Ethylene is one of the most important organic materials; it is used as a building block to produce fibers, plastics, and other chemicals. Global ethylene production capacity was around 1.48 × 10⁸ t in 2014, representing a 32% increase over the past decade [1]. At present, ethylene is almost exclusively produced via the steam cracking of gaseous and liquid hydrocarbon feedstocks such as ethane, naphtha, and gas oil. Due to its high endothermicity and complex product-separation steps, steam cracking is one of the most energy-intensive processes in the chemical industry. Although ethane is the most common feedstock for cracking in the United States and Middle East, more than 80% of the ethylene produced in Europe and the Asian-Pacific region is from naphtha [2,3].

Naphtha is a mixture of hydrocarbons with a boiling point range of 30–200 °C. In naphtha-based steam-cracking processes, naphtha is first fed into the convective section of the furnace for preheating and vaporization. At elevated temperatures in the radiant section of the furnace (750–900 °C or higher), naphtha is cracked into smaller molecules in the absence of catalysts. Gaseous light olefins are formed via reactions involving free radicals [4,5]. Steam cracking is highly endothermic and requires significant external heat input. Coke deposition, which is inevitable in the complex cracking reaction scheme, represents another significant challenge. Diluting the feedstock with steam lowers the propensity for coke formation. However, cracking furnaces still need to be periodically shut down and regenerated with air to avoid the accumulation of coke. The parasitic energy consumption due to steam usage and the need for periodic shutdowns result in an increase in capital and operating costs [6,7]. From an energy consumption standpoint, the pyrolysis section accounts for approximately 2/3 of the total process energy required in naphtha steam crackers, and leads to about 3/4 of the total exergy destruction [3]. Furthermore, large amounts of carbon dioxide (CO₂) are generated in this process due to the need to combust carbonaceous fuels to satisfy the parasitic energy requirements. Therefore, novel technologies that address these shortcomings of steam cracking are highly desirable.

The oxidative cracking of naphtha using a heterogeneous catalyst is a promising option [8–10]. However, the high reactivity of the olefins poses a challenge in identifying a suitable catalyst with high selectivity [11]. Moreover, there is an additional difficulty of separating the oxygenate byproducts [6]. The flammability of the gaseous mixture, which contains oxygen (O₂) and fuels, also presents safety concerns. Hot spots can result from over-oxidation of the hydrocarbons to carbon monoxide (CO) and CO₂, lowering the product selectivity [3,12,13]. The current study investigates a two-step, redox oxy-cracking (ROC) process for naphtha conversion in an O₂-free environment. The first step involves selective oxidation of the hydrogen (H₂) produced from hexane cracking with active lattice oxygen in a mixed-oxide redox catalyst. The redox catalyst particles, which enter the reactor at high temperatures, provide the sensible heat to compensate for the endothermicity of the cracking reactions. Meanwhile, the catalyst is reduced during this step. The reduced redox catalyst is later oxidized, regenerated, and reheated in air to complete the redox loop. The absence of gas-phase oxygen in the naphtha conversion step has the potential to ① inhibit the non-selective reaction pathways, ② reduce the potential safety hazards [13–15], and ③ decrease parasitic energy losses by eliminating the air-separation unit and providing tighter temperature control [16]. The overall process also significantly decreases the energy consumption and, hence, the nitrogen oxide (NO_x) and CO₂ emissions. Despite its advantages over steam cracking and conventional oxy-cracking, this novel ROC approach has yet to be investigated in detail.

We recently reported a chemical looping-oxidative dehydrogenation approach for ethylene production from ethane using operational principles similar to those of ROC [17,18]. The redox catalyst development, performance, reaction mechanism, and overall process efficiency were investigated [17–19]. However, these studies were limited to ethane conversion. The oxidative cracking of n-hexane to olefins using lattice oxygen of VO_x/Ce-Al₂O₃ has been investigated, showing up to 60% olefin selectivity at 30% n-hexane conversion [11]. The performance of sol-gel-synthesized Li/MgO in the oxidative cracking of hexane in the presence of gaseous oxygen, with 28% hexane conversion and 60% selectivity to light olefins, has also been shown [8]. The same authors also studied an integrated plasma-Li/MgO system to improve the oxidative conversion of hexane at the relatively low temperature of 500 °C [9]. Although the integration of plasma was effective to reduce the operating temperature, the light olefin yield was still limited to about 35%. Overall, existing oxy-cracking processes have demonstrated relatively low olefin yields (< 35%). Moreover, studies on the oxy-cracking of naphtha via the chemical looping scheme are limited.

In the current study, a detailed process simulation and analysis is conducted to investigate the performance of the ROC process. Experimental data obtained from a novel redox catalyst composed of alkali-doped mixed oxides were used in the simulation [20]. Single-pass hexane conversions of 79% with 61% light olefin selectivity were achieved in a laboratory-scale reactor. The experimental results were used in the process simulation, and the performance of the ROC process was compared with a state-of-the-art naphtha steam-cracking process. Using ASPEN Plus®, the effects of the reactor and process conditions were investigated (The flowsheets of ASPEN Plus® for hexane cracking and ROC process are given in Figs. S1 and S2 in the supplemental file, respectively). The potential advantages of the ROC process were validated using detailed process analyses. When compared with the steam-cracking process, the net energy demand of the upstream section of the ROC reference case (79% naphtha conversion) was shown to be 67% lower, and to require 25% less compression work. These decreases result in a 52% drop in the overall energy demand and a corresponding reduction in CO₂ emissions, as compared with the conventional cracking process.

Naphtha can vary in composition depending on its source and refinery conditions. It is a complex mixture of hydrocarbons that ranges from straight-chain alkanes to aromatics. One of the major components of naphtha is n-hexane. Reaction (1) illustrates one of the representative cracking reactions of n-hexane to ethylene. In addition to its high endothermicity, the vaporization of naphtha can be energy consuming—the enthalpy of vaporization for n-hexane at 298 K is 359 kJ·kg⁻¹. A simplified process flow diagram for the conventional naphtha steam-cracking process is shown in Fig. 1(a). As the first step, preheated naphtha is thermally cracked in the presence of steam. Since naphtha cracking is equilibrium limited and tends to form coke, steam is used as a diluent to enhance naphtha conversion and inhibit coke formation. In the cracking step, naphtha is passed through the tubes of the cracking furnace. The furnace is operated at high temperatures (∼1000 °C), so the heat for the endothermic cracking reactions is provided by the fuel combustion. The furnace consists of the radiant zone, where the cracking reaction occurs at > 750 °C, and the convective zone, where the feedstock preheating takes place at up to 650 °C. The product stream is at a temperature of approximately 800 °C. Using transfer line heat exchangers, the stream is rapidly quenched to 300–400 °C. This avoids further cracking and preserves the product compositions. As the product contains significant amounts of heavy hydrocarbons, a primary fractionator is installed to remove tar and oily material (boiling point > 200 °C) from the cracked gases. The product is further quenched to near-ambient temperature to remove the “heavies.” This is followed by a multistage compression system to raise the pressure of the lighter gaseous products to around 4 MPa. Drying and acid gas removal (AGR) units are present at the intermediate compressor stages, as they need to be operated at high pressures. An adsorption column of molecular sieves facilitates the drying, followed by a caustic wash to remove acid gases such as CO₂. The formation of hydrates is prevented by having the drying unit before the caustic wash [2].

The light components from the compressed stream are sequentially separated from the heavy fractions using a separation train of distillation columns that ultimately yield high-purity ethylene and propylene. At the beginning of the sequence, methane (CH₄) and lighter components (primarily H₂) are selectively separated using a demethanizer (DM). A deethanizer (DE) column separates acetylene, ethane, and ethylene from the bottom product of the DM. The distillate of the DE is reacted with H₂ to convert acetylene to ethylene, and then further distilled using a C2 splitter to generate the ethylene product. The C3+ components exit at the bottom of the DE, and are further purified in a depropanizer (DP) to separate the C4+ components. The debutanizer (DB) separates the C4 materials, raw pyrolysis gasoline, C5 materials, and aromatics from the bottom of the DP [2]. The C3 splitter follows the DP and recovers propylene, which is a valuable product.

Fig. 1(b) depicts the naphtha ROC process. A reactor-regenerator scheme is used instead of the energy-intensive furnaces. An experimental product distribution for the ROC scheme, using an alkali-doped calcium-manganese (Ca-Mn)-based oxide similar to that reported in our previous work [20], is used in the current work. For the ASPEN Plus® simulations, an oxide mixture containing calcium oxide (CaO) and manganese oxide (MnO_x) is used because the specific mixed-oxide is not available in the ASPEN Plus® database. Mn₃O₄ acts as the oxygen carrier phase participating in the oxidative dehydrogenation (ODH) reaction through lattice oxygen donation. Mn₃O₄ is reduced to MnO following Reaction (2). Naphtha is converted to a variety of products; however, for simplicity, Reaction (2) portrays n-hexane to ethylene as the model reaction, which involves the combustion of H₂ to water. In a regenerator, the reduced redox catalyst (oxygen carrier) is replenished with air via Reaction (3) (which is highly exothermic). The ASPEN Plus® database values indicate that the surrogate catalyst exhibits similar heat of reaction and heat capacity values to those of the actual redox catalysts. In the current work, the reducer is operated at 775 °C and 1 atm (1 atm = 1.013 × 10⁵ Pa), while the regenerator is operated at 935 °C and 1 atm. The redox loop is completed by recirculating the re-oxidized particles to the reducer. The redox catalyst provides both the necessary lattice oxygen and the heat required for ROC reactions. The naphtha and the air are preheated to 650 °C. The oxygen-depleted hot air from the regenerator is used to generate high-pressure steam to compensate for parasitic energy requirements. The gaseous output from the reducer passes through a series of downstream separation units similar to steam cracking. An amine scrubbing unit for CO₂ removal is required for the ROC process, as this process produces more CO₂ than traditional cracking. The heat requirements of the reducer and naphtha preheating are sufficiently provided by the heat stored in the re-oxidized solids. This is possible due to their high heat capacity and to the temperature difference (ΔT) between the reducer and regenerator. The reducer is endothermic, whereas the regenerator is exothermic. The reducer-regenerator system is operated under adiabatic conditions. The cracking furnace constitutes a major share of the total energy requirement; since it is replaced by the ROC reactors, the energy demand is reduced. In the following sections, the ROC process is compared with naphtha cracking based on the overall energy requirements.

The performance of ROC was evaluated and compared with naphtha steam cracking using the ASPEN Plus® simulator [21]. Two kinds of naphtha feeds were used to model the cracking process: ① a complex naphtha consisting of a variety of hydrocarbons, based on literature data; and ② n-hexane. Since n-hexane is a major component of naphtha, the process model can be significantly simplified, provided that the results using the n-hexane feed provide a reasonable representation of those using the complex feed. The composition of the complex naphtha is shown in Fig. 2, which has been optimized to mimic “true” naphtha. For the simplicity of the model, pure n-hexane was used as a representative compound for naphtha in the ROC process. Table 1 summarizes the ASPEN Plus® modules, property methods, and physical property databanks, while Table 2 lists the carbon yields for the various processes. The cracking results are based on literature data [2,22], whereas the ROC results are experimental results that were published in our previous work [20]. A higher hexane conversion case was also considered, in which 97% of the hexane is assumed to be converted in the ROC process. Experiments on blank hexane cracking and the ROC case revealed an increase of roughly 30% in ethylene and propylene yield for the ROC. Keeping the CO_x yield similar to that in the base ROC case, the ethylene and propylene yields were proportionally increased for the high ROC case, with the overall hexane conversion being 97%. The other products were assumed to have yields similar to that of the base ROC. The product distributions of this hypothetical high ROC case are given in Table S1 in the supplemental file. The supplemental file also contains general information with respect to process configuration in ASPEN Plus®, materials flow, etc.

The drying unit and the primary fractionator were modeled as flash columns, in which almost all of the water is condensed along with the heavy hydrocarbons. To simulate the AGR unit, a separator (Sep) was used, with its energy consumption accounted for. This was placed at the exit of the compressor. The solids used in the simulation include Mn₃O₄ and MnO, with CaO added as an inert substance in order to closely represent the actual redox catalyst [20]. Table 3 lists the key operating conditions and simulation assumptions.

In both the complex and n-hexane feed cases, the hydrocarbon feedstock is introduced to the cracking furnace after being preheated to 650 °C along with steam. A steam dilution of 0.5 kg per kilogram fresh feed is used [2]. The same preheating temperature is used in the ROC case, without steam dilution. In the process simulation, the Mn₃O₄-containing redox catalyst particles are initially fed into the reducer. They are reduced and flow to the regenerator, which is maintained at 935 °C, with the temperature of the reducer at 160 °C lower. In the regenerator, 10% excess air preheated to 650 °C is used to completely oxidize the reduced oxygen carrier (Reaction (3)) prior to recirculation to the reducer for another redox loop. Based on the product distribution in Table 2, the ROC involves the combustion of 75% of the H₂ to H₂O.

To prevent further cracking, the product gases in all the cases are quenched to a temperature of 100 °C. Heat released from the quenching is assumed to be utilized to generate steam and preheat the feed streams. The exothermicity of the reactor-regenerator system can also produce low-pressure steam, which can reduce the upstream energy load. The cooled gases are compressed to 2.5 MPa and flashed at 100 °C, in order to remove the heavy hydrocarbons and water from the stream. In the present model, this compressor, with the flash column, functions as the primary fractionator. The light stream is further compressed to a pressure of 42 atm [2] using a four-stage isentropic compressor. The primary component of acid gas is assumed to be CO₂, with no hydrogen sulfide (H₂S) present. Negligible energy consumption is assumed for the caustic wash. For the ROC, which has substantially more CO₂ in the product stream, an amine scrubbing process is assumed, with an energy consumption of 0.11 MW·h per tonne CO₂ [23]. Table 4 lists the operating conditions of the downstream columns for the cracking process [24]. The compressed stream is cooled to −100 °C prior to methane removal in the DM. This refrigeration step reduces the load on the DM by removing most of the H₂ and CO. A single pass is assumed in all the cases. The reflux ratio for each column is fixed as twice that of the minimum reflux ratio. For a fair comparison, the same number of stages and recovery were assumed for the C2 splitters in both cases. A deacetylenizer is used after the DE, in order to hydrogenate the acetylene in the DE distillate (see Table S1 in the supplemental file). Purified H₂ for this reaction is obtained through pressure swing adsorption (PSA), which operates with 80% H₂ recovery, an exit purity of 99% H₂, and a pressure drop of 0.1 MPa [25].

For process comparison, the different forms of energy (thermal, steam, and electricity) are converted into the same thermal basis. The higher heating value efficiencies used for different energy conversions are given in Table 3. The refrigeration power is calculated using the power-temperature graph provided in Ref. [26], and using the condenser and reboiler duties of the columns, as calculated by ASPEN Plus®. Table 4 [24] lists the refrigerants chosen for each unit.

The current study investigates an ROC process for light olefin production from naphtha. Based on our recently obtained experimental data using an Mn-based redox catalyst, the ROC process was simulated using ASPEN Plus® and compared with the traditional steam-cracking process. Compared with cracking, ROC replaces the energy-intensive cracking furnaces with an autothermal reducer-regenerator system. This results in significant energy savings. Moreover, the experimental results indicated significantly improved ethylene and propylene yields. For each tonne of ethylene product, the ROC process has the potential to reduce the overall thermal energy demand by 52% when compared with steam cracking. While the conventional naphtha cracking process is estimated to consume 32.5 GJ of energy in order to produce one tonne of ethylene, the ROC only requires a net thermal energy input of 15.5 GJ. The ROC can also provide a reduction of over 50% in CO₂ emissions. Specifically, the upstream section of the ROC process consumes approximately 67% less energy while producing 28% more ethylene and propylene for every kilogram of naphtha feedstock. The absence of the cracking furnaces also results in a considerable reduction in NO_x emissions for the ROC process. Overall, the current study indicates that the novel ROC concept has the potential to be an attractive and environmentally friendly option for ethylene and propylene production from naphtha.