Bubble dynamics, mainly including path instability, breakup, and coalescence, are significantly related to the bubble mass transfer. When a bubble rises freely in still water, it is imposed by buoyancy force, drag force, and lift force, as shown in Figure 2A. Its path can switch to a zigzag or helical trajectory after traveling a straight vertical line, usually accompanied by significant bubble deformations and surface oscillation [6] (Figure 3). The path instability is caused by the symmetry breakage of the wake vortex behind the bubble [7, 8], which is governed by the bubble size [8].

There are extensive studies on the dynamics and mass transfer of a single bubble in still liquids, most focusing on the effects of bubble dynamics, bubble shape and trajectory, etc. Generally, the bubble path instability promotes convective transport and thereby strengthens mass transfer [9]. The interface oscillations induced during path instability also enhance the mass transfer near the interface [10]. The zigzag rising trajectory can accelerate the slippage of bubbles and reduce the thickness of the bubble concentration boundary layer, leading to enhanced local mass transfer [11]. The helical rising bubbles can generate an asymmetric wake with a lower terminal velocity than non-helical rising bubbles, and thus, leads to a less efficient mass transfer. The mass transfer coefficient in the bubble columns is also known to be dependent on the bubble dynamics. However, bubble swarms perform a different path instability from a single bubble, due to the turbulence induced among the moving bubbles (Figure 2B). As bubble swarms rise in liquid, bubbles coalescence or rebound frequently, which causes difficulty in determining their trajectories [12, 13].

The behaviors of bubble in turbulent flow are usually characterized by breakup and coalescence. Bubble breakup usually occurs under the effects of turbulent fluctuation and collision, viscous shear stress, shearing-off process, and interfacial instability [14].

Bubble coalescence is more complex than bubble breakup [4], which is not only affected by hydrodynamics, but also dependent on the effect of interfacial characterization. Bubble coalescence is caused by the relative bubble motions induced by the turbulence in the continuous phase, mean-velocity gradients, and buoyancy [15]. Large-size bubbles formed by coalescence exhibit a lower mass transfer efficiency [16]. When they break up into smaller bubbles, the mass transfer process is promoted. In addition, Tse et al. [17] found that, in some instances, the coalescence of two bubbles was accompanied by the formation of a much smaller daughter bubble generated by the annular wave following the breakup process. In coalescence-dominated systems, such processes can generate significant numbers of small bubbles [17]. Essentially, the path instability affects bubble diffusion and controls the local concentration in still water, while the coalescence and breakup of bubbles affect the size distribution of bubbles in turbulent flow [4]. Great turbulence caused by bubble swarm motion under the effect of buoyancy, can result in an enhanced mass transfer.

Most of the industrial reactors are operated under turbulent flow conditions. The dynamic behaviors of bubbles in turbulent flows are dramatically more complex than those in still water due to the random character of the turbulent fluctuations. The flow is not steady and the fluctuating conditions within turbulent flows can lead to deformation, breakup and self-sustained oscillations in the bubble [18]. Secondly, different from still water, there are regions of high vorticity and low-pressure vortices in turbulent flows [4]. The spatial structure of the velocity field around the bubble is unknown and depends on different eddies. For a single bubble in turbulent flow, the bubble mass transfer is strongly dependent on the Reynolds number and the Schmidt number [19]. The bubble Reynolds number based on the bubble slip velocity is used to reflect the bubble-liquid relative motion, and the liquid Reynolds number to characterize the turbulent intensity [20]. The relative velocity between bubble and liquid, and the liquid turbulence both have an important effect on the mass transfer. An increase in the liquid turbulence intensity contributes to the improvement of the surface renewal rate, and thus the enhancement of mass transfer [20].

The mass transfer of a single bubble in still water is the basis for understanding bubble mass transfer under complex conditions and in various gas-liquid reactors. The mass transfer of a single bubble in still water is a combination of molecular diffusion and convective transport [11]. For a single bubble, the mass transfer coefficient K _L is mainly affected by bubble-liquid contact time, liquid viscosity and bubble equivalent diameter [21]. It decreases with the increase of the bubble-liquid contact time, probably due to the thickened bubble-liquid interface. The increase of liquid viscosity reduces the diffusion coefficient, and results in a lower mass transfer coefficient [22]. The bubble size has an essential effect on the mass transfer coefficient [23]. An increasing bubble equivalent diameter resulted in a decreasing bubble coalescence rate, a larger gas-liquid interface area, and a higher gas-liquid mass transfer coefficient. Smaller bubbles lead to less mass transfer due to small pulsations of turbulence around them. More specifically, for small bubbles with an equivalent diameter of 1–2 mm, the mass transfer coefficient is usually low, as the small bubbles have a rigid surface with negligible surface oscillations and internal circulation [24], which are essential to promote the gas-liquid mass transfer [25]. However, Hori et al. [26] found that, for spherical cap bubbles with an equivalent diameter greater than 5 mm, the mass transfer coefficient decreased with the increase of the equivalent diameter, due to the increasing flattening and aspect ratios of the bubble shape.

The bubble mass transfer is also dependent on the oscillation of the interface. As the size increases, bubbles become deformed and partially circulating. Bubble internal circulation increases with increasing size, and the shape oscillations occur at Re > 200 [11]. The oscillations of the bubble-liquid interface have been reported to improve the mass transfer [10, 27]. The reasons for the reduced mass transfer coefficient can also be the slippage at the bubble-liquid interface and the prevention of mass transfer resulting from the inert gases inside the bubble as well as the surfactants in the liquid phase [28]. For example, Bao et al. [11] reported that impurities (surface-active pollutants) in a contaminated system can influence the mobility of the bubble surface and increase the mass transfer resistance, which results in a smaller mass transfer coefficient than that in clean systems.

At large gas flow rates, bubbles in liquid become dense and the flow is featured with bubble swarms [29]. In practical industrial processes, gas-liquid mass transfer is usually in the form of bubble swarms, where the bubbles present various sizes, shapes and dynamics [30, 31]. Compared to a single bubble, the mass transfer characteristics of bubble swarms are more complex and the number of relevant studies has been much smaller. The mass transfer efficiency of bubble swarms is dependent on various parameters. First, the bubble size affects the gas exchange rate, bubble residence time, fluxes of volume and momentum, and thereby the mass transfer efficiency. Specifically, small bubbles have a high specific surface area, and can significantly improve the mass transfer from the gas phase to the liquid phase. Additionally, the large bubbles can induce considerable turbulence in the fluid, and promote mass transfer in the liquid [4]. According to Sahoo and Luketina [32], small bubbles with a radius of about 1 mm exhibited a higher oxygen transfer efficiency compared to larger bubbles.

Second, the way of bubble injection can affect the mass transfer efficiency. As reported by Gong et al. [33], concentrated injection of bubbles into liquid increased the bubble-induced liquid velocity, while it reduced the mass transfer efficiency. Uniform injection of bubbles performed higher mass transfer efficiencies than concentrated injection [33, 34]. The gas injection nozzle size also has an essential effect on mass transfer efficiency. For bubble swarms, the mass transfer rate increases with the decrease in the nozzle diameter [35]. Third, the size and height of the liquid container have an influence on the mass transfer efficiency of bubble swarms, and the effect also varies with the bubble size [36].

It has been widely accepted that the traditional single-bubble mass transfer theories cannot be applied directly to quantify the complex mass transfer processes of bubble swarms. For bubble swarms, the mass transfer performance is often empirically correlated with the volumetric mass transfer coefficient K _L A [11]. The models for mass transfer of bubble swarms will be summarized in the subsequent sections.

In non-Newtonian fluids, the mass transfer of bubbles is more complicated compared to that in Newtonian fluids. A non-Newtonian fluid has a varying viscosity with the change of shear rate. The viscosity variation of non-Newtonian liquid with shear rate varies with its rheological behavior. Non-Newtonian fluids can be further divided into Bingham fluids, pseudoplastic fluids and dilatant fluids [37]. Taking pseudoplastic fluids as an example, after a bubble being injected into the non-Newtonian fluid, the space expelled by the bubble movement does not recover immediately, which results in a vacuum behind the bubble. In a pseudoplastic fluid, restoring the vacuum region requires more energy than it does in a Newtonian fluid, thus slowing down the bubble movement. This causes a more significant drag coefficient of pseudoplastic fluids on the bubble than that of Newtonian fluids, and the bubble mass transfer coefficient is reduced as a result. However, some researchers found that, once the bubble size in non-Newtonian fluids is increased to a certain degree or the flow index reduced to a certain level, the wake effect, rise velocity and flow field turbulence can be enhanced, thereby increasing the mass transfer coefficient and improving the mass transfer efficiency [38]. The shape of bubbles can also be affected by non-Newtonian fluid. As reported by Bao et al. [11], the shape transition of bubbles in a pseudoplastic liquid is from ellipsoid to upside-down spherical cap, different from that in the Newtonian fluid (from ellipsoid to spherical cap).

The quantity and rheological properties of non-Newtonian material bring about the local acceleration or deceleration on slippage, shear rate, and thus irregular shape shifting and mass flux of gas–liquid interface. In addition, the gas–liquid mass transfer investigation in colored or non-transparent non-Newtonian fluids can be even more difficult, because the usual non-intrusive measurement methods can hardly be able to visual volume change or concentration distribution from these non-transparent fluids. Moreover, some non-Newtonian fluids, such as polyacrylamide solution, can exhibit both non-Newtonian behavior and surfactant-like effect, which may also decrease the mass transfer coefficient and increase the system complexity [39]. Therefore, the understanding of bubble mass transfer in non-Newtonian fluid is not as clear as that in Newtonian fluids.

Researchers proposed various theoretical models and empirical correlations to predict the mass transfer coefficient. The most recognized models for gas-liquid mass transfer are the two‐film model proposed by Whitman [40], the penetration model by Higbie [41] and the surface renewal model by Danckwerts [42]. The two-film theory assumes a stable liquid-gas interface with retardation films on both gas and liquid sides, with solute diffuses through the two films by molecular diffusion (Figure 4A). The two-film theory revealed the mechanism of gas-liquid mass transfer and laid a solid foundation for the subsequent studies. But it is only applicable when the Schmidt number (Sc = ν/D, ν = kinematic viscosity, D = diffusion coefficient) is small. Higbie [41] proposed the penetration theory, which assumes that the fluid is composed of elements, and the gas-liquid mass transfer is completed by the fluid elements. However, the penetration theory regards the residence time of the fluid elements at the gas-liquid interface as a constant, which is inconsistent with reality and difficult to obtain. As an improvement, the surface renewal model assumes that solute diffusion into the liquid phase is unsteady. Also, the residence time of fluid elements at the gas-liquid interface is considered varied, while the probability of elements being renewed is the same.

However, the actual mass transfer process is more complicated, and the assumptions made by the three models are inconsistent with reality to some extent. Additionally, some parameters in the models, like the surface renewal rate and the residence time of fluid elements, are difficult to measure, making the models challenging to be applied to practice. Despite the limitations, the three classical models provide a clear view of the mass transfer process and lay the foundation for studying the gas-liquid mass transfer mechanism.

Based on the three classical models, researchers have developed numerous models for mass transfer. For example, Ma and Yu [43] proposed the three-film theory of gas-liquid mass transfer based on the two-film theory, by considering three resistance films for gas-liquid mass transfer, namely, gas film, liquid film and interfacial resistance film (Figure 4B). The interfacial tension effect in the region near the gas-liquid interface was also included in the three-film theory, which caused a different molecules transportation mode compared to that in other regions. According to the three-film theory, the solute concentration close to the interfacial resistance film is C i = C ∧ e − Δ E R T (1) where ΔE = difference of interfacial energies between the two sides of the interfacial resistance film, T = temperature, and C ^∧ = equilibrium concentration. Perlmutter [44] improved the surface renewal model by considering the flow of fluid elements from the liquid bulk to the interface as a tandem process. In extreme cases, Perlmutter’s model can be converted into the classical models. That is, when there is only one fluid element for renewal, the model becomes the same as the classical surface renewal model, while it turns into the classical penetration model when the number of fluid elements in tandem approaches infinity. Shen et al. [21] proposed an improved surface renewal model by including the instability factors of surface film and the diffusion process in the film. In addition to the above models extended from one of the classical models, models developed by combining two or three of the classical models have also been reported. For example, Hanratty [25] proposed the film-penetration theory as a combination of the two-film theory and the penetration theory.

In addition to the three classical mass transfer models and those extended from them (e.g., Eq. 1), there are models established based on different theories like eddy diffusion theory and whirlpool theory. Eddy diffusion theoretical and whirlpool theoretical models consider the effects of the turbulence structure on mass transfer.

According to the similarity between mass transfer and momentum transfer, Levich [45] proposed an eddy diffusion model by assuming that the interphase mass transfer in semi-infinite space was mainly dominated by molecular diffusion and turbulent diffusion. Based on Levich’s model, King [46] reported that eddy flow dominated mass transfer enhancement in turbulence, where the large-scale eddies generated surface renewal, whereas the effect of those small ones can be damped by surface tension. He obtained the solute concentration distribution in the eddy by combing the eddy velocity equation and the convection-diffusion equation, and then calculated the eddy diffusion coefficient as K L = a 1 n D 1 − 1 n f a 2 n t D 2 n − 1 , n (2) where a and n are constants independent of t.

In the whirlpool models, the high-speed turbulent eddies near the interface are considered to play a leading role in mass transfer, and the mass transfer rate is greatly affected by the degree of turbulence. More specifically, the whirlpool models can be divided into large-scale eddies models (e.g., Fortescue and Pearson [47]), small-scale eddies models (e.g., Lamont and Scott [48]) and individual eddy models (e.g., Luk and Lee [49]). The large-scale eddies models consider that, in turbulent flows, among the energetic eddies with different scales near the interface, it is the large-scale energetic eddies controlling the gas-liquid mass transfer. The formula for the mass transfer coefficient is: K L = 1.46 D U Λ (3) where Λ = integral length from the interface and U = turbulent velocity near the interface. However, it is not easy to measure the eddy size distributions at the interface directly, which limits the practical application of the large eddy model.

The small eddy models consider that, although small-scale eddies contain lower energy, they can be fully mixed with adjacent large-scale energetic eddies, thereby transferring matter and energy among eddies of different scales, through motions including rotation, jetting, etc. Therefore, for the small eddy models, in a fully-developed turbulence field, the minimal viscous dissipative eddies control the mass transfer process. Lamont and Scott [48] developed the equation for the mass transfer coefficient by describing eddy velocities based on Kolmogorov’s eddy theory: K L = A 2 D 1 2 ε ν − 1 2 (4) where A₂ = model constant determined from experimental data, and ε = turbulent kinetic energy dissipation rate.

Both the large and the small eddy models use the statistical averaging method to determine the eddy size and velocity, which poses difficulties in explaining the mass transfer mechanism of an individual eddy. In addition, whether large eddies or small eddies control the mass transfer process in the turbulence field is still debatable. Considering these concerns, Luk and Lee [49] proposed the individual eddy model, which was established based on a local equilibrium hypothesis. That is, although the mass transfer process across the interface was unsteady, the mass transfer within an individual eddy could be considered as stable. By assuming the flow velocity in an individual eddy as u, the component transport equation of the process is solved. Thus, the gas-liquid mass transfer coefficient of the model is obtained (Luk and Lee [49]): K L = 0.9 D V L (5) where V is the velocity fluctuation amplitude, and L is the velocity and length scale of an eddy. Compared with the large-scale eddy theory (e.g., Eq. 2 and 3) and the small-scale eddy theory (e.g., Eq. 4 and 5), the individual eddy theory further deepens the mass transfer mechanism to the level of single eddies.

In addition to the models above, there have been extensive studies on K _L , which has been reported to be a function of various factors, including gas diffusivity, liquid density and viscosity, gas-liquid affinity, as well as aeration and hydraulic conditions [50, 51]. Equations for calculating the mass transfer coefficients of a single bubble under different conditions have been widely reported, some as listed in Table 1. For example, Crift et al. [56] proposed an equation for the mass transfer coefficient of a spherical bubble as follows K L = 2 π 1 − 2.89 max 2.89 , Re 1 2 Re 1 2 S c 1 2 D D m (6) where Re = Reynolds number (Re = ρvL/μ), and D _m = bubble diameter. Kendoush [59] developed the equation for the Sherwood number of ellipsoidal bubbles by improving Crift’s equation (Eq. 6), with Z being the bubble radius function that changes with the bubble quadrant angle: S h = 0.564 Re S c ∫ 0 π 1 + 2 Z 3 θ ⁡ sin 3 ⁡ θ Z 3 θ d θ (7)

In Eq. (7), θ = quadrant of the bubble.

Generally, the mass transfer models for a single bubble are unlikely to be extrapolated to bubble swarms directly. Although the mass transfer of bubble swarms is highly complex, several researchers have made efforts to establish models. Alves et al. [60] established a model of the average mass transfer coefficient for bubble swarms, by using averaged bubble size, gas holdup, specific interfacial area and bubble residence time in bubble swarms, as Eq. (8): K L = 1.13 t m v s D D m + t R − t m c v s D m D 2 3 v l − 1 6 t R (8) where t _R = bubble residence time and t _m = the time span where bubbles behave like mobile fluid particles. Bork et al. [61] proposed an equation for the Sherwood number and verified the accuracy of the model experimentally, as Eq. (9) S h = 4 π F D Re S c S r (9) where F_D = enhancement factor (F_D = 2 for bubble swarms), Sr = Strouhal number ( S r = f d h v s where f is the bubble rise path and d _h is the bubble diameter of the major axis). Additional representative models are listed in Table 2 for reference.

In non-Newtonian fluids, the mass transfer of bubbles is more challenging to compute than Newtonian fluids. Baird and Hamielec [65] developed an equation for the Sherwood number of a single-bubble mass transfer in Newtonian fluids, as Eq. (10) S h = 2 π P e ∫ 0 π v θ v max sin 2 ⁡ θ d θ (10) where Pe = Peclet number defined as P e = V L D , v _θ = surface velocity, and v _max = velocity relative to continuous phase. Based on their equation, Hirose and Moo‐Young [66] proposed a semi-empirical mass transfer formula for a single spherical bubble in non-Newtonian fluids, as Eq. (11) S h = 0.65 P e 1 − 4 m m − 1 2 m + 1 (11)

Here, m = correction factor of non-Newtonian fluids. Kawase and Moo-Young [62] derived equations for the mass transfer coefficients of bubbles moving freely under gravity in non-Newtonian fluids based on the early work of Calderbank and Moo-Young, as Eq. (12) K L = 0.975 D g D m 1 4 (12)

Note that the bubble mass transfer in colored or non-transparent non-Newtonian fluids can be even more challenging to investigate, because the usual non-intrusive measurement methods can hardly be able to visualize the volume change or concentration distribution in these non-transparent fluids [11]. Moreover, a few types of non-Newtonian fluids, e.g., polyacrylamide solution, can exhibit both non-Newtonian behavior and surfactant-like effect, which can decrease K _L and cause additional complexity to the system [67]. Due to the difficulties, most existing models for mass transfer coefficient are developed by fitting the experimental data, which makes them only suitable for some specific working conditions and lack generality.

One of the common intensifications of mass transfer in bubble columns is to construct horizontal or vertical baffles in the reaction chamber, which changes flow direction and enhances turbulence, thereby improving mass transfer. Yin et al. [68] constructed rectangular baffles with staggered arrangement in the channel to improve the gas-liquid mass transfer efficiency of the CO₂-water system. With baffles, the highest bubble mass transfer coefficient reached 2.8 times the coefficient without baffles. The mass transfer is enhanced by the strengthened turbulence in the liquid phase due to the baffle disturbance, which causes bubble breakup and promotes surface renewal. Gu et al. [69] proposed a new type of rotating packed bed to enhance turbulence in the gas phase. With such a rotating packed bed, the mass transfer coefficient was twice that of traditional designs. Another solution is to install stirrers or vibration exciters inside the reaction chamber of a bubble column to enhance circulation and, thereby mass transfer. However, this method is only effective in a narrow range of process parameters [70]. It can also cause undesirable hydraulic conditions in the bubble column, like back mixing and phase separation, which can reduce the mass transfer coefficient [70].

Introducing energy into mass transfer equipment is another essential means to enhance mass transfer, including stirring, oscillation, and adding energy fields (magnetic fields, sound fields, supergravity fields, etc.). For example, introducing pulsations inside a bubble column or subjecting the entire column to pulsations can cause bubble breakup and elimination of bubble rise, which increases the liquid-bubble interfacial area and the contact time, and thereby enhancing mass transfer [70]. Zou et al. [71] reported that stirring and aeration could weaken the phenomenon of concentration polarization and improve mass transfer efficiency significantly. Zhang et al. [72] found that solid particle oscillation enhanced the radial mixing of the fluid, and the mass transfer coefficient was increased by more than 50%. Reichert et al. [73] found that the use of external alternating magnetic fields along with suspended magnetic particles could lead to forced and highly intensive particle movement in liquid-gas mixture, which increased the mass transfer coefficient by up to 200%. However, the methods of stirring and oscillation are not suitable for traditional equipment such as bubble columns and packed columns. There are also limitations to introducing the energy field, e.g., electric and magnetic fields are only suitable for substances with a certain charge and magnetism. Also, the high cost of adding an energy field needs to be taken into consideration for industrial applications [74].

Introducing mass separation agents, dispersed particles as the commonest, can significantly reduce the mass transfer resistance, and accelerate the mass transfer process [75]. Ferreira et al. [76, 77] experimentally studied the effects of adding expandable polystyrene (EPS) particles (hydrophobic) and hollow glass beads (hydrophilic) on the mass transfer in a bubble column. The results showed that the hydrophobic micron-sized EPS particles always harmed the mass transfer process. The effect of hydrophilic hollow glass beads was a function of the solid holdup, i.e., the mass transfer coefficient increased with an increasing solid holdup of up to 10% and decreased with further higher holdups. The fine particles promote surface renewal thus the mass transfer at lower solid holdups, when the change in liquid viscosity is negligible. However, when the solid holdup is over a certain level, the viscosity of the liquid phase increases, and the fine particles at the gas-liquid interface prevent the gas diffusion to the liquid phase. Thus, the mass transfer coefficient decreases.

With the development of nanotechnology in the last decades, nanoparticles have become a popular choice as a mass transfer promoter. Colloids composed of ultrafine nanoparticles (100 nm or smaller) are called nanofluids [77]. Olle et al. [78] experimentally investigated the effects of aqueous suspensions of 20–25 nm magnetic (Fe₃O₄) nanoparticles on the bubble mass transfer in an agitated and sparged reactor. Their results showed that, with nanoparticle volume fractions below 1%, the bubble mass transfer could be enhanced up to 600%. Park et al. [79] measured the chemical absorption rate of CO₂ into an aqueous solution of nanometer-sized colloidal silica (0–31 wt%) and 2-amino-2-methyl-1-propanol in a stirred vessel. They found that the volumetric liquid-side mass transfer coefficient and the absorption rate in the nanofluid decreased as the nanoparticle concentration increased, which could be due to the increase of system viscosity.

There are still problems for using nanoparticles to enhance gas-liquid mass transfer in practice. Specifically, (1) the mass transfer effect of the same type of nanoparticles varies under different experimental conditions, which can be related to the nanoparticle properties, preparation methods and stability of nanofluids [69]; (2) the enhancement mechanism of nanoparticles on mass transfer is still unclear, which is simply attributed to the micro-convection caused by Brownian motion of the nanoparticles currently; (3) few devices are available to observe the movement of nanoparticles under ordinary solid content [80]; (4) due to the limitations of the stability and visualization of nanoparticles, the universality and accuracy of current calculation methods for mass transfer with nanoparticles still require improvement.

Microreactors featured with bubble mass transfer performing in micrometer-sized channels have also been reported as a way to enhance bubble mass transfer [81]. Due to the short diffusion path and huge specific surface area of bubble mass transfer in micrometer-sized channels, the volumetric mass transfer coefficient in microreactors can be one to three orders of magnitude higher than that in traditional gas-liquid reactors [82]. The structure of microchannel is closely related to the mass transfer efficiency. It is usually designed with convergent-divergent channels, bend, built-in obstacle, etc., to increase the gas-liquid interfacial area and enhance the flow disturbance, and thereby improving mass transfer efficiency. External energy was also introduced to disturb the flow, like ultrasound [83]. Many researchers have explored designs of microreactors for various purposes. For example, Ganapathy et al. (2016) [82] studied the mass transfer performance of a microreactor which was with 15 straight parallel channels of 456 μm diameter in a cross-flow inlet configuration. They found that the efficiencies of CO₂ absorption into aqueous diethanolamine nearly reached 100% under certain operating conditions, which indicated the significant effect of mass transfer intensification of the microreactor. Yin et al. (2022) [81] proposed a split-and-recombine (SAR) microreactor, composed of convergent–divergent arc channel and rectangular obstacles. The bubble dynamic features and mass transfer characteristics were revealed, which showed that the SAR microreactor performing well in enhancing the gas–liquid mass transfer with rapid chemical reaction. Liu et al. (2023) reported a bubble-based microreactor (BBMR), and it also performed excellently in promoting the mass transfer process [84]. Nevertheless, there are still many technical barriers for microreactors, like commercialization and specific reaction integration problems. Also, critical analysis including synthesis ability, measurement analysis, extraction, and detection should be conducted before the microreactors being applied widely in industry[85, 86].