Sanjay Gupta

| Independent Director | Former CEO-Dangote Refinery & Petrochemical Project| Former - C&MD Engineers India Limited | Author

 

 

 

 

 

The discussion towards the optimization of the configuration for the Dream Refinery is a sustained affair. As part of the series of podcasts on PP plus and the follow up exchanges on various articles it was felt necessary to examine the various aspects associated with Gasification. In view of the same,  based on the various exchanges, particularly with respect to the bottoms upgradation, and the need to examine Gasification as a possible optimization module in the configuration, spurred a study at our end. After detailed study, we began to see the merit of Gasification, particularly as the gas prices have been a crucial factor in terms of the impact that it has on the economics of the complex. It is well known, that the hydrogen requirements in the modern large complexes is quite high, and unless gas is available competitively, the overall economics of the complex is likely to be adversely impacted.

After preliminary studies, we released an article in which various bottoms Gasification options were examined. In the base configuration, we had envisaged a Coker and linked the Pet coke to an CFBC plant with co-firing of coke + coal. Later, we examined a coke + coal based Gasification unit, to meet the Hydrogen, steam and power requirements of the complex. Also the Gasification facility provided an opportunity to produce a lot of urea as well, to utilize the Co2 produced as a waste stream. This improved the overall carbon footprint of the complex besides incremental revenue generation. The coke based Gasification unit however was large , capital intensive and involved a lot of solids handling. While the merits of Gasification could be foreseen, there were complications too, both in terms of size, plot area, technology and the associated nuances. The inbuilt apprehension of the Refineries indulging in so much of solids handling remained a deterrent.

Upon further study the SDA based pitch Gasification offered an alternative solution. Solids handling was eliminated and the Gasification perse became a lot more simpler. Upon detailed examination it was realized that 65% DAO recovery was ideal both to offer a good solution, wherein the recovery was good for the downstream Hydro-processing, as well as it generated pumpable pitch which would ideally not have hydrocarbon carryover to impact the Gasifier performance. All solids handling was eliminated as a single/dual train SDA of 3.2-3.4 MMTPA seemed to also offer a good solution to cater to all the captive hydrogen, steam, power solution of the complex. In addition,it also enabled urea to be produced from the residual syn gas. All in all, a very economical and revenue generating solution emerged even while all gas import to the complex stood eliminated.

In many ways the SDA slipped in to replace the Coker with a slightly higher capacity. This enabled the SHCU capacity to be rationalized thereby resulting both in Capex and Opex savings. All in all, a win win in terms of offering a better bottoms upgrader.  In context of the same we chose to revisit the configuration to examine the changes that would be prompted out of the replacement of Coker by SDA, followed with a gasification island to meet the captive hydrogen and other requirements of the complex.

 

The base configuration remains unchanged and we shall run through the same for reiteration:

1. The Crude Distillation Unit and Vacuum Distillation Unit for a 20 MMTPA (Million Metric Tonnes Per Annum) refining complex processing a 50:50 blend of Arab Light  and Arab Heavy. For this capacity we proposed a single train CDU for minimizing the Capex of the unit.

2. Feed-stock Blend Properties 50:50 AL/AH considered for the study:

Blend API Gravity: 30.2° to 30.5° API (Medium Sour Crude)

Blend Sulfur Content: 2.3% to 2.4% m/m

3. Atmospheric Charge Furnace: Horizontal tube, multi-pass split furnace designed to heat crude up to 350°C – 370°C before entering the flash zone of the main column.  The furnace will be equipped with an energy conservation scheme for maximizing the system efficiency.

4. Atmospheric fractionation Column - equipped with high-efficiency valve trays or structured packing in the upper rectifying sections to ensure sharp cut points with suitable Pump-Around. Unit to be equipped with Side Strippers for Naphtha, Kerosene, and Diesel to control flash points and initial boiling points.

5.Vacuum Distillation Unit (VDU) - The VDU processes the Atmospheric Residue (AR) from the CDU bottoms under reduced pressure to recover high-value heavy distillates without thermal cracking of the Atmospheric Residue received from CDU bottoms at 340°C–350°C.

6. Vacuum Furnace: Low-residence-time box or vertical cylindrical heater designed to elevate AR temperature to 400°C – 420°C. Steam injection is utilized inside the coil to lower the partial pressure and prevent coking.

7. Vacuum Column may be  maintained at 25 – 40 mm Hg using a vacuum system. Multi-stage steam jet ejectors combined with surface condensers or hybrid systems using liquid ring vacuum pumps can be utilized to create vacuum in the system. Low-pressure-drop structured packing is generally used in the wash zone and fractionation sections, to minimize pressure drop and maximize heavy gas oil recovery.

8. Vacuum column yields Light Vacuum Gas Oil and Heavy Vacuum Gas Oil for upgradation in VGO hydrotreater to generate feed for PFCCU.  The bottoms of the vacuum column is the vacuum residue.

9. Vacuum Residue or Bottoms (>565∘C), is proposed to be routed to Solvent Deasphalting (SDA) unit to generate two distinct cuts. One the DAO at 65% recovery to be routed to VGO Hydrotreater to augment the VGO HDT capacity and correspondingly the PFCCU capacity. The second cut the pitch from SDA to be routed directly to a close coupled Gasifier for generation of syn gas. The SDA capacity is limited to 3.2 MMTPA in a single train configuration.

10. The balance Vacuum Residue is routed to the Slurry Hydrocracking unit. This configuration remains unchanged from the base case, except that the unit becomes a little smaller to facilitate overall optimization. The total VR pool comprises of VR + Decanted oil from PFCCU and the Pyrolysis fuel oil from the Steam Cracker.

11VGO Hydrotreater prepares the combined heavy fractions for the Petrochemical FCC (PFCC), all vacuum gas oils, DAO, and LVGO/HVGO are pooled into the VGO HDT unit. The VGO hydrotreater may operate at a mild severity of 15-20%. The LPG and the naphtha produced from the unit could be routed to the Steam Cracker, while the uncracked hydrotreated bottom could be routed as clean feed to PFCC. The VGO hydrotreater may also receive the VGO from SHCU and enable its proper hydrotreatment for removal of nitrogen etc while maximizing feed to PFCC.

12. To achieve a minimum 20% propylene yield by weight, the Petrochemical Fluid Catalytic Cracking Unit (PFCCU) operates in a high-severity petrochemical mode, using high catalyst-to-oil ratios, elevated riser temperatures above 565∘C, and specialized zeolite-containing catalysts like ZSM-5 to promote secondary cracking of naphtha-range molecules into light Olefins.  The unit is large and perhaps for the given severity will have to be split into two parallel chains of Reactor Regenerator, dedicated MAB/WGCs and a common product factionator and a gas plant. This unit certainly will be of mega scale and will be equipped with an elaborate Tertiary separator system and an over head expander to generate clean green power. The residual heat from the Regenerator off gases will be recovered in dedicated flue gas coolers which would produce significant quantum of HP steam. The off gases from  PFCCU shall be integrated with the steam cracker for recovery of ethylene and cracking of all residual ethane to extinction. Polymer grade propylene shall be recovered in the unit in a dedicated PRU section and all the propane so recovered shall also be recycled to the steam cracker to maximize the Cracker feed.

13. PFCCU naphtha yield due to high PFCCU severity is rich in aromatics (Benzene, Toluene, Xylenes - BTX), making it an ideal feedstock for an integrated aromatic extraction unit rather than direct motor gasoline blending.  After selective hydrotreatment the aromatic rich naphtha s routed to the Aromatic complex, where it combines with CCR effluent, Pyrolysis gasoline from Steam Cracker to maximize the production of Benzene and PX at the expense of toluene and ortho and meta xylene. This adds considerable economic value to the plant. All the PX produced can be directed to a PTA complex while the Benzene could be partially utilized in the LAB facility. Balance of course would be available for merchant sale.

14. The C4 mix from PFCCU along with the I- butane from LPG splitter shall be routed to the Alkylation unit. The alkylate from the unit shall combine with the non aromatic stream from the Aromatic complex to form the gasoline pool.

15.Straight-Run Naphtha split;- SR Naphtha from the CDU is Hydrotreated and fractionated into two distinct cuts to serve different downstream objectives. The capacity of the naphtha Hydrotreater will reduce as compared to the base case as Coker naphtha is eliminated. Light Naphtha from the splitter is routed directly to the Steam Cracker and the Heavy Naphtha is routed partly to the Continuous Catalyst Regeneration (CCR) Reformer to produce high-octane reformate and aromatic extraction feed (BTX) for the aromatic complex. The CCR capacity is so adjusted to ensure that the combined PX from PFCCU naphtha, Pyrolysis gasoline and CCR does not exceed 1.5 MMTPA to support a single train PTA unit. All the balance heavy naphtha will be routed to the Steam Cracker.

16. All the straight run diesel and the cycle oil from PFCCU is upgraded in DHDT. Given the changing market dynamics and the need to swing the production between distillates and petrochemicals , the DHDT unit is operated as a mild Hydro-cracker with 35% severity, such that the LPG and naphtha from the same could be the feed to the Steam Cracker as additional feed.The idea is to maximize the cracking capacity to maximize the production of Petrochemical intermediates. The total diesel product will comprise of DHDT diesel combined at the outlet of DHDT with the diesel from SHCU and from VGO HDT.

17. All LPG i.e. straight run, from VGO Hydrotreater, from SHCU and DHDT is led to a central LPG pool. The LPG could be split in C3 and C4 cut. The C4 cut could be further separated in n- butane and i - butane. The i-butane could be routed to the Alkylation unit and the n-butane could be routed to SDA, both to serve as initial fill as well as continuous make up to . All C3 from LPG and the surplus n - butane shall be routed to Steam Cracker. The idea is to address the maximize the cracker feed and address zero commercial LPG production.

18. The propane from PFCCU also to be routed to cracker. The C4 mix from PFCCU to be sent to the Alkylation unit where in conjunction with i-butane Alkylate will be produced . This will be mixed with the non aromatic cut from the aromatic complex to form the gasoline pool.

19.  As is evident from the above, the idea is to maximize  Steam Cracker feed. Combining all designated naphtha streams, PFCC-derived propane, and the surplus straight-run kerosene yields after capping the ATF production to about 1 mmtpa the total feed allocated to the Steam Cracker complex shall comprise of Hydrotreated Light and heavy naphtha from the NHT, Hydrotreated naphtha from VGO HDT, Mild Hydrocracker and SHCU, hydrotreated Kerosene from CDU, Propane from PFCCU and LPG splitter, surplus propane and off gases from PFCCU.  By shifting the entire naphtha pool and surplus kerosene away from conventional gasoline/diesel/reforming pools and directly into the Steam Cracker, the complex transitions heavily into a refinery-petrochemical integration model. Processing this mixed naphtha-kerosene-propane - butane  feedstock in a modern high-severity steam cracking complex yields high ethylene and propylene. The total feedstock of the Cracker exceeds the minimum economic size of a single cracker. A brief estimate reveals that two Olefin plants of about 1.5 MMTPA capacity each could be considered to handle the total feed and maximize the production of ethylene and propylene. To further augment the Olefins and aromatics the C4 mix from he Steam Crackers is combined into a common K cot unit.

20.Hydrogen requirement: Operating CCR in aromatic mode produces significant volume of high-purity hydrogen, which is vital for offsetting the high hydrogen consumption of the downstream NHT, VGO HDT, DHDT, and Slurry Hydrocracker (SHCU) units. Also the cracker will produce some hydrogen after catering to the internal hydrogenation requirements. The balance hydrogen requirement of the complex shall be catered from the Gasification Island.

21. C4 mix from the steam cracker can either be sent to the Alkylation unit after selective hydrogenation or routed to the K-COT unit to maximize Ethylene, propylene and Aromatics. In this case a K- COT unit is foreseen to maximize the production of petrochemicals. This adds value to the complex. The propylene purge from the Alkylation unit can also be routed to the product propylene pool.

22. Routing the PFCCU off-gases directly into the Steam Cracker’s fractionation and recovery train (upstream of the demethanizer) is a  high-efficiency petrochemical integration. Instead of losing valuable light olefins and paraffins to the refinery fuel gas system, this captures and upgrades them and improves the economics of the plant.

23. All PX will go to PTA for maximizing production. The Stoichiometric Conversion Factor is 1.565

24. Total ethylene produced in the complex will be used to produce world scale LLDPE/HDPE, Butene 1, HDPE, LDPE and impact PP unit multiple trains.

25.The total propylene produced in the complex is used to produce world scale capacity of Homo/Random PP, Impact PP, Acrylates and PO/PG.

26. Given Economic considerations SAP, Cumene/Phenol, MEG and Bis phenol, polycarbonates and VCM/PVC  are not considered though they were considered in the base case configuration. All these plants seemed to be high Capex units with relatively uncertain returns.

27. Amongst the niche Petrochemicals only Pure Esters are preferred due to low execution risk, steady cash flow, and zero solid-handling bottlenecks. Esters plug the vast manufacturing landscape of paints, adhesives, textiles  with minimal sales risk. Additionally, to feed the integrated oxo-alcohol unit producing n-butanol and 2-ethylhexanol from Glacial Acrylic Acid, the synthesis gas (CO+H2) can be sourced from the Partial Oxidation (POX) unit. This serves as good horizontal integration and saves Capex.

28. A single, large-scale Kerosene Hydrotreater (KF) and a massive central Molex n-paraffin extraction unit to pull the required linear paraffins, 500+ KTPA feed from the refinery kerosene pool is foreseen for LAB. A common high-purity Benzene storage and feed distribution header shall supply the benzene for the LAB production. Optimal trains of 165 KTPA are foreseen to achieve 495–500 KTPA. The global Linear Alkylbenzene (LAB) market is  between 3.7 to 4.7 million metric tons annually. Demand is fundamentally oriented to the production of Linear Alkylbenzene Sulfonate (LAS), the world's most widely used biodegradable surfactant for laundry powders, dish-washing liquids, and industrial cleaners. Growth remains steady at a 3.1% to 4.7% CAGR, sustained by global consumer shifts toward high-efficiency washing machines and stricter environmental mandates favoring readily biodegradable surfactant. Rather than selling raw LAB entirely as a merchant commodity, it is proposed to co-locate Linear Alkylbenzene Sulfonic Acid (LABSA) conversion units. This is a high selling product. Diversify beyond household laundry detergents into high-performance industrial sectors, such as Textile & Leather Processing Emulsifiers and wetting agents. Oilfield Chemicals: Enhanced oil recovery (EOR) surfactants. Lubricant Additives: Specialized detergent sulfonates for heavy-duty industrial and automotive oils.  A slipstream of dry, high-purity gaseous SO3 from the sulphuric acid converter can be diverted directly to the LABSA falling-film sulfonation units, leading to cost savings. This facility adds significant revenue to the project.

29. Deploying a  captive sulphuric acid plant running on WSA framework can successfully serve four core demands: alkylation make-up/spent acid regeneration, LABSA gaseous SO3 supply, refinery acid gas treatment, and merchant sales. Alkylation licensors traditionally mandate their proprietary SAR (Sulfuric Acid Regeneration) units because their performance warranties are tied to strict acid strength, 98.5–99 wt% minimum and purity. The sulphuric acid unit can be sourced from others meeting a specified purity and strength standard, ≥98.0 wt% H2SO4 with controlled organics/iron. WSA can meet the objective.

30. Spent alkylation acid contains heavy hydrocarbons. The combustion furnace must feature precise oxygen-trim control and high residence time, 1,000∘C+ operating temperature, to ensure complete cracking of hydrocarbons. Incomplete combustion will result in carbon soot, which poisons the downstream vanadium catalyst.

Gasification island

1. Solvent Dealphalting - The largest single-train capacity for a Solvent Deasphalting (SDA) unit—specifically utilizing supercritical extraction technology—ranges between 70,000 to 80,000 BPSD roughly equivalent to 3.5 to 4.0 million metric tons per annum of vacuum residue feed. KBR (ROSE® Technology) is the undisputed global market leader, holding over 90% of the worldwide installed SDA capacity. Their proprietary Supercritical ROSE, Residuum Oil Supercritical Extraction, process is the standard for high-capacity single-train designs due to its superior energy efficiency and solvent recovery footprint.

2. At 65% lift, the unit will be operating in a deep-extraction zone,  using an n-butane solvent blend, to pull a high yield of metal-reduced, highly paraffinic DAO to feed downstream hydroprocessing unit  while rejecting 35% as a concentrated asphaltene pitch. Preheated vacuum residue is blended inline with a lean supercritical solvent stream in a Solvent-to-Oil volumetric ratio of 5:1 to 7:1.

3. The Asphaltene Extractor is equipped with proprietary high-efficiency internals. The heavy, asphaltene molecules precipitate and descend as a dense phase to the bottom, while the lighter, aliphatic hydrocarbons dissolve into the rising solvent to form the Rich Solvent/DAO Phase exiting the top. The rich overhead stream is heated via heat exchangers and sent to a Supercritical DAO Flash Drum where DAO instantly is separated.

4. Pitch: The bottom slurry from the extractor goes through a high-temperature flash drum and a steam stripper to recover remaining solvent, discharging a hard or semi-solid Pitch product (35%) suitable for Gasification feed. At a high lift of 65% the levels of trace metals (Ni/V), asphaltenes, and CCR in the product are to be carefully controlled to avoid faster catalyst deactivation in the downstream VGO Hydrotreater.

5. Rejecting 35% of a heavy vacuum residue as pitch leaves a very thick, highly viscous bottom product. Temperature control must be carefully controlled to avoid pitch to solidify. The pitch product loop must be designed with high-temperature steam tracing, and progressive gear pumps with shorter leads for quick handling.

POx gasifier

1. Rejected Asphaltene Pitch (35%): 1,120 KTPA forms the feed to the gasifier. This heavy, metal-rich, highly aromatic asphaltene pitch serves as the feed for the Partial Oxidation (POX) gasification block. To process 1,120 KTPA of pitch, a multi-train entrained-flow non-catalytic partial oxidation configuration is foreseen with 2+ 1 configuration of the gasifiers.

2. Typical specific oxygen demand for heavy asphaltic pitch POX is roughly about 0.85 to 0.95 wt% of pure O2per unit weight of pitch. This implies that about 980 KTPA of pure O2 will be required for the Gasification block. The Oxygen required for the PO/PG complex could be additionally included in the ASU plant sizing for integration. Total ASU capacity may be split in two parallel trains to meet the complex requirement.

3. Raw Syngas Production (CO+H2) to the tune of  2.5 to 2.7 standard cubic meters (Nm3) per kg of pitch shall be produced.  This syn gas shall be rich in carbon monoxide 48–52% and hydrogen 42–45%, with minor amounts of CO2, H2S and trace COS.

4.Syngas Coolers High-Pressure Steam Generation- The raw syngas exits the POX reactor core at extreme temperatures (1,300∘C–1,500∘C). To recover this massive thermal energy safely before acid gas removal, specialized Syngas Coolers are deployed on each train. Radiant and convective syngas coolers quench the gas down to roughly 250–300∘C while generating Very high-pressure steam, 125 bar and 540 Deg C. Each active train generates approximately 1.2 to 1.5 tons of high- pressure steam. This massive high-pressure steam generation can be routed directly to utility turbines for power generation or used to drive large turbo-compressors within the complex, significantly lowering the overall utility footprint.

5. After the syn gas is cooled and cleaned up it is sent to the shift section, from where first the hydrogen could be recovered to meet the complex requirements. For the purposes of the study a conservative estimate of 155 KTPA of hydrogen recovery is foreseen from the Gasification block.

6. The remaining syngas is directly connected to the power island comprising of the high-pressure steam boilers and Steam Turbine Generators - STG. The system monetizes the carbon-rich tail gas while capitalizing on the massive 125 bar/540C steam generated directly by the syngas coolers.

7. Combining the syngas cooler steam with the power island boiler generation creates the utility powerhouse. All high-pressure steam headers feed a central Steam Turbine Generator (STG) train operating with condensing/extraction turbines to satisfy the complex's internal electrical demand.  The typical Gasification island balance is as below:

• Pitch Feedstock: 1,120 KTPA from the 3.2 MMTPA unit.
• Oxygen Feed from ASU: 980 KTPA pure O2.
• Total Raw Syngas generation: ∼230 MT/h of raw CO+H2 mix.

Process Complex Steam Demands - for the purposes of the study the total steam requirements for the complex are estimated as below:

• VHP Steam (125 bar): 300 TPH
• MP Steam (15 bar): 600 TPH (extracted/let-down from STGs).
• LP Steam: 500 TPH (low-pressure process heating and deaeration).

Power Generation STG Configuration

• STG 3+1 Configuration (3 operating + 1 standby spare), with each turbine rated at 175 MW gross capacity.
• Running 3 active STGs yields 525 MW gross generation, delivering the required 500 MW net export power to the refinery and petrochemical complex after accounting for internal loads of  ASU and Oxygen compressors.
• Steam Turbine Extraction: The VHP steam expands through the turbines, providing controlled extractions for MP (15 bar) and LP steam headers.

Total VHP Steam Generation (125 bar/540∘C)

To support a power island generating 500 MW net while simultaneously providing steam extractions (600 TPH MP + 500 TPH LP) plus direct process steam (300 TPH VHP) the total  VHP steam would be significant of which close to 300 TPH of VHP steam would be generated by the syn gas coolers from the hot reactor effluent from Gasifiers. The boilers must therefore generate about 1800 TPH of VHP steam by firing the balance syngas and PSA tail gas. 4+1 boilers, generating 500 TPH of steam from each boiler may be installed.

Downstream Syngas Allocation: Acrylates, Urea Potential

After extracting 155 KTPA of pure H2 and routing the required tail gas/syngas balance to the power island boilers, the remaining surplus syngas (CO+H2) and byproduct CO2 can unlock a high-value chemical potential:

A. Acrylates Plant Feed (Oxo-Synthesis / Carbonylation)

Requirement: The complex's 390 KTPA Acrylate Esters & Oxo-Alcohols block requires a dedicated stream of purified Carbon Monoxide (CO) and hydrogen. A slipstream of unshifted or selectively scrubbed syngas is routed to a CO cold-box / cryogenic separation unit to supply the exact molar ratio required for oxo-synthesis.

B. Urea Production Potential

Raw Material Sources from the Gasifier:

CO2 Supply: Massive quantities of pure CO2 are scrubbed out during the Acid Gas Removal and Water-Gas Shift purification steps. Ammonia is produced via a standard Haber-Bosch loop by combining surplus nitrogen (N2 separated from ASU) with a dedicated fraction of hydrogen from the Gasification block. The leftover carbon and hydrogen balance from the pitch can support a world-scale urea facility yielding approximately 800 to 1,000 KTPA of granular urea, helping the carbon waste into a premium solid fertilizer product.

High-Level Economic Analysis

The total investment in the Gasification island is estimated to be around $1.5 to $1.8 billion. The raw feedstock is asphaltene pitch from the SDA unit —a heavy, highly viscous, metal-contaminated refinery byproduct that traditionally commands low economic value. Gasifying this low-value burden converts the low value pitch into prime synthesis gas avoiding the need to source high value gas.

 

• Gasification eliminates reliance on merchant hydrogen or building a standalone Steam Methane Reformer (SMR) dependent on natural gas feedstocks.
• Generates exportable electricity to power the entire refinery complex.
• Utilizes the liberated balance syngas, ASU nitrogen, and AGR-scrubbed CO2 to produce granular urea turning carbon waste into a premium solid fertilizer product with global market demand.
• High-purity sulfuric acid from the WSA plant and acrylates slipstream integration further diversify and maximize asset profitability.
• By utilizing PSA tail gas and balance syngas to fire the auxiliary boilers, the power island operates entirely on internal residues, insulating the complex from external natural gas price volatility.
• Integrating syngas cooler steam with auxiliary boiler generation across extraction-condensing turbines ensures high thermodynamic efficiency, minimizing overall energy loss per barrel of crude processed.

Summary of changes

 

1. Coker replaced with a slightly higher capacity SDA.
2. Slurry Hydrocracker capacity Reduced.
3. POx gasification replaces CFBC system.
4. Gas connectivity to the plant is dropped and so is the SMR facility. All hydrogen sourced from Gasification.
5. NHT capacity reduced due to deletion of Coker.
6. PFCCU capacity is generally the same. However Cracker capacity maximized . all LPG eliminated and routed to Cracker after splitting.
7. DHDT capacity also reduced due to deletion of Coker.
8. Overall Diesel, gasoline reduced and petrochemical make increased.
9. LAB capacity augmented and LABSA included in configuration.
10. MEG, Cumene/phenol dropped from the configuration. Merchant Benzene sale preferred.
11. PO/PG capacity reduced and Acrylates restricted to alcohols and conventional esters.
12. Urea production ensured without gas import.
13. Overall revenue increased and economics improved substantially.
14. Overall Capex of the plant reduced directionally.