rate changed remarkably with the m Figure 13 shows the fractional molecular curve of membranes prepared from chitosan with different molecular weights in 1.0 mol/L NaOH. The rejection rate increased with the molecular weight of the chitosan. The inhibition rate of particles with molecular weights above methyl orange (327 Da) was 100%. At 60 ≤ MW ≤ 600, the blocking rate changed remarkably with the molecular weight, which indicates that a fractional molecular weight was identified. This range includes the molecular weights of many functional food components such as amino acids, saccharides, and food polyphenols.

**Figure 13.** Fractional molecular curves of the chitosan membranes.

### **4. Conclusions**

327 Da. At 60 ≤ MW ≤ 600, the blocking rates changed remarkably with the Membranes were successfully prepared from chitosan powder of different molecular weights of chitosan and with different alkali treatments. The volumetric water content of the chitosan membranes increased with the NaOH concentration regardless of the molecular weight of the chitosan. The membrane prepared from chitosan with a high molecular weight exhibited greater mechanical strength than the other membranes. The molecular weight and alkali treatment significantly affected the water permeation flux and mass transfer the prepared chitosan membranes. The water permeability was highest in the membrane prepared from chitosan with a low molecular weight. The water permeation flux increased 1.8-fold as the NaOH concentration was raised from 1.0 to 5.0 mol/L. The membranes had an inhibition rate of 100% for tested components with molecular weights above MW = 327 Da. At 60 ≤ MW ≤ 600, the blocking rates changed remarkably with the molecular weight, which indicates that a fractional molecular weight was identified.

In this study, the volumetric water flux increased with the NaOH concentration and molecular weight of the chitosan. The findings regarding the dominant role of the alkali treatment on both the physical properties and water permeability of chitosan membranes will help facilitate the production of chitosan membranes as a separation technology for water treatment and environment-compatible engineering.

**Author Contributions:** Conceptualization, R.-i.N.; methodology, R.-i.N.; formal analysis, K.K. and Y.N.; investigation, K.K. and Y.N.; writing-original draft preparation, R.-i.N.; writing-review and editing, R.-i.N.; visualization, R.-i.N., K.K. and Y.N.; supervision, N.N. All authors have read and agree to the published version of the manuscript.

**Funding:** This research received no external funding.

**Acknowledgments:** The authors sincerely thank Masanao Imai of Nihon University, who provide the rheometer for measuring mechanical strength of the membrane.

**Conflicts of Interest:** The authors declare no conflict of interest.

### **References**

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### *Article* **Treatment of Olive Mill Wastewater through Integrated Pressure-Driven Membrane Processes**

### **Aldo Bottino, Gustavo Capannelli, Antonio Comite \* , Camilla Costa, Ra**ff**aella Firpo, Anna Jezowska and Marcello Pagliero**

Department of Chemistry and Industrial Chemistry, University of Genoa, via Dodecaneso 31, 16141 Genova, Italy; bottino@chimica.unige.it (A.B.); gustavo.capannelli@gmail.com (G.C.); camilla.costa@unige.it (C.C.); firpolella00@gmail.com (R.F.); aniajez@yahoo.com (A.J.); marcello.pagliero@unige.it (M.P.)

**\*** Correspondence: antonio.comite@unige.it; Tel.: +39-0103536197

Received: 14 August 2020; Accepted: 9 November 2020; Published: 11 November 2020

**Abstract:** The disposal of wastewater resulting from olive oil production (olive mill wastewater, OMW) is a major issue for olive oil producers. This wastewater is among the most polluting due to the very high concentration of organic substances and the presence of hardly degradable phenolic compounds. The systems proposed for OMW treatment are essentially based either on conventional chemical-physical, biological and thermal processes, or on membrane processes. With respect to conventional methods, membrane processes allow to separate different species without the use of chemicals or heat. This work deals with the use of the integrated pressure-driven membrane processes for the treatment of OMW. They consist of a first stage (microfiltration, MF) in which a porous multichannel ceramic membrane retains suspended materials and produces a clarified permeate for a second stage (reverse osmosis, RO), in order to separate (and concentrate) dissolved substances from water. Laboratory scale experiments with different small flat sheet RO membranes were first carried out in order to select the most appropriate one for the successive bench scale tests with a spiral wound module having a large membrane surface. The aim of this test was to concentrate the dissolved substances and to produce water with low salinity, chemical oxygen demand (COD), and reduced phytotoxicity due to a low content of phenolic compounds. The trend of the permeate flux and membrane retention as a function of the volume concentration ratio was investigated. The influence of OMW origin and its aging on the membrane performance was also studied.

**Keywords:** olive mill wastewater; membrane separation process; microfiltration; reverse osmosis; water recovery

### **1. Introduction**

Olive oil mill wastewater (OMW) is a by-product of the olive oil extraction process produced seasonally in a large quantity. Niaounakis and Halvadakis in their book [1] estimated a generation of OMW in the range of 10–30 million m<sup>3</sup> /year in 2006 and we should expect that since then, its quantity has increased in accordance with the increase in world olive oil consumption, which from 2006 to 2019 has grown from about 2.6 to 2.97 million tons [2,3].

### *1.1. OMW Composition*

The OMW consists mainly of olive fruit vegetation water (more than 50% of the fruit) and water added during the extraction process. The composition of OMW is affected by the variety and ripeness of the olives and by the system used for their processing (pressure or centrifugation mills). For example, the centrifugation step in three-phase olive mill processing, the most common olive oil extraction

system, generates an amount of OMW more than two times higher than that of olive oil produced. An average OMW composition can be given as 83.2% of water, 1.8% of inorganic salts and 15% of organic constituents, among which 7.5% of sugars [4]. OMW is characterized by a low pH, a high electrical conductivity and a chemical oxygen demand (COD), which can be as high as 200 g/L. The three-phase process (3P) generates the greatest amount of OMW, about 1–1.2 m<sup>3</sup> /tons of olives, while the two-phase process generates the least amount, about 0.085–0.1 m<sup>3</sup> /tons of olives. The batch-pressing process produces about 0.4–0.6 m<sup>3</sup> /tons of olives of OMW [5]. Nevertheless, all the three types of OMW are highly pollutant. Due to the presence of several organic compounds, among which there is a phenolic fraction, untreated OMW has broad-spectrum toxicity against bacteria, plants and animals [6], which implies treatment and environmental problems. However, phenols presence in OMW makes this problematic by-product (wastewater) a potential source for recovery of precious antioxidants. For the abovementioned reasons, OMW treatment systems are not only supposed to be flexible and efficient in reducing COD and salinity, they also should be a viable alternative for recovery of high added value phenolic compounds.

### *1.2. OMW Membrane-Based Treatment Processes*

The systems proposed for OMW treatment are essentially based either on conventional biological, chemical, physicochemical and thermal processes [7–9] or advanced membrane processes [10–16]. The latter, especially pressure-driven processes (microfiltration (MF), ultrafiltration (UF), nanofiltration (NF) and reverse osmosis (RO)), offer several advantages over traditional technologies, mainly in terms of low energy consumption, no additive requirements and no phase change, and thus, the possibility to preserve the original characteristics of treated effluents.

Gebreyohannes et al. [17] in 2016 reviewed both the literature and patents about the application of integrated membrane technologies for OMW treatment and they highlighted the polarization and fouling problems occurring in the pressure-driven membrane processes, which are mainly related to the particular composition of the OMW (e.g., solids, pectins, etc.). Again in 2016, Pulido [18] reviewed in detail the open literature on the application of membrane technologies in OMW treatment as well as on the main obstacles for their cost-effective utilization, namely the related fouling problems. He highlighted the need for a pretreatment before the integrated membrane process to limit the fouling phenomena and to achieve more stable operating permeate fluxes.

Typically, the proposed integrated pressure-driven membrane processes are based on the combination of steps for the removal of suspended solids (e.g., microfiltration or ultrafiltration) and of a second step aimed at the pollutant concentration and clean water recovery (e.g., nanofiltration and/or reverse osmosis). A fractionation of the pollutants contained in the OMW is technically feasible [19,20] but the application of such a process scheme composed by several steps of MF and UF with different molecular weight cut-off (MWCO), NF and RO is very expensive and often quite sophisticated for its implementation into small and medium olive mills.

Membrane processes have been applied to all the three types of OMWs (Table 1). The content of total suspended solids (TSS) and others minor components such as fats and pectins makes imperative a feed pretreatment before the NF or RO processes. Considering that the pH for the types of OMWs is similar, the electrical conductivity (EC) reflects the concentration of organic electrolytes and salts. The two-phase process shows the lowest electrical conductivity or solid residue. On the other hand, the two-phase process generates a solid, known as alperujo, which is a very pollutant waste to handle since it contains most of the organic compounds that in the three-phase process are released in the wastewater [5].


**Table 1.** Analytical characteristics of olive mill wastewater (OMWs) from the batch (BP), two-phase (2P) and three-phase (3P) processes and some integrated membrane processes proposed in literature.

EC = Electrical conductivity; TSS = Total Suspended Solids; TOC = Total Organic Compounds; COD = Chemical Oxygen Demand; Ph = Polyphenols; Ch = Carbohydrates.

Cassano et al. [9] studied the application of UF polymeric membranes (MWCO between 4 and 10 kDa) and they observed a flux decrement up to 50% over 300 min operating time. The best performing membrane was made of regenerated cellulose. The flux recovery after cleaning with an alkaline detergent at 40 ◦C was claimed enough to recover the initial water flux. In any case, the raw OMW was subjected to preventive microfiltration step at 0.2 micron. Garcia Castello et al. [11] studied the combination of MF, NF followed by an osmotic distillation. In the MF step a 0.2 µm membrane was used and a strong flux decrease was observed without any tendency of stabilization. The cleaning procedure was carried out by using a concentrated alkaline solution of 20 g/L NaOH at 40 ◦C for 30 min followed by tap water rinsing. An irreversible fouling was observed with a loss of flux of about 50%. The flux reduction in the NF membrane (Nadir N30F spiral-wound membrane module) was about 35% after about 1 h operation at a volume reduction factor of about 3. The NF membrane after cleaning with 1g/L of NaOH as done for MF completely recovered its initial water flux. From the cited investigations it seems that although UF underwent severe fouling the initial flux could be recovered in most of the cases by a chemical alkaline cleaning procedure.

Bazzarelli et al. [28] proposed an integrated membrane process based on a MF/NF and osmotic distillation and membrane emulsification. For the MF step, a 0.14 µm ceramic membrane was used and the good results in the MF flux stability were ascribed to an acidification step at pH 1.8 and a subsequent filtration on a stainless steel filter [29]. The chemical cleaning protocol was still based on the use of an alkaline detergent at 40 ◦C for 30 min. Chemical physical pretreatments before the integrated pressure-driven membrane processes were studied in order to improve the performance of the integrated membrane process. Pulido et al. [30] applied a pretreatment based on a Fenton process, then directly followed by NF. Nevertheless, the direct application of tighter membrane processes (NF or RO) after a physical chemical secondary treatment can lead to cake formation on the membrane surface as reported for RO membranes [31].

Recently, the possibility of using a water-ethanol mixture for the extraction of polyphenols and their purification by integrated membrane processes was explored [32]. Although it opens up an interesting perspective, additional investigations should be carried out in order to define the quality of the reverse osmosis permeate and its ethanol content. With the aim of recovering valuable polyphenols, most of the studies investigated the integration of ultrafiltration (UF) and (NF). De Almeida et al. [22] showed that despite the combination of UF and NF, the COD and total phenols removal can be 83.3 and 93.1%, respectively. Despite the interesting results, the quality of the permeate water is still far from being disposed in the sewage under the parameters imposed by the legislation. Therefore, to meet the current disposal regulations a further treatment or filtration step of the NF permeate is clearly necessary.

Another integration scheme relied on the direct use of RO instead or in addition to the NF. Tundis et al. [33] showed the recovery and classification of polyphenols by using a MF step on a 0.1 µm TiO<sup>2</sup> membrane followed by a NF step and a RO step based on a membrane typically applied to brackish water. Although a flux decay was observed for all the membrane filtrations, as the aim of the work was about the characterization of the polyphenols in the concentrate, the quality of the final NF and RO permeates was not assessed by the authors. Zagklis et al. [34] in a recent paper mentioned the design of a full system based on UF/NF/RO integrated with adsorption steps and solid-liquid extraction with the aim of recovering the polyphenolic fractions from both OMW and other types of phenolic containing wastewater (e.g., grape marc and olive leaves). Coskun et al. [35] in their lab scale study proposed centrifugation as a primary step followed by UF and finally by RO. Their study was exclusively focused on the rejection performance of the different membranes. Petrotos et al. [36] studied some relevant operational parameters on a pilot scale, a process integrating MF followed by a NF (or open RO) and then by RO using tubular membranes.

The problem of OMW is clearly urgent from the point view of its environmental impact and the technological solution that requires it to be simple, cost-effective and reliable, especially in countries where the size of working olive mills is still small. Integrated pressure-driven processes, which include RO as a final step, should enable the production of a permeate water of sufficient quality not only for its safe discharge into sewage, but also for any kind of reuse into a farm or olive oil production process.

Membrane processes were shown to be very effective in the treatment of numerous industrial effluents and wastewaters. However, their successful application depends on the proper choice of process configuration and process conditions, and these are the focus of the experimental study presented here. In this work, two consecutive pressure-driven membrane processes, namely microfiltration and reverse osmosis, are proposed for OMW treatment in order to obtain a RO permeate composed of water with a low salinity, COD, and reduced phytotoxicity due to very low content of phenolic compounds, which are retained and concentrated by the RO membrane. To this aim, laboratory scale tests were first carried out with small flat sheet RO membrane samples in order to select the most suitable membrane for the successive pilot scale investigation with a spiral wound element with a large membrane surface area. The high concentrations of suspended materials in OMW imposed the use of microfiltration as a pretreatment system for the RO, in order to avoid plugging of the feed spacer of the spiral wound elements. Ceramic multichannel elements with excellent thermal stability and chemical resistance to withstand severe cleaning cycles were used for microfiltration in order to easily remove particulates that can foul the membrane or plug the channels. This work deals with practical aspects and problems connected to the concentration of large volumes of OMW with MF/RO pilot plants and to OMW storage that were not investigated enough in the literature.

### **2. Materials and Methods**

Since characteristics of OMW may differ significantly from mill to mill, OMWs from three different olive mills, two located in Liguria and one in Tuscany, were employed. The names of these mills cannot be revealed for confidentiality reasons and a generic code composed of letters and numbers will be used to identify the three types of OMWs. OMWs were first stored in reservoir tanks to allow sedimentation of a large part of suspended materials and separation of a supernatant fluid, which was filtered through a filter bag with an opening of 200 µm prior to microfiltration.

Microfiltration of prefiltered OMW was performed in a batch operation mode with the plant schematically shown in Figure 1a, using three ceramic membranes (Membralox EP19-40, Pall Corp., Port Washington, NY, USA) arranged in parallel into a stainless steel housing (Figure 1b). The main properties of these membranes are shown in Table 2.

**Figure 1.** (**a**) Schematic of the plant used for MF tests: V1-9 ball valves; GV gate valve; P1,2 manometers, T thermometer; F flowmeter. (**b**) Membralox module: 1—ceramic multi-channel membranes, 2—stainless steel membrane housing, 3—module end-cup, 4—clamps.

**Table 2.** Main properties of Pall–Membralox EP19-40 membrane used for MF tests.


The MF retentate was completely recycled to the feed tank while the clean permeate was continuously withdrawn to be used for RO test. As can be seen from Figure 1a, OMW is fed by the centrifugal pump to the membrane module with a velocity v = 3.9 m/s (calculated from the ratio between the feed flow rate measured by the flow meter, F, and the membrane channels cross-section) at an average pressure P = 2.3 bar, unless otherwise reported, measured by two manometers, P1 and P2, located before and after the membrane module, respectively. The permeate flow rate is simply evaluated by measuring with a graduated tank the time necessary to produce a given permeate volume. Permeate flux is then calculated from the ratio between the permeate flow rate and the overall filtration surface area. The temperature measured by the thermometer, T, is kept constant at 30 ◦C by a cooling device immersed into the feed tank. An electric immersion heater in the cleaning tank provides a rapid heating of the cleaning solutions (NaOH and/or NaOCl solutions) used to remove foulants from the membrane.

The scheme of the RO plant is very similar to that of the MF plant shown in Figure 1a. The main differences are related to the feed pump (piston), the pressure control valve (globe valve), and the use of a variable frequency drive 'inverter' to control the feed flow rate, Q (speed pump). The experimental conditions adopted for RO experiments were: P = 30 bar (unless otherwise reported), Q = 1000 L/h, T = 25 ◦C. A small cell was used for preliminary tests with flat sheet membrane samples (surface 0.0066 m<sup>2</sup> ) listed in Table 3. A cylindrical vessel was employed to house a spiral wound membrane module (SW30HR Dow-Filmtec, now DuPont, Wilmington, Deleware; 4" diameter, 40" length, membrane surface 7.9 m<sup>2</sup> ) during the successive bench scale experiments. Preliminary tests with small flat membranes were carried out keeping the feed concentration constant, and continuously recycling both permeate and concentrate streams to the feed tank. Concentration tests with spiral wound module were performed in a batch operation mode, following the same procedure previously described for microfiltration of OMW. During both RO and MF experiments samples of different streams were collected for analysis.


**Table 3.** NF and RO membrane used during test cell experiments.

Electrical conductivity, pH and suspended solids content were measured according to Standard Methods [37]. COD was determined with the spectrophotometric method using Merck Spectroquant@ test kits (Merck KGaA, Darmstadt, Germany). The method is analogues to EPA 410.4, US Standard Methods 5220 D, and ISO 15705. Phenols were determined with Folin–Ciocalteu reagent [38].

### **3. Results**

### *3.1. Feed Pretreatment and Microfiltration*

OMWs with quite different characteristics were received from three mills. In particular, the OMW3-SG was characterized by a very high load of suspended solids of small size with a negligible settling velocity and poorly retained by the filter bag as can be seen in Figure 2a, where the images of the three types of OMWs after settling and filtration treatment are reported for comparison. The darker color of samples OMW1-FR and OMW2-CA is connected to a high particle removal efficiency. However, even in these two cases (especially for OMW2-CA) the produced filtrates did not satisfy the requirements for the RO feed. This is apparent from the images of Figure 2b, where deposited solids after centrifugation (8000 rpm for 10 min) can be observed on the bottom of the centrifuge tube. Therefore, a post-filtration treatment with ceramic membrane with 0.2 µm pore size was employed for the removal of fine suspended solids and production of a suitable feed for RO [39]. The main physicochemical characteristics of the three types of OMW (after settling and bag filtration) fed to the microfiltration plant are reported in Table 4.

(**a**) (**b**)

**Figure 2.** (**a**) Images of OMWs after settling and bag filtration. (**b**) Images of OMWs after centrifugation (h = TSS volume).


**Table 4.** Physicochemical properties of OMWs fed to the microfiltration plant.

Figure 3 refers to the microfiltration tests and shows the behavior of permeate flux as a function of volume concentration ratio, VCR (i.e., the ratio between the volume of the initial feed and the volume of the final concentrate) for the three different pretreated (settled and filtered) OMWs. Gentle heating at 30 ◦C makes the feed (especially the OMW3-SG) more fluid with consequent reduction of the friction loss along the plant and improved performance of the centrifugal pump. The permeate flux at the beginning of the MF tests appears very close for the three OMWs, while increasing the VCR, OMWs behave differently, especially OMW3-SG. As far as OMW1-FR and OMW2-CA are considered, the permeate flux first slightly decreases and then tends to level off. With OMW3-SG, which contains a relevant amount of suspended materials, a strong and almost proportional decline of permeate flux with increasing the VCR is observed.

**Figure 3.** Permeate flux versus volume concentration ratio (VCR) during the MF tests with different OMWs.

After a VCR = 2.1, the high viscosity of the concentrated OMW3-SG (Figure 4a) considerably reduces the performance (head and flow rate) of the centrifugal pump, thus the fluid velocity through the membrane channels is progressively lowered and membrane channels begin to plug. The permeate flux first falls and then continues the decrease slowly. Immediately after VCR = 2.85, a sudden increase of the pressure occurred and the pump reached its shut-off pressure. The test had to be stopped immediately and it was necessary to use a metal probe to unclog membrane channels (Figure 4b).

≈ Moreover, an intense membrane cleaning with NaOH solution (2% *w*/*w*) and NaOCl (500 ppm Cl) at 60 ◦C for at least 60 min was used to remove the deposit remaining on the membrane surface and into the membrane pores. By measuring pure water flux before (JW,0) and after (JW,F) OMW filtration, a flux recovery ratio FRR = (JW,F/JW,0)·100 very close to 100% was achieved, thus demonstrating the effectiveness of the cleaning procedure. The other two types of OMWs (1-FR and 2-CA) did not plug membrane channels but severely fouled the membrane. The pure water flux after MF was around 30% of the original membrane flux, but even in this case a FRR ≈ 100% was obtained after the cleaning with NaOH and NaOCl.

**Figure 4.** (**a**) OMW3-SG viscous concentrate after VCR = 2.1. (**b**) Removal of OMW3-SG muddy concentrate from the membrane channel with a metallic probe.

The main physicochemical properties of feed (FD) and permeate (PR) samples collected at increasing VCR during the MF of the three types of OMW are listed in Table 5. Both pH and electrical conductivity of feed and permeate are substantially similar since dissolved ions pass through the pore of the membrane while a given retention is observed for COD due to the removal of suspended organic part, which contributes to this parameter. Phenol retentions seem to be high for the MF membrane, but according to previous literature findings [8], this fact can be ascribed to fouling, which may deeply alter the retention characteristics of membrane by itself.


**Table 5.** Main physicochemical properties of feed (FD) and permeate (PR) samples collected during the MF tests with three different types of OMW.

### *3.2. Nanofiltration and Reverse Osmosis*

The results of NF/RO screening tests with small flat sheet membranes carried out by using the MF permeate of OMW2-CA as feed are reported in Table 6. Except for Desal DK, all the other membranes present very high solute retention. To obtain useful products from OMW such as purified water (permeate) and a polyphenols rich solution (concentrate), a membrane with the highest possible retention to salts, COD and phenols are required. Table 6 reveals that SW30HR membrane (DOW) completely meets these requirements. Therefore, this membrane in a spiral wound configuration was selected for successive bench-scale tests.

**Table 6.** Results of the RO screening tests with different types of NF and RO membranes (P = 30 bar, T = 25 ◦C). Feed: MF permeate of OMW2-CA (Conductivity = 13,200 µS/cm; COD = 40,180 mg/L; Phenols = 1070 mg/L).


The results of the RO concentration test carried out with the MF permeate of OMW1-FR are shown in Figure 5. By increasing the VCR, the permeate flux decreases first rapidly and then slowly until reaching VCR = 10.5, a value (around 1 L/(m<sup>2</sup> ·h)) ca. 30 times lower than that of the initial flux (VCR = 1). The observed flux decline with increasing VCR can be ascribed to the increase of the osmotic pressure of the feed, as well as concentration polarization and fouling phenomena.

**Figure 5.** Permeate flux as a function of VCR during RO test with the MF permeate of OMW1-FR.

The osmotic pressure of the concentrated solution (VCR = 10.5) can be estimated by measuring the permeate flux at increasing pressures and a constant feed concentration as shown in Figure 6. The differences in pure water flux (Figure 7) measured before and after OMW treatment are connected to the membrane fouling. Only a moderate cleaning with a NaOH solution (pH = 11) at 40◦C was sufficient for eliminating fouling and achieving FRR around 100%.

Figures 8 and 9 show the behavior of the permeate flux as a function of VCR during the RO concentration of the permeates produced by microfiltration of OWM2-CA and OWM3-SG. The trends are similar to that shown in Figure 5. The permeate flux improves at higher pressure but continues to fall with the increase in the VCR. A worse membrane performance is observed according to the considerably higher solute content of these OMWs as shown in Table 7. Further inspection of Table 7 reveals high retention values for conductivity and COD and an excellent abatement of phytotoxic phenol fraction. As expected, the retention worsens with VCR and improves with the pressure (since water flux through the membrane increases with the pressure while the solute diffusion is independent of pressure).

**Figure 6.** Permeate flux as a function of operating pressure at VCR = 10.5 (Feed: MF permeate of OMW1-FR).

**Figure 7.** Pure water flux as a function of operating pressure before and after RO test with MF permeate of OMW1-FR.

**Figure 8.** Permeate flux as a function of VCR during RO test with the MF permeate of OMW2-CA.

*Membranes* **2020**, *10*, 334

**Figure 9.** Permeate flux as a function of VCR during RO test with the MF permeate of OMW3-SG.

**Table 7.** Main physicochemical properties of feed (FD) and permeate (PR) samples collected during RO test with the MF permeates of three different types of OMW.


The effect of OMW age on the performance of the membrane is shown in the following Figures 10 and 11 and in Table 8. It is worth noticing that olive oil extraction is a seasonal operation whose duration is around 4–5 months during the winter. The amount of OMW is much higher than that of olive oil produced, and consequently very large plants are necessary for the treatment of all the wastewater generated daily, otherwise it must be stored. To obtain preliminary information on the influence of OMW storage/aging on the performance of the integrated membrane process, a given amount of OMW2-CA was allowed to rest for ca. 4 months. After this long settling period, the supernatant liquid was filtered through the usual filter bag and the resulting filtrate was sent to the MF plant. From the results reported in Figure 10, a given increase of the permeate flux of the stored OMW is observed. This increase is connected to a lower content of suspended material (TSS = 420 mg/L) due to the 4 months settling period. Conversely, only a moderate variation of the permeate flux during the RO experiments (Figure 11) occurs since the amount of dissolved solids does not practically change

during the storage, as can be seen from physicochemical characterization results shown in Table 8. From the same table it is apparent that the storage period does not affect membrane retention.

**Figure 10.** Effect of storage of OMW2-CA on the permeate flux as a function of VCR during MF test.

**Figure 11.** Effect of storage of OMW2-CA on the permeate flux as a function of VCR during RO test with the MF permeate of OMW2-CA.

**Table 8.** Main physicochemical properties of feed (FD) and permeate (PR) samples collected during the RO test with "aged" OMW2-CA.


### **4. Discussion**

### *4.1. Pretreatment and Microfiltration*

As reported in Tables 1 and 3, OMW contains relevant concentrations of TSS. Therefore, any type of membrane process aimed at polyphenol recovery as well as water reuse needs a pretreatment to remove TSS. The removal of TSS is of crucial importance for the fouling control of the NF or RO process.

Cassano et al. [12] pretreated the raw OMWs by using a commercial tubular MF membrane module (pore size 0.2 µm, polypropylene, 5.5 mm inner diameter). Then, UF polymeric membranes were used to produce a clear permeate to be fed to the nanofiltration unit. Nevertheless, for all the UF membranes, a flux decay was observed. Bazzarelli et al. [29] studied the change of pH to destabilize the solid suspension in OMW and they showed that a pretreatment based on MF or UF can be effective at removing the suspended solids. MF exhibits higher fluxes than UF and ceramic membranes showed the highest fluxes. In another interesting approach of MF by using polymeric hollow fiber membranes, a fouling control was attempted by the deposition on the membrane surface of a photoactive gel [40].

Garcia–Castello et al. [11] reported the performance of a 0.2 µm alumina membrane after several filtration runs. During each run, a consistent flux decay was observed and even after a cleaning procedure with 20 g/L NaOH at 40 ◦C the water flux of the virgin membrane was not fully recovered.

In our work, we proposed the use of a ceramic MF membrane due to its high chemical and thermal stability during the cleaning procedures to restore its performance. We also observed an evident decay of the flux (Figure 3), but on the other hand, a chemical cleaning with alkaline agents combined with the use of a sufficiently high temperature (about 60 ◦C for at least 1 h) we restored the initial membrane performance. The effect of the temperature during the chemical cleaning was remarked in a recent work also by Fraga et al. [41], where the use of high permeability silicon carbide MF membranes was investigated.

The MF ceramic membrane module tested with all the three OMWs was able to preserve the subsequent NF or RO spiral wound modules from plugging problems. Since MF can seriously suffer from plugging and fouling phenomena at high TSS, the use of chemically-resistant membranes seems to be essential, especially if the plant is designed to be used only seasonally.

### *4.2. Nanofiltration and Reverse Osmosis*

SW30HR membrane showed the best retention of COD and phenols among the tested NF and RO membranes. By increasing the VCR, the retention to phenols was always very high (>99.3%) and the retention to COD had generally been about 95%. The highest VCR obtained was limited by the increase of the osmotic pressure. As shown, at VCR = 10.5 the experimental osmotic pressure was approaching 29 bars. The electrical conductivity of pristine OMW1-FR was 5310 µS/cm, while for OMW2-CA and OMW3-SG, the electrical conductivities were very close, 13,940 µS/cm, 12,780 µS/cm, respectively (Table 4). Since the electrical conductivity is mainly related to the concentration of dissolved salts, with the OMW1-CA it was possible to achieve a higher VCR than for the other two OMWs. The different behavior between OMW2-CA and OMW3-SG during the RO concentration is therefore mainly related to the different level of organic compounds, considering that the ratio of COD between the OMW3-SG and OMW2-CA is about 2. A pressure increase seems to be beneficial to both the flux and the retention.

### *4.3. Remarks*

On the basis of our results and of the findings reported in the literature, integrated membrane processes are able to efficiently produce a polyphenols-rich concentrate. The recovery of polyphenols is very interesting, since they are valuable compounds that can be supplied to cosmetic and pharmaceutical industries. Nevertheless, the exploitation of polyphenols-rich streams is still facing some technological challenges related to the polyphenol fractionation. The main driver to develop processes for the treatment of OMW is the environmental pressure in order to limit the pollution related to their production and disposal. A clean water stream can be obtained when an RO process is considered as a final step. In the proposed integrated membrane process, the high retention of polyphenols can allow the separation of good quality water already after a first RO stage, which can be more easily accepted by a sewage depuration system since the residual COD is no longer related to the presence of polyphenols. The permeate water can be considered also for an internal reuse in the olive mill after and eventual refining treatment (a second RO stage or adsorption) as well as for irrigation purposes. The main process issues are related to controlling the fouling. Ceramic membranes have proved their suitability since they can withstand aggressive chemical cleaning procedures, and although their cost is still high compared to that of polymeric membranes, they can guarantee a longer lifetime. Since in many countries olive mills are still small enterprises, the investment costs for a membrane-based treatment plant can be more easily faced if there is the possibility of storing part of the OMWs generated during the milling season. In this work, we proved that the aging of the OMW does not critically affect the performance of the integrated membrane process.

Pulido and Martinez-Ferez in their review [42] identified the control of fouling as one of the limits of membrane technologies applied to OMW. Another bottleneck for a wide field application of the integrated membrane process remains, related to the options available for either the disposal or chemical/energetic valorization of the concentrate stream. These options should be evaluated on the specific characteristics and constraints of the olive mill willing to apply membrane technology.

### **5. Conclusions**

MF/RO integrated membrane processes have been proposed for the treatment of OMW. The MF can be considered a suitable pretreatment for RO process since it provides a clean permeate, which does not cause plugging of the spiral wound element. RO separates dissolved substances from water, thus allowing the concentration of valuable products and produces water with a low salinity, COD, and phytotoxicity. Channel plugging and fouling of MF membrane represent a serious problem during the treatment of OWM, characterized by a high load of fine particles, which cannot be properly removed by simple settling or bag filtration. Therefore, ceramic membranes capable of withstanding hard cleaning agents are necessary. Membrane performance is not deeply affected by OWM aging and consequently, the wastewater may be treated gradually, without the need of large plants operating only a few months a year. This may involve important benefits connected to the reduction of investment costs and of bactericide solutions, which are necessary for a long-term storage of delicate RO membranes. The commercial RO membrane for seawater treatment, SW30 HR (Dow), showed a very high retention to polyphenols and dissolved species, which contribute to electrical conductivity. The increase of both osmotic pressure and organics concentration limited the maximum volume concentration ratio that could be achieved.

**Author Contributions:** Conceptualization, A.B. and G.C.; data curation, C.C. and A.J.; funding acquisition, G.C.; investigation, A.B., R.F., A.J. and M.P.; methodology, A.B. and A.J.; project administration, G.C. and A.C.; resources, A.C. and A.J.; supervision, A.B. and A.C.; visualization, A.J. and M.P.; writing—original draft, A.B. and A.J.; writing—review & editing, A.C. and M.P. All authors have read and agreed to the published version of the manuscript.

**Funding:** This research was conducted in the framework of the ENPI CBC Med Project: MEDOLICO I-B/2.1/090 (Mediterranean Cooperation in the Treatment and Valorisation of Olive Mill Wastewater, OMW).

**Conflicts of Interest:** The authors declare no conflict of interest. The funders had no role in the design of the study; in the collection, analyses, or interpretation of data; in the writing of the manuscript, or in the decision to publish the results.

### **References**


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